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Pressurized fluidized bed combustion and gasification

power systems

VALTION TEKNILLINEN TUTKIMUSKESKUS STATENS TEKNISKA FORSKNINGSCENTRAL

TECHNICAL RESEARCH CENTRE OF FINLAND ESPOO 1988

8ISIR1BUTI0N OF THIS DOCUMENT IS UNLIMITED FOREIGN SALES PROHIBITED

DISCLAIMER

This report was prepared as an account of work sponsored by an agency of the United States Government. Neither the United States Government nor any agency thereof, nor any of their employees, makes any warranty, express or implied, or assumes any legal liability or responsibility for the accuracy, completeness, or usefulness of any information, apparatus, product, or process disclosed, or represents that its use would not infringe privately owned rights. Reference herein to any specific commercial product, process, or service by trade name, trademark, manufacturer, or otherwise does not necessarily constitute or imply its endorsement, recommendation, or favoring by the United States Government or any agency thereof. The views and opinions of authors expressed herein do not necessarily state or reflect those of the United States Government or any agency thereof.

DISCLAIM ER

Portions of this document may be illegible in electronic image

products. Images are produced from the best available

original document.

UDK 662.74:662.76:66.096.5

Keywords:combustion, fluidized bed processors, pressurizing, gasification, coal, peat, electric power plants

VTT SYMPOSIUM 83

VTT-SYMP—83

DE89 766970

rr ]

Pressurized fluidized bed combustion and gasification power systems

Espoo, 23.-24.11.1987

Editor: Maija KorhonenTJ£.

Organizers:Helsinki University of Technology,Laboratory of Energy Economics and Power Plants Technical Research Centre of Finland,Laboratory of Fuel Processing and Lubrication Technology Imatran Voima Oy

MASTERVALTION TEKNILLINEN TUTKIMUSKESKUS STATENS TEKNISKA FORSKNINGSCENTRAL TECHNICAL RESEARCH CENTRE OF FINLAND Espoo 1988

BtSTRiBUTION GF THIS DOCUMENT IS UNLIMITEDFOREIGN SALES PROHIBITED

ISSN 0357-9387Copyright © Valtion teknillinen tutkimuskeskus (VTT) 1988

JULKAISIJA - UTGIVARE - PUBLISHER

Valtion teknillinen tutkimuskeskus (VTT), Vuorimiehentie 5, 02150 Espoopuh. vaihde (90) 4561, teleksi 122972 vttha sf r j

L . . _v4.

Statens tekniska forskningscentral (VTT), Bergsmansvagen 5, 02150 Esbo tel. vaxel (90) 4561, telex 122972 vttha sf

Technical Research Centre of Finland (VTT), Vuorimiehentie 5, SF—02150 Espoo, Finland phone intemat. + 358 0 4561, telex 122972 vttha sf

r r

VTT, Polttoainejalostus-ja voitelutekniikan laboratorio, Biologinkuja 3-5,02150 Espoo puh. vaihde (90) 4561

VTT, Laboratoriet for bransleforadlings- och smoijteknik, Biologgranden 3-5, 02150 Esbo tel. vaxel (90) 4561

VTT, Laboratory of Fuel Processing and Lubrication Technology, Biologinkuja 3—5, SF—02150 Espoo, Finland

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VTT OFFSET?AINO, ESPOO 1988

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FOREWORD

Antero Jahkola, Prof.Helslnlcf University of Technology Espoo, Finland

The application of fluidized bed combustion to energy production Is

cornnerclally established and Its Importance Internationally recognized.

So far It has based on atmospheric combustion technology (AFBC).

Pressurized fluidized bed combustion (PFBC) has not yet been applied to

commercial energy production 1n any country. However, test plants have

been constructed, extensive research projects carried out and the pilot

plant stage has already been reached.

In Finland, research Into pressurized fluidized bed combustion was

initiated In 1985, at first by studying thoroughly research results

obtained in other countries and by evaluating the advantages of

fluidized bed combustion and gasification technology within the context

of Finland's energy economy. The preliminary studies warranted the ini­

tiation of experimental research and the construction of test equipment

for pressurized fluidized bed combustion and gasification. A research

project of several years' duration was Initiated In 1986 as a joint

project between the Helsinki University of Technology, the Technical

Research Centre of Finland, and Imatran Voima Oy. The first pressurized

combustion experiments with peat were carried out In summer 1987. The

work 1s funded mainly by the Finnish Ministry of Trade and Industry.

It is also of crucial Importance to the progress of the research work,

to Increase the level of expertise of our researchers and hence inter­

national cooperation and Interchange of views. International seminars

and conferences are especially well suited for this purpose. The PFGC/G

Seminar organized in Otanleml on November 23 - 24, 1987, was the first

attempt towards this goal. The Seminar was planned and organized by the

executive group of the BFBC/G research project.

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The foreign lecturers Invited to the Seminar represent the highest

international expertise. They reviewed the present results and level

of international research and development in the field of pressurized .

fluidized bed combustion and gasification, and problems that require

further research. The aim of the Finnish lecturers was to inform about

the research work, its aims and results.

The first PFBC/G seminar aroused gratifyingly, extensive interest

in Finland and also in the other Nordic countries. The seminar was

attended by 75 participants, mainly specialists and researchers in

the field.

I wish to extend iqy best thanks to the participants, organizers, and

lecturers, in particular to the foreign guests, for the successful

seminar.

Espoo, January 1988

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Antero Jahkola

Chairman of the Organizing Committee

CONTENTS

FOREWORD 3

Antero Jahkola

OPENING ADDRESS 7

Seppo Hannus

RESEARCH INTO PFBC IN FINLAND

POWER AND HEAT GENERATING PFBC COMBI PROCESSES AND 9

THEIR POTENTIAL MARKETS IN FINLAND

Antero Jahkola, Seppo Hulkkonen S Karl Savlharju

GASIFICATION COMBINED CYCLE POWER GENERATION - 37

PROCESS ALTERNATIVES

Esa Kurkela

THE OTANIEMI PFBC/G TEST RIG AND RESEARCH PROJECT 53

Seppo Hulkkonen, Antero Jahkola & Esa Kurkela

PRESSURIZED COMBUSTION TECHNOLOGY

PFBC RESEARCH AND DEVELOPMENT - BRITISH EXPERIENCES

H. Raymond Hoy 59

PRESENT STATUS AND FURTHER DEVELOPMENT OF PFBC 99

COMBINED-CYCLE PLANT

Krishna K. Pillai

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PRESSURIZED FLUIDIZED BED COMBUSTION OF PEAT

Andras Horvath

121

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GASIFICATION COMBINED CYCLE POWER GENERATION

NEW COAL TECHNOLOGIES 133

Rainer Muller

THE RHEINBRAUN HTW-GASIFICATION TECHNOLOGY FOR 161

COMBINED CYCLE POWER GENERATION

Eberhard Nitschke

THE PRESSURIZED BED PEAT GASIFICATION PLANT OF 193

KEMIRA OY IN OULU

Kalevi Kuorelahti

TECHNO-ECONOMIC EVALUATION OF DIFFERENT GASIFICATION 199

COMBINED CYCLE PROCESSES

Yrjo Solantausta

l _E _

HOT GAS CLEANUP

HOT GAS CLEANUP: THE STATE-OF-THE-ART, AND PROSPECTS 213

FOR THE FUTURE

Roland Clift

DIDIER HOT GAS FILTER SYSTEMS 237

Ulrich Weidner

PARTICLE MEASUREMENTS AT THE bTAHlE^I PFBC/G TEST RIG 251

Seppo Hulkkonen

THERMODYNAMIC STUDY OF ALKALI METALS RELEASE IN PRESSURISED 263

FLUIDISED-BED COMBUSTION AND GASIFICATION OF PEAT

Wahab Mojtahedi & Rainer Backman

APPENDIX

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OPENING ADDRESS

Seppo Hannus, Dr.

Ministry of Trade and Industry

Energy Department

00131 Helsinki, Finland

On behalf of the Ministry of Trade and Industry and its Energy Depart­

ment I have the prlvlledge to open this seminar on pressurized

fluidized bed combustion and gasification power systems.

The subject of this seminar has a high priority both Internationally

and here in Finland. There are several reasons for this and they origi­

nate from the present situation in the world energy scene. With the

continuying growth of energy consumption and. In particular, electrici­

ty consumption more energy production capacity is needed. This Is true

both in the near and long term future. At the same time the available

energy sources are few. The present trends indicate that a large part

of the additional energy needs will or have to be satisfied with solid

fuels, especially coal. And these prospects are not problem-free. En­

vironmental concern already plays a major role in energy policy, and in

a situation with growing coal consumption the concern is likely to

become more acute.

Looking beyond the year 2000 we can see the present electricity capaci­

ty becoming old and needing replacement. In Finland like In many other

industrialized countries a large part of the electricity capacity has

been constructed In the 1970s and will have to be replaced in the first

or second decade of the next century. Much more decisions on the con­

struction of new capacity will have to be made In the 1990s than has

been necessary In the 1980s. In this situation the first priority 1s to

maximize the electricity capacity in combined heat and power genera­

tion. Since solid fuels will, most probably, play a significant role In

future power systems extensions we have to be prepared to meet high

environmental emission standards, too.

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These prospects and concerns are the main motivation why the Ministry

of Trade and Industry has set a high priority on combustion research

and development work. The information available to the Ministry indi­

cates that there is a considerable long-term technological potential

for developing better energy production and combustion processes. It

has also been estimated that the Finnish technological and scientific

expertise in this field has the possibility for becoming competitive on

the international level. And thus we could be able to develop our own

technologies instead of being totally dependent on technology transfer

from other countries. This, however, implies active cooperation with

other countries carrying out similar research activities. The future

will show whether we have been realistic in estimating our resources

and capacities.

In practical terms, the Ministry has decided to launch an extensive

long-term research programme on advanced combustion technologies. It

has been estimated that the total cost of this five-year programme

would be in the order of FIM 100 million, of which about one half would

come from the public energy research funds and the other half would be

draw from other sources. This project is called Liekki (which is Fin­

nish and means flame).

Like all the participants of the Seminar well know the most promising

new combustion technologies are related to pressurized fluidized bed

combustion and gasification power systems which are the subject of this

seminar. I will not now go into the technical substance of these topics

but on behalf of the Ministry of Trade and Industry wish the best suc­

cess to this Seminar.

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POWER AND HEAT GENERATING PFBC COMBI PROCESSES AND THEIR POTENTIAL MARKET IN FINLAND

A. Jahkola, Prof., S. Hulkkonen, M.Sc.(Eng.) K. Savlharju, M.Sc.(Eng.)*Helsinki University of Technology and Technical Research Centre of Finland* Otaniemi, Espoo, Finland

Abstract

Pressurized fluidized bed combustion/gasification makes it possible to construct gas turbine power plants based on solid fuels. By connecting, after the gas turbine, a steam-generator and a back-pressure turbine we/obtain a combined process, which has a good thermal efficiency and a two-threefold power to heat ratio compared with conventional back-pressure power plants utilizing solid fuels. The ratio, however, decisively depends on the allowable inlet temperature of the gas turbine. Low calorific value fuels are also suitable for such a power plant. The PFBC combined processes are of special interest from the Finnish point of view, as the growth of heat load is saturating in industry and in cities while the demand of electric power is growing further rapidly.

In the report alternative combined PFBC-cogeneration power processes are compared. The best power to heat ratio can be obtained by means of a partial gasification process equipped with an after-burning combustor. In case of a cogeneration power process the power to heat ratio is 1.0-1.3 whilst it is only 0.5 for conventional cogeneration power plants.

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According to the results of a market study summarized in the report nearly 3000 megawatts of PFBC combined cogeneration power plant capacity can be built in industrial or municipal power plants, of which 1400 megawatts are the increase in the capacity of conventional back-pressure power plants.

According to preliminary profitability calculations the economic competitiveness of the PFBC-cogeneration processes also seems to be fairly good compared to that of the conventional power plants.

1 INTRODUCTION

Pressurized fluidized bed combustion has some advantages compared with the conventional combustion methods and also with atmospheric fluidized-bed combustion. In a fluidized bed boiler it is possible to use a great variety of fuels, including low-grade fuels, efficiently and with low emissions. Due to high pressure and efficient heat transfer the pressurized fluidized bed boiler is small-sized.

Pressurized combustion makes it possible to construct gas turbine plants based on solid fuels. The gas turbine process preceding a conventional steam power plant process improves the efficiency of a condensing power plant and the power to heat ratio of a cogeneration power plant.

2 CHARACTERISTIC PROPERTIES OF PFBC COMBINED CYCLE PROCESSES

Figure 1 shows a process scheme of a small-scale PFBC district heating power plant based on steam-cooled combined cycle system.

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Air

Stack

Figure 1. PFBC-combined cycle cogeneration power plant process.

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Figure 2. Conventional cogeneration power plant process.

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The combustion chamber of the steam generator is pressurized to an operating pressure of 6 - 12 bar. After cleaning in a hot gas dust filter, the flue gases expand through the gas turbine.The heat remaining in the flue gases is transferred to the feed water of the steam cycle. The generated steam is used to drive a steam turbine and finally to generate district heat.

The fuel (peat) input of the power plant is 50 MW, resulting in 28 MW district heat and 15 MW electric power. Approximately 30% of the electric output is generated in the gas turbine.

The power to heat ratio is a very characteristic feature in case of cogeneration power plants. It is defined as the output of electricity in relation to the output of district heat.

The effect of the gas turbine inlet temperature on the power to heat ratio can be clearly seen in figure 3 where the compressor pressure ratio is 8:1. The power to heat ratio is growing sharply when the gas turbine inlet temperature is increased.

The effect of the compressor pressure ratio on the power to heat ratio is shown in figure 4. At higher pressures the power to heat ratio could be improved by using intercooling in the compressor.

The effect of the moisture content of peat on the power to heat ratio is shown in figure 5, when the pressure ratio of the compressor is 8:1.

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W :S0%

1000 tec) 4

FlgunS. The aftoct of gas turblnt Infat twnptratum Figure 4. The effect of the compressor pressure(t) on the power lo heet retio (a)- ratio (x) on the power to heat ratio (a).

Moisture o( peat w - 50%; Pressure ratio x - 8:1

port

oc.

70 w(%)

Figures. The effect of the moisture contentfw) of Figures. Overall thermal effiency ft) as a function peat on the power to heat ratio (a). of the moisture content of peat (w).Gas turbine Inlet temperature as parameter

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The gas turbine inlet temperature is used as the third parameter. The power to heat ratio increases when the moisture content of the fuel is increased. This is due to a large amount of water vapour in the flue gases which has a positive effect on turbine output. It should, however, be noted that the gas turbine inlet temperature also increases, when the moisture content is decreased.

The effect of the fuel moisture on the overall thermal efficiency is shown in figure 6. The thermal efficiency decreases when the moisture content of the fuel is increased. The reduction in efficiency is due to increased flue gas losses.

3 THE MOST ATTRACTIVE PROCESS ALTERNATIVES OF PFBC COMBINED CYCLE POWER PLANTS

For extensive utilization of combined cycle power plant processes in Finnish conditions one of the basic requirements is the capability of cogeneration of power and heat. It is advantageous if hereby the power to heat ratio is as high as possible. In addition, it should be possible to use indigenous low-grade fuels as the main fuels.

Principally, there is quite a number of process alternatives applicable to PFBC combined cycle power plants. In the following only the most attractive ones will be reviewed.

For the small-scale cogeneration power plants the PFBC process alternative given in figure 1 is rather attractive. It has steam cooling of the fluidized bed, and, therefore, a simple process scheme. It is suitable for low-grade fuels of high moisture content. However, due to the moderate inlet temperature of the gas turbine the power to heat ratio is relatively low as seen in figures 3-5.

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For the larger scale cogeneration power plants the following four process alternatives can be seen as the most. attractive ones. They are suitable for the thermal output of appr. 200 MW.

3.1 Steam-cooled process

Figure 7 shows a PFBC combined cycle cogeneration power plant process where the bed is cooled with a water/steam coil (supercharged boiler cycle).

The bed temperature at 100% load is 850°C, and the flue gas temperature before the stack 150°C. The pressure ratio is 12:1 and the air coefficient 1.25.The steam turbine is provided with reheating. The district heat output of the plant is 105 MW, divided almost equally between two heat exchangers. The electric power output of the steam turbine is 49 MW and of the gas turbine 24 MW, the power to heat ratio being 0.693 for the whole plant. The temperature of the district heating water leaving the plant is 120 °C in the case reviewed. A higher power to heat ratio could be obtained with a lower water temperature.

3.2 Air-cooled process

Figure 8 shows an air-cooled PFBC cogeneration power plant process. In this case some of the by compressor pressurized air is conducted into the heat exchanger to cool the fluidized bed (air heater), and the rest is used for combustion in the fluid bed. After the gas cleaning the air and gas flows are combined and fed into the gas turbine.The fluidized bed combustor is pressurized to 12 bar. Approximately 1/3 of the compressed air goes into combustion and 2/3 are conducted through the cooling coils of the bed.

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Process-values at dimensioning point

Steam cycle Gas cyclem P t h m p t h

kg/s bar •c kJ/kg kg/s bar •c kJ/kg1 46.3 160 565 3475 1 97.5 1 15 163 46.3 42 565 3590 2 97.5 12 351 3714 1.84 4.5 315 3090 4 118 11.5 850 10506 21.4 0.6 135 2758 5 118 1 450 541

16 46.3 160 330 1530 7 118 1 150 17017 358 5019 358 120

0f -200 MW 0dh - 105.3 MWpot - 24.0 MW Pst - 49,0 MW

a - 0.693

w TL. L

Figure 7. Steam-cooled PFBC cogeneration power process.

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Process values at dimensioning point

Steam cycle Gas cyclem P t h m P t hkg's bar *€ kJ/kg kg/s bar °C kJ/kg

1 48.3 160 565 3475 1 97.5 1 15 163 48.3 42 565 3590 2 97.5 12 351 3714 1.8 4.5 315 3090 4 43.7 11.5 850 10506 22.3 0.6 135 2758 5 105.1 11.5 850 1050

16 48.3 160 330 1530 6 105.1 1 460 54117 375 50 7 118 1 473 55619 375 120 9 118 1 150 170

11 61.3 11.5 850 1050

0, = 200 MW (tot) ®dh = 110.2 MWpgt!= 17.2 MW Pst = 51.4 MW

a = 0.628

Figure 8. Air-cooled PFBC cogeneration power process.In exhaust-gas boiler supporting burning.

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The air coefficient for combustion is 1.25. The bed temperature is 850 °C, and the temperature of flue gases before the stack 150 0C.Steam is produced in an exhaust-gas boiler equipped with additional combustion. Approximately 2/3 of the total fuel input of the plant are fired in the exhaust-gas boiler. An atmospheric fluidized bed boiler can be used as the exhaust-gas boiler.

The district heat output of the plant is 110 MW, which is divided almost equally between two heat exchangers. The electric power output generated by the steam turbine is 51 MW, and that of the gas turbine 17 MW. This gives a power to heat ratio of 0.628.

3.3 Steam-cooled process with partial gasification of fuel

Figure 9 shows a steam-cooled combined cycle cogeneration power scheme, where a PFB-reactor is used as a pyrolysing gasifier operating with understoichiometric air. The gasifier is cooled by superheated steam. The gas turbine inlet temperature is 1100 °C. Hot product gases are cleaned by a clean-up system and conducted into the combustors of two parallel gas-turbine units. The temperature of product hot gases is kept relatively low (600 °C). After the gas turbine flue gases are conducted into two exhaust gas boilers which produce steam for a back-pressure steam turbine. The boilers are also used as additional heaters of the district heating water. The gas turbines generate 2 x 40 MW electric power and the steam turbine 34 MW. The total net electric power output is 108 MW. The respective district heat output is 113 MW. The power to heat ratio is thus 0.956.

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boiler boiler

113MW

Hot go& clean-up

Figure 9. Combined cycle steam cooled PFBC/G cogeneration power plant process.

Electric power output = 108 MW;District heat output =113 MW Power to heat ratio a = 0,956

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3.4 Air-cooled process with partial gasification of fuel

A scheme of an air-cooled process provided with partial gasification is given in figure 10. The fuel is fed into the gasifier in which, due to the prevailing temperature, partial gasification takes place. The remaining combustible char is separated in the cyclone, and conducted into the air-cooled PFB-boiler. Some of the flue gases formed in the boiler are conducted into the gasifier, and the remainder is combined with the heated cooling. The gas released from the gasifier is burned together with the air/flue gas mixture in the after-burner before the gas turbine. The temperature of the fluidized bed is 850-900 °C, from where the gas temperature is raised in the after-burner to 1090 °C before the gas turbine. The quantity of pressurized air conducted into the process is considerably larger than the air flow of preceding processes. The total air coefficient used in this case is about 2.4. Steam is produced in the exhaust-gas boiler.The district heat output of the plant is 73 MW, the electric power output generated by the steam turbine is 34 MW and by the gas turbine 57 MW, which is considerably higher than in other alternatives. The power to heat ratio of the plant is thus 1.25.

3.5 Comparison between processes

The clearly highest power to heat ratio among the processes under comparison is obtained with the processes provided with partial gasification. Thus the proportion of electric power output generated by the gas turbine is fairly large due to the high inlet temperature of the gas turbine. Due to the high temperature level the temperature of flue gases after the gas turbine is also sufficient for generating high-standard steam.

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Process values at dimensioning point

Steam cyclem P t hkg/s bar ‘C kJ/kg

1 32.0 160 565 34753 32.0 42 565 35904 1.2 4.5 315 30906 14.8 0.6 135 2758

16 32.0 160 330 153017 248 5019 248 120

0f = 200 MWPgt = 57.2 MW

a =1.25

Gas cyclem P t hkg/s bar *0 kJ/kg

1 183.7 1 15 152 183.7 12 351 3715 204.2 11.5 1090 13136 204.2 1 610 6987 204.2 1 302 3199 204.2 1 150 170

°dh = 73,0 MW

Pst = 33.8 MW

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Figure 10. Air-cooled PFBC cogeneration power process with partial

gasification.

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The highest power to heat ratio can be achieved by the air-cooled process, but due to the large quantity of flue gases the flue gas losses of the process are greater than those of other alternatives. As a consequence, the overall efficiency of energy production is somewhat lower. As a whole, the air-cooled process is quite complex, and according to a reference /3/ it is less economic than the other alternatives. It should, however, be borne in mind that the significance of the high power to heat ratio of cogeneration has not been properly considered in this comparison. As a more attractive alternative, however, can be considered a steam-cooled process provided with partial gasification.

The power to heat ratio of the process based on air-cooled fluidized bed is relatively low, 0.628. A mere 25 % of the total electric power output is generated by the gas turbine." The process has the advantage that the amount of flue gases to be cleaned is small. In addition, it is possible to select a higher ratio of the gas turbine than in a conventional air-cooled process since the operation of the exhaust-gas boiler is enhanced by additional combustion. A weakness of the process principle is that additional combustion passes the first part of the overall process. For this reason the process is neither economic with natural gas.

For comparison of the advantages achievable with pressurized combustion and combined-cycle processes a respective conventional cogeneration process was also evaluated. The scheme and essential process values of this process are given in figure 11. The process follows a similar principle as the steam cycles of PFBC-processes.

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Process values at dimensioning point

Steam cycle Gas cyclem P t h m p t hkg/s bar ’C kJ/kg kg/s bar *C kJ/kg

1 57.7 160 565 3475 1 97.5 1 15 153 57.7 42 565 3590 2 118 1 150 1704 2.60 4.5 315 30906 24.2 0.6 135 2758

16 57.7 160 190 810 0f -200 MW17 408 50 0dh - 120.0 MW19 408 120 Pst * 59,0 MW

a = 0.491

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Figure 11. Conventional cogeneration power process (Reference).

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The electric power output generated by the plant is 59 MW, and the district heating output 120 MW. This gives a power to heat ratio of 0.491. The overall thermal efficiency of the process is about 89 %.

4 THE ECONOMIC PROFITABILITY OF PFBC COMBINED CYCLE PROCESSES

From the utility point of view various PFBC processes have been compared with a conventional district heating process both in small and medium-size plants. It is assumed that the compared process alternatives satisfy the same heat demand, whilst the electric power output varies depending on the process scheme used.

4.1 Profitability of small PFBC-district heatingcogeneration plants

The process scheme of the compared small-scale PFBC-district heating plant is shown in figure 1. The profitability of the PFBC combined cycle process is compared with a conventional back-pressure power plant process, shown in figure 2. Its fuel input is of the order of 50 MW in all cases.Based on the duration curve of figure 12 the total energy generated over the year is calculated. The generated district heating output and energy are the same for all the compared alternatives: the maximum output 28 MW and the produced annual energy 146 GWh/a. The maximum electric power output of the conventional cogeneration process is 11 MW, and the produced annual energy is 53 GWh/a. The respective values for the PFBC-combined cycle processes are 14 - 16 MW and 70 - 75 GWh/a. The duration curves of the electric power generation with respective process alternatives are shown in figure 13.

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56 MfWOPKBC 60

PFBC 50

Load factor3150 h/m Conv. back-pr

<Ref«r«nc»)28 MW

Lead factor

4000 6000 t(h>

Figure 12. Duration curve of a figure 13. Duration curves of the district heat demand. power generation with

different power alter­natives.

Allowed investment costs, FIM/kW

Electricity price

18000

16000

14000

12000

10000

8000

6000

Fuel price, FIM/MWh

Hs = 160 mk/MWh -°- Hs - 180 mk/MWh * Hs - 200 mk/MWh

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Figure 14. The allowed additional investment costs of the PFBC-combined cycle cogeneration power plant as a function of the fuel and electricity prices.

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The supposed fuel is peat with a moisture content of either 50% or 60% alternatively. In the case of 50% peat (PFBC 50) the process is dimensioned for peat of 50% moisture content. The electric power output is 15.1 MW with the power to heat ratio of 0.53. The overall thermal efficiency is 86% and the fuel input is 50 MW.

In the case of a 60% moisture content (PFBC 60) the rated electric power output is 16.3 MW with the power to heat ratio of 0.58. The overall thermal efficiency is 85% and the fuel input 52 MW.

The highest output of electric power is generated by the PFBC 60 process, the power output of which is 16.3 MW and the annual energy 76 GWh/a. The respective values for the reference plant are 11.3 MW and 49 GWh/a.

The economic comparison of the different alternatives is shown in table 1.

The fuel price is depending on the moisture of peat.The price of generated surplus electric power has been assumed to be the medium price of electricity at the Finnish national grid, in 1986 approximately 160 FIM/MWh.

The biggest profit can be achieved by the PFBC 60 alternative, in which the price of the 60% moist peat is 35 FIM/MWh. The total profit over the year is approximately FIM 6 million/a compared with the reference plant. The annual profit of the PFBC 50 alternative is FIM 2.2 million/a.

The allowed cost for the additional investment of the PFBC 60 plant is FIM 1075/kW ‘a when the fuel price is FIM 35/MWh. The corresponding cost for the PFBC 50 alternative is FIM 505/ kW‘a.

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Table l Energy economical comparison of peat firedPFBC combined cycle and conventional district heating cogeneration power plants.

Conv.plant

PFBC-50 PFBC-60S PFBC-60b

Fuel price FIM/MWh 48,0 48,0 43.5 35Electricity price FIM/MWh 160 160 160 160

Fuel consumption GWh/a 226,83 249,12 260,88 260,88Fuel cost, total MFIM/a 10,668 11,958 11,348 9,131Fuel cost decrease MFIM/a - -1,070 -0,460 +1,757

Electricity output max. KW 10438 14724 15995 15995Annual electr. generation MWh/a 49182 69399 75524 75524Electricity output increase kW - 4286 5557 5557Electr. gener. increase MWh/a - +20217 +26342 +26342Electr. incomes increase MFIM/a * +3,235 +4,215 +4,215

Net coverage returns MFIM/a . +2,164 +3,754 +5,972Net unit cover, returns FIM/kW.a - 505 676 1075

Profitable additional investment costs

(i = 5%;T«20a) FIM/kW 6292 8419 13993MFIM * 26,69 46,84 74,27

250

Gener.costsFIM/MWh

A. Coal fired conventional condensing plantB. Peat fired conventional back-pressure plantC. Peat fired combined cycle cogeneration plantD. Nat. gas fired combined cycle cogener. plant

Flgun 15. Electricity genentlon costs ofdlffennt power plant processes.

- 28 -

Figure 14 shows the allowed investment costs as a function of the price of peat with 60% moisture content using the electricity price as a parameter. The assumed operating time is 20 years and the rate of interest 5%. The strong effect of the electricity price on the allowed investment costs is visible.

4,2 Profitability of medium-size district heating powerplants

The profitability of medium-size (200 MW) combined cycle cogeneration plants was evaluated for the case of a steam- cooled PFBC with partial gasifier (figure 9).

A cost estimate for the peat-fuelled cogeneration power plant is shown in table 2, The compared process alternatives are a 500 MW coal-fired condensing power plant with desulphurization, a conventional 60/120 MW peat-fired cogeneration power plant, and a natural gas fuelled combined cycle cogeneration system. The annual costs of electricity generation with different power plant alternatives are illustrated in figure 15.

According to the comparison in table 2 and in figure 15 the combined cycle power plant fuelled with milled peat is clearly more economic than the coal-fired condensing power plant, and generates electric power at about the same price as the conventional peat-fired cogeneration power plant with equal heat output. If we consider the higher electric power output of the plant with pressurized combustion, and the difference has to be covered by condensing power, the pressurized combustion, instead, makes it possible to get 10-15% lower overall production costs.A natural gas fuelled combined cycle power plant is, however, the most economic at current gas prices.

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29

Table 2. Comparison of power and heat production costsbased on different process alternatives.

ProcessFuel

Steam turbine plant Gas/steam turbine plantHard coal Milled beat Milled oeat Natural oas

Plant capacity- electric power output MW 500 60 107 104- district heat output MW 0 120 114 108

Process efficiency- electric power % 40,3 81,3 76,9 78,1• district heat % 0.0 87,0 89,3 90,9- demand of supp. fuel % 3.0 8,0 0,0 0.0

Investment costs- machineries MFM 1289,7 186,8 363,2 185,5• buildings MFM 241,6 63,9 63,4 61,8• other costs MFM 561,6 91.1 154,7 89,9Total investment MFM 2092,9 341,8 581,3 337,2- Sales tax compens. MFM 0,0 •36,9 -63,2 -36,4Capital demand MFM 2092,9 304,9 518,1 300,8

Energy production costsAnnua/ fixed costs- capital costs MFIM/a 148,5 21,6 41,6 24,1-other fixed M FIM/a 90,3 13,2 16,1 11.8Total fixed, actual MRM/a 238,8 34,8 57,7 35,9- from it heat gener. MFIM/a 0.0 13,5 5.4 5.4- from it power gener. MFIM/a 238,8 21,3 52,3 30,5Fixed costs, comparable- heat generation MFIM/a 0.0 13,5 12,8 5,4- power generation MFIM/a 238,8 21,3 44,9 30,5Fixed unit costs, comparable- district heal FIM/kW.a 0,0 112,5 112,5 50,0- electric power FIM/kW.a 477,6 408,9 419,4 293,3

Variable costs -Fuel price-main fuel FIM/MWh 39,0 40,0 40,0 47,0- supporting fuel FIM/MWh 47.0 47.0Variable unit costs, actual- district heat FIM/MWh 0,0 48,4 46,6 46,7- electric power FIM/MWh 102,1 51,7 54,1 54,4Variable unit costs, comparable- district heat FIM/MWh 0,0 46,6 46,6 46,6- electric power FIM/MWh 102,1 73.8 54,1 54,4

Comparable energy production c ostsDistrict heat FIM/MWh Lf-4000 h/a 75 75 59

Lf-5000 h/a 69 69 57Lf-6000 h/a 65 65 55Lf-7000 h/a 63 63 54

Electric power FIM/MWh Lf-4000 h/a 222 176 159 128Lf-5000 h/a 198 156 138 113Lf-6000 h/a 182 142 124 103Lf-7000 h/a 170 132 114 96

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4.3 Profitability of industrial cogeneration power plants

The competitiveness of small industrial PFBC power plants is considerably better than the figures above, due to a higher relative growth of the power to heat ratio.

By complementing an existing industrial back-pressure power plant with a preceding PFBC boiler and gas turbine the achievable additional electricity output will cost FIM 6400/kW in investment.The average cost of generated electricity will be FIM 60/MWh with 5% interest and 20 years' operating time. If a reinvestment or new capacity are concerned, the additional need for capital drops down to 3 000 FIM/kW, corresponding to 28 FIM/MWh.

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5 MARKET POTENTIAL OF PFBC IN FINLAND

New fuels for combined cycle power plants, being of interest in Finnish conditions, are wood waste and black liquor in the pulp and paper industry, peat at heating power plants and peat or coal at condensing power plants. When natural gas is available, the process based on pressurized combustion of solid fuels is not yet competitive, in the next 10 years, excepting the use of process fuels (waste wood and liquors).

In the long term, the combined cycle power plant processes using solid fuels seem to be competitive even in Finland. However, it is still feasible to build natural gas combined cycle technology, if a gas net is available. By replacing natural gas with eg. gas produced from coal a significant proportion of the savings obtained from the above price difference can be put into investments.

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- 31 -

In the short and medium terms possible objects o£ use include the wood waste boilers of the pulp and paper industry and the district heating power plants outside the natural gas network. Possible objects of use in the long term are all the base load power plants with a thermal output of 50 MW or more.According to the Finnish energy statistics the industry consumed in 1985 1.7 million toe of waste liquors, i.e. 19 TWh, which is equal to 1070 MW of fuel input.

Until the year 2000-2010 the fuel input obtained with black liquor will rise up to 2700 MW and that obtained with wood waste up to 1200 MW, if the annual production of chemical and mechanical pulp will grow in accordance with development programs for the pulp and paper industry.

Along with the rationalization of the pulp and paper industry the production of chemical pulp will be concentrated on estimated 12-14 plants. The average waste wood and black liquor input per plant will be about 300 MW. At the moment the average size of recovery boilers is less than 100 MW. The possible fuel input of the largest industrial power plant would be as high as 650 MW.

It is possible to build a maximum of 12-15 plants with a fuel input of 20-60 MW in conjunction with paper mills using mechanical pulp. The overall thermal output can rise in the largest units up to 100 MW.

]

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EE

In Finland, in 1983, the total thermal capacity of boilers with over 50 MW was 19 GW. Of this, the proportion of base load power was perhaps 12-14 GW. In the long term, thisbase load capacity and the thermal capacity of newcondensing power produced with fossil fuels is thepotential to which the pressurized combustion can beapplied.

- 32 -

The theoretical electric power generation capacity of the industrial back-pressure power plants is today about 1500 MW of which about 1000 MW is premium capacity. The figures indicate about 5000 MW to process heat output. This figure will not grow to a decisive degree. The back-pressure power output of the district heating power plants is about 1700 MW corresponding to about 3500 MW of district heat output. Until the year 2000 the output will grow to 4300-4500 MW. The thermal output of the back-pressure power plants will be 9-10 GW on the level of the year 2000.

As the combined cycle power plant involves a notable rise in the power to heat ratio, the theoretical additional capacities of back-pressure power until the years 2000-2010, compared with the conventional steam power plants, will be as follows:1) industrial back-pressure power plants 2000 MW2) municipal cogeneration power plants 2500 MW

In reality, construction activities are decisively affected by

1) ageing of existing units,2) additional need of capacity, and3) competing fuels (natural gas).

Pressurized combustion based on peat would become topical in 2000-2010 in twelve towns in northern and central Finland.Coal-fired power plants based on pressurized combustion could be considered in three towns on the western coast of Finland.

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By including the process fuels of the pulp and paper industry in the possible potential of the year 2010 and the fuels outside the natural gas network as well as the municipal power stations within the same area, pressurized combustion makes it possible, in principle, to increase power production according to table 3 by the year 2000-2010. The figures include the hypothesis that, by the year 2010, 3/4 of the pulp mills utilize the process fuels in a gas turbine process, i.e. ageing of the plants has been criteria for classification.

Hence, it is in theory possible to produce about 1400 MW more cogeneration power with pressurized combustion compared to conventional solutions.Thus the electric power capacity produced by pressurized combustion would be 2950 MW.In the production of condensing power the pressurized combustion may also be a competitive alternative in 2000-2010.

6 CONCLUSIONS

In the future, the pressurized combustion will offer a com­petitive method of energy production which is suitable for the environmentally beneficial utilization of low-grade fuels with a good overall efficiency. In addition, by utilizing pressurized technology it is possible to improve considerably the power to heat ratio of cogeneration power plants, which, again, makes this type of plants well-suited for an alternative form of future energy production also in Finland.

The most characteristic techno-economic advantages of combined cycle processes in comparison with respective steam processes are shown in table 4.

34 -

Table 3. Electric power generating potential of PFBC/G processes until the year 2010.

(Electric MW, in brackets increase compared to conventional cogeneration power plants, MW)

Unit size (fuel input) < 100 MW 100...300MW >300 MW

Fuel Unitnr.

Grossoutput

Unitnr.

Grossoutput

Unitnr.

Grossoutput

•Peat and bark 5 300 (150) 3 200 (100) 4 700 (350)

Coal - - - - 3 600 (300)

Black liqueur* - - 8 930 (370) 1 220 (120)

Total 5 300 (150) 11 1130 (470) 8 1520 (770)

' Black liqueur alone or together with other fuels

Table 4. The impact of the combined cycle process on the overall economy of power plant processes.

Steamprocess

Combined cycle process

Separate electricity generation plant

- efficiency % 35...40 40...50

Industrial cogenaratlon plant

• power to heat ratio - overall efficiency %

0,20...0,2580...90

0,5...0,780...88

District heating cogeneration plant

- power to heat ratio- overall efficiency %<

0,4...0,585...90

0,55...1,2584...88

35 -

The power to heat ratio of a PFBC cogeneration power plant is considerably higher than that of a conventional one. An increase of the gas turbine inlet temperature or the pressure ratio has a positive effect on the power to heat ratio. Additionally, an increase in the power to heat ratio can be achieved by using wet fuels.In a comparison of different combined cycle cogeneration power plant processes utilizing pressurized fluidized bed technology it was found that air-cooled partial gasification and a steam-cooled PFBC-process with partial gasification are the most economic in particular if the thermal capacity of the plant is at least 200 MW. The power to heat ratio achievable with a partial gasification process is 1.25 resp. 1.0, whereas with an air- or water-cooled PFBC-process without afterburning it is less than 0.7.

According to economic calculations the fixed unit costs of additional electric power generating capacity achieved by pressurized combustion are less than those of coal-fired condensing plants, although the fuel consumption in the cogeneration process drops to half. The economic competitiveness of pressurized combined cycle power plants is also good if compared with conventional cogeneration power plants.

According to the conducted market survey, it will be theoretically possible, by the year 2010, to build about 1400 ’ MW more cogeneration power capacity using the pressurized combustion compared with the conventional solutions. A prerequisite is, however, that the process operates technically with various types of fuel. Thus the electric power capacity of power plants operating on pressurized technology would be 3000 MW. The PFBC combined cycle power plants can also be a competitive alternative in the production of condensing power after the year 2000.

36 -

REFERENCES

1. Hulkkonen, S., Makela, J., & Jahkola, A.: The properties and economy of the 50 MW peat burning PFBC district heating power plant. TKK/EVO 1986.

2. Olesen, C., et al.: The commercial status of ASEA-PFBC technology. 8th Int. Conf. on Fluidized Bed Combustion. Houston, Texas, 1985.

3. Huber, D. A. & Costello, R. M. : Conceptual desing andeconomic evaluation of various fluidized bed power plant concepts. Pressurized Fluidized Bed Combustion Technology Exchange Workshop. DOE Conf-7906157, 1980.

4. Jahkola, A., Hulkkonen, S., Horvath, A., Boij, J., Kurkela, E., Mojtahedi, W., Saviharju, K.: Development of pressurized fluidized bed combustion. Research Report Otaniemi 1987. Helsinki University of Technology, Department of Energy Engineering. Report 22, 213 p (In Finnish)

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GASIFICATION COMBINED CYCLE POWER GENERATION - PROCESS ALTERNATIVES

Esa Kurkela, M.Sc., Senior Research Scientist

Technical Research Centre of Finland

Laboratory of Fuel Processing and Lubrication Technology

02150 Espoo, Finland

Abstract

Interest in Integrated Gasification Combined Cycle (IGCC) power plants

has recently increased also in Finland. The IGCC systems offer the

potential of superior efficiency and environmental performance over

conventional pulverized coal or peat fired boilers. Potential applica­

tions are both large-scale electricity production from coal and medium-

scale combined heat and electricity production. In the latter case, the

gasification process should also be applicable to peat and wood.

Several IGCC processes have been developed in USA and in Europe. These

processes differ from each other in many respects. Nearest to commer­

cialization are processes, which employ oxygen gasification and cold

gas cleanup. The Cool Water plant, which was brought into operation in

1984 in USA, has demonstrated the feasibility of an IGCC system using

Texaco entrained-bed gasifier. Several pressurized fluidized-bed and

fixed-bed gasification processes have also reached a pilot or de­

monstration stage with a wide variety of coals from lignite to hard

coal. Pressurized fluidized-bed gasification of peat (Rheinbraun-HTW-

process) will also be demonstrated at the peat ammonia plant of Kemira

Oy, which will be commissioned in 1988 in Oulu, Finland. Oxygen gasifi­

cation and cold gas cleanup are, however, economically viable only in

large-scale applications. Technology is being developed to simplify the

IGCC system, in order to reduce its capital costs and increase its

efficiency. Air gasification combined with hot gas cleanup seems to

have a great potential of improving the competitiveness of the IGCC

system.

r r iil. _-is. _Jl

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1 GENERAL

The use of solid fuels in gas turbine processes is based either on

pressurized combustion or on gasification. Combined processes are also

under developement. Pressurized fluidized-bed combustion (PFBC) has

been considered the best alternative of the combustion methods and

other combustion methods have hardly developed at all, whilst the deve­

lopment of processes based on gasification includes several alter­

natives of different type. The aim of ventures underway in the United

States and Europe is most often the production of electricity in large

(> 300 MWe) coal-fired power plants. Advantages of the gasification

combined cycle processes are primarily smaller emissions and higher

efficiency than those of the conventional power plants. In addition to

base load electricity generation from coal, co-generation of electri­

city and heat in the size class of < 200 MWg is of special interest in

Finland. The gasification process should also be capable of using indi­

genous fuels, such as peat, wood and different waste fuels.

The Gasification combined cycle plants have advanced closer to a com­

mercial level than those based on fluidized-bed combustion. A plant

based on air gasification operated in Liinen, F.R.G. in the years 1973 -

1982, and an industrial-scale demonstration plant based on oxygen gasi­

fication was put into operation by Texaco in Cool Water, U.S.A., in

1984. However, the advancement of the gasification combined cycle pro­

cess to a level that is more profitable than the other electricity

production alternatives still requires the development of both gasifi­

cation processes and gas turbines.

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2 THE IGCC SYSTEM

2.1 Process alternatives wi

In conventional power plants (Figure 1) the solid fuel is combusted

under atmospheric pressure and steam is generated; electricity is pro­

duced at an overall efficiency of 36 - 39 % and the flue gas is cleaned

\

- 39 -

AtmosphericCombustion

Pressurised Fluidised Bed Combustion

Gasification Combined Cycle

Air/Oxygen Solid fuelAir Solid fuelAir Solid fuel

Steam turbine

Gasturbine Steam

turbineSteamturbine

Flue gas CondenserFlue gas CondenserFlue gas

GasCleaning

HeatRecovery

HeatRecovery

PressurisedGasification

Pressurised.Combustiorv

GasCleaning

AtmosphericCombustion

HeatRecovery

GasCleaning

Boiler

i r i

Figure 1. Schematic diagrams of a conventional power plant and ofpressurised fluidised-bed combustion (PFBC) and integrated gasification combined-cycle (IGCC) power generation schemes.

at about 150 °C to protect the environment. In an Integrated Gasifi­

cation Combined Cycle (IGCC) power generation scheme the solid fuel is

first gasified under pressure. The produced fuel gas is first cleaned

and burned in the combustion chamber and then expanded in the gas

turbine.

The main subsystems of the IGCC system are gasifier, heat exchanger,

gas clean-up, gas turbine generator set, heat recovery boiler, and

steam turbine generator set. To maximize the efficiency of the plant,

these subsystems should be carefully integrated.

The IGCC plant can be constructed in many different ways, which deviate

from each other in details (Figure 2). The most significant factors

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DEGREE OF REACTOR METHOD OFGASIFICATION TYPE HEATING

V-A.-

OXYGEN

INDIRECTHEATING

FLUIDIZEDBED

ENTRAINEDBED

FIXEDBED

OTHERTYPES

TOTALGASIFICATION

PARTIALGASIFICATION

IE

Figure 2. Gasification methods in IGCC systemsT-i; r

affecting the efficiency of the whole process and its economy are:

degree of gasification: total and partial gasification,

method of heating: air gasification, oxygen gasification and

indirect heating,

reactor type: fixed-bed, fluidized-bed, entrained-bed,

gas clean-up: cold and hot clean-up,

properties of the gas turbine,

properties of steam process.

Estimates of the efficiency of the PFBC and IGCC systems presented in

various references /4, 12, 13, 14/ are shown in Figure 3. The turbine

inlet temperature has a fairly crucial effect on the efficiency of the

IGCC plant. The superiority of hot gas clean-up to cold one is based on

the fact that the efficiency available from the gas turbine process

increases, as it is not necessary to cool the gas.

I'l.

- 41 -

EFFICIENCY (LHV)

% 55

50

45

40

35

30

COMBINED CYCLE SYSTEM BASED ON NATURAL GAS IGCC-SYSTEMS WITH HOT GAS CLEANUP IGCC-SYSTEMS WITH COLD GAS CLEANUP

I I ]

GAS TURBINE INLET TEMPERATURE

Figure 3. Efficiency comparison of powergeneration systems.T r -j

2.2 Gasifier types

There are three main types of gasifiers: fixed-bed, fluidized-bed, and

entrained-bed reactors (Figure 4). There are experiences available from

several commercial applications. The present stage of the most well-

known processes is presented in Table 1.

Fixed-bed gasifiers

The Lurgi process developed in Germany before the World War II was the

first pressurized gasifier. Lurgi reactors have been used industrially

in various places, for example, at the Sasol plant in South Africa and

in a Great Plains project in the United States, where the production of

synthetic natural gas of high calorific value was demonstrated /3/.

Lurgi gasification was also used in the IGCC demonstration plant in

Liinen, which operated in 1973 - 1982. In this process the lump coal is

- 42

FIXED BED FLUIDIZED BED ENTRAINED FLOWPRESSURE 100 bar 25 bar AO barCAS TEMPERATURE 700 - 1100 C600 C 1000 - 1500 C

(5 - 50 (0-5 inn)

FINES FINES

0,+ STEAMAIR/O.

HT-HINKLERU-GASKRU

TEXACOPRENFLOSHELLVEW

PROCESSES LURGI (dry ash) BGC-LURGI

Figure 4. Gasifier types and processesf r■r- _j

Table 1. Gasification processes: stage of development.

PROCESS development statusPILOT DEMO-PLANT COMMERCIAL

FIXED BEDLurgi, dry ashBritish Gas - Lurgi Hestfleld,CB(500t/h>

SASOL(S.Africa,44000t/d)

FLUIDIZED BEDHT - HinklerKRHU - GasMINO

Hachtberg,FRG(20t/d) Waltz Mill.USA(200t/d)Ch1cage,USA(20t/d>Studsvik,SHE(12t/d)

Berrenrath,FRG(720t/d) Kemira,FIN(650t/d)

ENTRAINED FL(H(Texaco

Prennflo

Montebel1o,USA(20t/d)

Furstenhausen,FRG(48t/d)

Oberhausen,FRG(150T/d) Tennessee,USA('83)UBE,Japan!’84)Cool Hater,USA<’84) etc.

Shell Amsterdam,Netherl. Hous ton,USA(150t/h)VEW Herne,FRG(12t/d) Herne,FRG(240t/d)

ri

- 43 -

pressurized in a multiphase lock hopper and fed to the upper zone of

the reactor. The mixture of air or oxygen and steam is fed to the lower

zone of the gasifier. The fuel flows slowly downwards through drying,

pyrolysis, reduction and combustion zones. As a consequence of the

counter current flow of the gasification material and the gases, the

temperature of the gas is low (< 600 °C for coal, < 200 °C for peat)

and the major part of energy is obtained as chemical energy.

However, an abundance of condensing tars are developed in the process,

and therefore, cold gas clean-up and related waste water treatment

would be a complicated operation. On the other hand, the particle and

alkali content of the gas can be fairly low, and hence, the alternative

of hot gas clean-up is an interesting potential.

The most wellknown fixed-bed process designed for combined cycle appli­

cations is British Gas Slagging Lurgi, where the temperature of the

lower gasifier zone is kept above the ash melting temperature. The

process may be suitable, in addition for coal gasification, also for

piece-shaped peat products, although tar separation would be more

troublesome. The aim of British Gas is to apply this process in the

size class of > 200 MWe and to use o*ygen/steam gasification and cold

gas clean-up.

Entrained-bed gasifiers

In these gasifiers finely pulverized coal and air or oxygen are led to

a reactor, where gasification takes place at > 1 200 °C temperature.

The residence time is extremely short. As a consequence of the high

temperature the CO and H2 contents of gas are high and tar content

insignificant. The ash is removed molten. The high carbon conversion

requires dry and finely pulverized fuel. The fuel is fed as slurry to

the pressurized reactor in most processes. The pressurized entrained-

bed gasifiers do not suit particularly well for gasifying biomass or

peat, as these fuels require pretreatment to a considerable extent and

the feed causes problems. Due to the high temperature of gas, the major

t. I. 1

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part of the fuel energy is obtained as sensible heat of the gas. Hence,

the chemical efficiency is lower than in fixed-bed processes.

The most significant commercial or developing entrained-bed gasifica­

tion processes suggested for combined cycle power plant applications

are Texaco, Prenflo, Shell and VEW. The last-mentioned one is a partial

gasification process. For the time being, the process of Texaco, based

on slurry feed and multiphase gas cooling and heat recovery, has been

the most successful one. Before the construction of the Cool Water com­

bined cycle power plant, experiences had already been obtained from a

number of synthesis gas plants. The slurry feed, which decreases effi­

ciency and increases oxygen consumption, and expensive heat recovery

equipment can be considered to be disadvantages of the Texaco process.

The Prenflo process of Krupp-Koppers /3, 5/ was developed from the

unpressurized commercial Koppers-Totzek process. Development work was

done in 1974 - 1981 in co-operation with Shell at a 150 t/d pilot plant

in Harburg, F.R.G. Krupp-Koppers has continued to develop the process

by constructing a new 48 t/d demonstration plant in Furstenhausen. In

this process, dry coal is fed pneumatically to the reactor.

Fluidized-bed gasifiers

In these reactors the fuel, particle size 0 - 10 mm, is typically gasi­

fied at 750 - 1 000 °C. The method suits in particular for reactive

fuels, such as lignite, biomass and peat.

The most wellknown pressurized fluidized-bed processes are the High

Temperature Winkler (HTW) developed by Rheinische Braunkohlenwerke AG

and Uhde GmbH, F.R.G., the U-gas process developed by the Institute of

Gas Technology, U.S.A., and the KRW process originally developed by

Westinghouse, U.S.A. All these processe employ bubbling bed reactors.

Lurgi and A. Ahlstrdm Oy have developed circulating fluidized beds,

using a higher rate of flow. However, Lurgi and Ahlstrom have so far no

pressurized equipment. The difference between the slow bubbling

fluidized bed and the circulating fluidized bed is apparently becoming

- 45 -

smaller, as there are solids recirculating systems in all these pro­

cesses, and the tendency is to change over to higher fluidizing veloci­

ties also in HTW and U-gas processes.

A special feature of the U-Gas and KRW processes is agglomeration based

on partial melting of ash. The methods were developed in particular for

fuels with a high ash content, when good carbon conversion requires an

efficient separation of ash and residual carbon. The U-Gas process has

been proved to operate also with peat in pilot-scale experiments /8/.

The U-Gas and KRW processes have not been demonstrated commercially so

far.

2.3 Partial gasification

In total gasification the aim is to convert the fuel totally to gas.

The gasification consists of pyrolysis by heat and of reactions of the

residual carbon with gasification gases. The last-mentioned stage re­

quires oxygen inlet into the process, usually as air, pure oxygen or

steam. As the gasification of the residual carbon is a rather slow

reaction, it is difficult to secure sufficient conversion of carbon in

certain gasification methods, especially in such fluidized-bed re­

actors, in which agglomerated ash removal is not applied.

Partial gasification is a process, where part of fuel energy is re­

covered as residual carbon from the reactor. The degree of gasification

can in principle range from slight pyrolysis (part of volatiles is

removed) to total gasification. The heat of endothermic drying and

pyrolysis stage can also be introduced in many different ways. In

combined cycle power production, the simplest way of oxygen input is

partial gasification by air. Compared to total gasification, the only

difference is that carbon conversion is not maximized.

Two different integrations of partial gasification to a power plant are

shown in Figure 5. In alternative a) the bottom ash from a pressurized

fluidized-bed gasifier and the fly ash separated in cyclones are fed to

- 46 -

AIR/OXYGEN FUEL

1___ LAIR

PRESSURIZEDGASIFICATION

ASH +

DUST

FUEL GAS CLEANING

AIR FUEL GAS COMBUSTION

rATMOSPHERICCOMBUSTION

ASH

FLUEGASCLEANING

FLUE GAS

a

GASTURBINE

HEAT T-n- RECOVERY ,

FLUEGAS

©

STEAMTURBINE

CONDENSER

FUEL b

ASH +

CHAR

GASTURBINE

STEAMTURBINE

FLUEGAS

CONDENSER

HEATRECOVERY

FUEL GAS COMBUSTION

PFBC

HOT GAS CLEANUP

GASCLEANUP

PARTIALGASIFICATION

Figure 5. Two Integrations of partial gasification to a power plant.a) IGCC combined with atmospheric combustion,b) PFBC combined with partial gasification.

- 47 -

an atmospheric fluidized-bed boiler, where combustion is completed.

Alternative b) is based on that weakness of the PFBC process that the

inlet temperature of the turbine cannot be higher than that of ash

deformation. This restricts the available efficiency. As part of the

fuel is gasified, the inlet temperature of the turbine can be raised to

the allowed maximum temperature by combusting the fuel gas.

3 EXPERIENCES FROM THE COOL WATER DEMONSTRATION PLANT

The Cool Water coal gasification demonstration plant was commissioned

in 1984. A five-year demonstration project is underway, the following

enterprises participating in the work: Texaco Inc., Southern California

Edison Company, General Electric, Bechtel Power Corporation, Electric

Power Research Institute (EPRI), and Japan CWP. A simplified block

diagram of the CWCG process is shown in Figure 6.

The system uses an oxygen-blown Texaco entrained-bed gasifier. Radiant

and convective heat exchangers cool the gas for cold gas clean-up and

raise saturated steam for superheat in a heat recovery steam generator

(HRSG). Particulates are scrubbed and the gas is cooled further for

sulphur removal by the Selexol process. The gas is then reheated and

saturated for combustion and expansion in a 65 MW General Electric

Frame 7F combustion gas turbine. Remaining energy is recovered in a

HRSG to produce superheated steam for a 55 MW steam turbine.

The reports on the performance of the Cool Water plant have been ex­

ceptionally good /l, 3, 10, 15/. The environmental performance has in

particular exceeded all targets. The measured emissions of the CWCG

plant compared to U.S. New Source Performance Standards for Coal-fired

Utility Plants are shown in Figure 7.

48

OEttINfftAUZED WATER

STACKGAS

IFWSYSTEM

COAL RECEIPT STEAMPUNT TURGINEIS.*.)

EXHAUSTSTEAMSTEAM

CLEANSYNGAS

CLEANSYNGAS

SULPUR SYNGASSYNGASCOOLING SATURATIONCOOLING

SOUR jACIO SATURATOR■LOWOOWN

STACK

ASH 4 WATER GAS

PUSH SULPUR

recycleT

TO DISPOSALHANOLING

EFPLUENTWATER

TREATING

QtfcSEL (ALT FUEL)

Figure 6. Cool Water coal gasification program, bloc flow diagramnot.

i r

y Tj

S02 (3 % wt S) S02 (0.5JS wt S) PARTICULATES

(X I U S - N S P S E3 COOL WATER

I E 1

wt

^' iure 7. Environmental performance of the Cool Water demonstrt.; plant /3, 10/.

- 49 -

4 ADVANCED IGCC SYSTEMS

The Cool Water Program has succesfully demonstrated on a commercial

scale an IGCC system, which 1s based on oxygen gasification and conven­

tional cold gas clean-up. Oxygen gasification and cold gas clean-up

are, however, economically viable only in large-scale applications.

Cooling and scrubbing the fuel gas also means a reduction of efficiency

compared to hot gas clean-up systems. So, technology is being developed

to simplify the IGCC system, in order to reduce its capital costs and

to increase its efficiency. Air gasification combined with hot gas

clean-up seems to have a great potential of improving the competiti­

veness of the IGCC system (Figure 8).

Advantages of hot gas clean-up compared to the conventional IGCC con­

cept are:

higher thermal efficiency,

simpler process,

no waste water production,

low investment costs in medium-size power plants.

However, there are still a number of unsolved technical problems

related to hot gas clean-up. Development work is required at least in

the following subjects:

temperature and corrosion resistant filter materials (for parti­

culate removal),

in-bed desulphurization and/or hot sulphur removal processes,

alkali control methods,

N0x control methods.

It is possible to remove about 70 - 85 % of sulphur in the fluidized-

bed gasifier by using dolomite. This has been demonstrated by experi­

ments carried out at the KRW fluidized-bed pilot plant in Waltz Mill,

Pennsylvania, U.S.A. /!/. However, if the aim is to reach the same

level as in conventional scrubbing processes (> 90 % desulphurization),

the H2S content of the fuel gas should be reduced further in separate

desulphurization equipment. Possible sorbents are, for example, iron

I

IT]

r ' T ’■ *

tflo

■q.jre 8. Comparison of eMSt'Ing and advanced IGCC sr—em-

f 4

t i rz* H

NM

- 51 -

oxide and zinkferrite (ZnFe204), the use of which has been studied at

the Morgantown Energy Technology Center /l, 12/.

Particulate removal and alkali control at high temperature is discussed

more closely in other papers of this publication. The major part of

fuel nitrogen is transformed into NH3 in gasification. This cannot be

removed in hot gas clean-up. Ammonia has a tendency of develop fuel N0x

in combustion. Hence, special design of combustion chambers may be

required to keep the N0x emissions low.

LITERATURE

1. Wieber, P. R. & Halow, J. S. Advanced IGCC power systems for the

United States. Energy Progress 7(1987)2, p. 119 - 125.

2. Thompson, B. H. et al. Conversion of coal to electricy by gasifi­

cation combined cycle systems - overview of GCC developments in

Europe. AICHE Nat. Meeting, Boston, 24.-27.8.1986. 25 p.

3. Conn, A. L. The integrated gasification combined cycle power

plant - power from coal with minimum environmental problems - an

American view. Energy World (1986) December, p. 5 - 12.

4. Nitschke, E. Combined-cycle power generation - a promising alter­

native for the generation of electric power from coal. Symp. on

Coal Gas - a Clean and Efficient Alternative for the Generation of

Electric Power, Stockholm 18.9.1986. 23 p.

5. Pohl, H. C. & Rohm, J. K. Pressurized entrained flow gasification

and its application to combined-cycle power plants. Palo Alto

1985, Electric Power Research Institute, EPRI AP-4257-SR, Vol. 1,

Section 20. 38 p.

6. Davies, H. S. et al. Application of the British Ga/Lurgi Slagging

gasifier for combined-cycle power generation. Palo Alto 1985,

Electric Power Research Institute, EPRI AP-4257-SR, Vol. 1,

Section 13. 32 p.

W ]

I T ]

T f 1JZ.

52 -

7. Weinzierl, K. Combined cycle with integrated coal gasification

VEW coal conversion process. Palo Alto 1985, Electric Power

Research Institute, EPRI AP-4257-SR, Vol. 1, Section 22. 25 p.

8. Kurkela, E. et al. Synthesis gas production from Finnish peat by

fluid-bed gasification. Energy from Biomass and Wastes X, April

7-10, 1986, Washington, D.C. Chicago 1986, Institute of Gas Tech­

nology. 15 p.

9. Corman, J. C. & Horner, M. W. Simplified IGCC attracts clean coal

cash. Modern Power Systems 1986: October, p. 37 - 41.

10. Clark, W. N. et al. Cool Water: performance & economics. Energy

Progress 7(1987)2, p. 99 - 104.

11. Salvador, L. A. Clean coal money staked on Applachian IGCC.

Modern Powpr Systems 1986: November, p. 33 - 37.

12. Finckh, H. H. & Muller, R. The development of integrated coal-

gasification power plants with clean combustion in Germany. New

York 1985, the American Society of Mechanical Engineers,

85-IGT-47. 75 p.

13. Vessonen, K. Kahden kivihiilta kayttavan kombivoimalaitossuunni-

telman analyysi (Analysis of two concepts of coal-fired combined-

cycle power plants). Diploma thesis. Espoo 1987, Helsinki Univer­

sity of Technology, Department of Mechanical Engineering. 75 p.

14. Solantausta, Y. Techno-economic evaluation of different gasifi­

cation-combined cycle processes. Espoo 1988, Technical Research

Centre of Finland, VTT Symposium 83. P. 199-21.2.

15. Keller, J. Development of coal gasification technology applicable

for combined-cycle power generation. Symp. on Coal Gas - a Clean

and Efficient Alternative for the Generation of Electric Power,

Stockholm 18.9.1986. 24 p.

IFk. i __

I F

t rSL -A.

IE

m

- 53 -

THE OTANIEMIPFBC/G TEST RIG AND RESEACH PROJECT

Seppo Hulkkonen M.Sc.(Eng), Antero Jahkola Prof., Esa Kurkela M.Sc.(Eng)*

Helsinki University of Technology

Technical Research Centre of Finland *

Otaniemi, 02150 Espoo, Finland

IT-A. ]

1 INTRODUCTION

The reseach project "Experimental Research of Pressurized Fluidized Bed Technology and

its Application to Combined Power and Heat Generation" was started in 1986. The project

is carried out in co-operation between Helsinki University of Technology (TKK), Technical

Research Centre of Finland (VTT) and Imatran Voima (TVO). The project is scheduled for

years 1986-1989 and is mainly financed by the Ministry of Trade and Industry (KTM).

The total costs of the project are 10 million FIM including the test facilities. The objective

of the project is to create knowledge of the basics of pressurized fluidized bed combustion/

gasification and their applications. The fuel of main interest is peat. The research is

realised partly theoretically based on literature studies and computer runs but primarily by

experimental measurements carried out using the PFBC/G test rig.

I

T T 1Jt. - -J

2 PFBC/G TEST RIG

The project was started by designing and building a PFBC/G test rig operating at pressures

up to 10 bar. The test rig was commissioned in summer 1987 and is located in Otaniemi in

VTT/POV Laboratory. The test rig can be used for combustion and gasification research

with thermal input of 130 kW and 400 kW, respectively. A general flow diagram of the rig

is shown in Figure 1.

The main part of the facility is reactor which is a 6 m tall vessel of 60 cm diameter.

Refractory lining brings the inside diameter down to 15 cm in bed and 25 cm in freeboard.

The bed height is 1.2 m and freeboard height is 2.3 m. The bed is cooled by vertically

oriented watertubes immersed in the bed. Hot gases leave the freeboard and pass through a

two-stage cyclone collector and then along high temperature ducting in which are located

the dust sampling system and main gas sampling points. The hot gases are then cooled in a

heat exchanger before pressure let-down valve which also controls the pressure in reactor.

I

WI

54 -

The system is pressurized using an electrically driven screw compressor, and fluidizing air

is admitted through a nozzle-type air distributor at the base of the bed. Air can be preheated

to 250 *C by electrical heater, if required. During start-up the fluidizing air is preheated to

900 *C using a propane burner. The main fuel is crushed peat pellets, which are continuously screw-fed from two pressurized hoppers.

GAS CLEAN .-P

LET-DOWNHOPPER

TKK VTT IVOPFftC/Q TEST RIG

STEAM

V.A

i r ;

T r

Figure 1. PFBC/G test rig.

The bed height can be varied by adding bed material from a pressurized hopper or by

discharging material through the air distributor. The main technical data of the rig are

presented in Table 1.

The test rig includes a wide range of measuring devices and sampling facilities.

Temperatures throughout the plant are measured using thermocouples and pressures are

measured by local pressure gauges and by differential pressure cells with electronic

pressure transmitters. The data are collected to a data logger and processed by a computer.

The rig, however, is operated mainly manually. Gaseous components (02, C02, CO,

S02, S03, NO, N02, CH4 , CxHy, NH3 and HCN) are measured from the flue gases, and

particle concentration and particle size distribution are also determined. Additionally, gas

and material samples can be extracted from the bed and freeboard. The measuring devices

and results in detail are discussed elsewhere.

- 55 -

Table 1. PFBC/G technical data.

Thermal input 130 kW (Gasification 400 kW)Fuel Peat pelletsMax. operating pressure 10 bar (abs)Bed temperature 850-950 'CFluidizing velocity 0.6-1.5 m/sBed material SandAirflow 140 m3n/hGas flow 160 m3n/h (Gasification 210 m3n/h)Fuel feed 30kg/h (Gasification 90 kg/h)Bed diameter 0.15 mBed height 1.2 mFreeboard diameter 0.25 mFreeboard height 2.3 mFuel feeding Screw feeding from pressurized hoppersGas cleanup Two cyclonesStart-up Propane burner

2.1 Operating experiences

The rig has so far been operated as a combustor for 150 hours during three test periods

including hot commissioning runs. The rig functions well and the planned operating

conditions are easily reached and maintained. The operation of the rig and measurements

have been performed around the clock basis by two 4-man teams working 12-h shifts.

The control of the rig is relatively easy. The temperature and pressure levels can be varied, which allows tests at different operating conditions. Feeding of peat pellets has functioned

properly, too. However, the moisture content of the fuel has been only about 20 %.

The biggest problem during the commissioning tests was too low a temperature of the gases

after water quenching which led to corrosion of the gas ducts. The corrosion problems

were solved by replacing the wet cooling system by a heat exchanger and keeping the gas

temperature above 200 *C.

3 RESEARCH PROJECT

T J

I l

The research project is divided into four subprojects:

a) Operation and maintenance of the rig. Feeding systems and safety aspects.

56 -

b) Pressurized fluidized bed combustion and hot gas clean-up.

c) Pressurized fluidized bed gasification and gas cleaning

d) Power plant processes and turbine technology

The responsibility for the subprojects a) and c) is taken by VTT, for b) by TKK, and for d)

by IVO. The project organisation is illustrated also in Figure 2.

EXPERIMENTAL RESEARCH OF PFB - TECHNOLOGY AND ITS APPLICATION TO COMBINED POWER AND HEAT GENERATION

IMATRAN VOIMA

MINISTRY OF TRADE AND INDUSTRY

TECHNICAL RESEARCHCENTRE OF FINLAND

HELSINKI UNIVERSITY OF TECHNOLOGY

(TKK)

D) POWER PLANT PROCESSES AND TURBINE TECHNOLOGY

B) PRESSURIZED FLUIDIZED BED COMBUSTION ANDHOT GAS CLEAN-UP

O PRESSURIZED FLUIDIZED BED GASIFICATION AND GAS CLEANING

A) OPERATION AND MAINTENANCE OF THE RIG. FEEDING SYSTEMS AND SAFETY ASPECTS

Figure 2. Project organisation and subprojects.

Subproject a) includes, besides operation and maintenance of the rig, also studies of fuel

feeding systems. Feeding systems will be reviewed using literature, and laboratory- scale

experiments with different fuels will be carried out. Pressurized fluidized bed combustion

research (b) includes studies of combustion and gaseous emissions at different pressures

using different fuels, and also studies of fluid bed hydrodynamics. Additionally, particle

elutriation is measured and some hot gas cleanup devices will be tested. Cyclone tests will

f f

57 -

probably be continued by ceramic filter tests. Pressurized gasification research (c)

includes, among others, basic research of gasification at high pressures and also

development of alkaline monitoring systems. Subproject d) includes computer modelling

of PFBC/G processes and also reviews of turbine technology.

The rig is planned to be used annually 6 monthts as a combustor and 6 months as a gasifier.

A general timeschedule of the project is shown in Figure 3. After the first combustion and

gasification test periods the rig will be completed by some secondary hot gas cleanup

device. The fuel in the first runs has been peat but coal will also be tested.

IKK VII ivn

PFBC/G

TIME SCHEDULE OF THEPFBC/G PROJECT

PLANNING OF THE RIG

BUILDING OF THE RIG

COMMISSIONING

COMBUSTION TESTS

GASIFICATION TESTS

i: i. i

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Figure 3. Project timeschedule.

r :c;

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I I

w1

- 59 -

PFBC RESEARCH AND DEVELOPMENT - BRITISH EXPERIENCES

H. Raymond Hoy, Director T’' WHoy Associates Ltd ^

Leatherhead, England

Abstract

The UK involvement in the development of PFBC is outlined with parti­

cular reference to the period up to the end of 1984.

1 BACKGROUND

The main motivation for developing PFBC is its potential for enabling

coal to be used in combined cycle power generating plant, thereby in­

creasing power generating efficiency, and at the same time meeting at­

mospheric emission regulations as regards sulphur and nitrogen oxides.

As is well known, a gas turbine converts heat into power at a higher

temperature than is practicable in conventional steam power plant. The

steam plant however rejects heat at a lower temperature, and hence

combining the two extends the range of temperature over which heat is

converted into power and hence increases power generating efficiency.

Combined cycle plant burning clean gas or liquid fuels are now being

used by a few power utilities, but for this system to be widely applied

the fuel will need to be coal.

Interest in the possibility of using coal as the fuel for power plant

involving gas turbines was first shown in the UK in 1947, and a number

of R A D programmes were intiated by the UK Ministry of Fuel and Power,

and by the Department of Scientific and Industrial Research.

Two direct combustion systems were investigated. In the first pulve­

rized fuels (75 - 80 % smaller than 76 microns), were burned and the

r r i

v r '

E

intent was to keep the combustion temperatures below those at which the

ash was molten. In the second, crushed coal was burned in a cyclone

combustor at temperatures at which the ash from the fuel was molten,

and from which over 80 °k of the ash was discharged as liquid slag.

The first route was followed at a number of establishments in the UK,

USA, and Australia, and following upon a 2000 hour test run at the US

Department of Energy's Morgantown research station in 1964 work was

discontinued. The turbine blades suffered from both erosion and accumu­

lation of ash deposits on the blades. The higher temperatures prevail­

ing in the cyclone combustion system resulted in greater volatilisation

of the coal ash and when in the early 1950s the English Electric Com­

pany operated a gas turbine at Rugby with this type of combustor, the

passageways through the turbine blades became choked with cohesive ash

after only a few hours operation. This approach was discontinued in the

late 1950s, and the author and his colleagues concluded that direct

firing of coal for gas turbines could not be seriously considered until

there was a radical new development in combustor technology.

This possibility arose with the evolution of fluidised bed combustion

in 1963 - 1964, which had the potential of achieving high combustion

efficiency at combustor temperatures below those likely to sinter the

ash and cause major volatilisation of alkalis, together with providing

a more compact heat exchange system, and the possibility of meeting the

anticipated limitations on emissions of S0X and N0X.

In the latter part of 1967 my colleagues and I inspired by the late

Prof. Douglas Elliott set about obtaining the necessary support to

build a pressurised combustor at the British Coal Utilisation Research

Association's laboratory in Leatherhead. The type of cycle envisaged is

shown in Figure 1.

The philosophy was to build the largest combustor with a rectangular

cross section that could be contained within an existing 1.8 m diameter

pressure vessel. The result was a combustor, Figure 2, with a cross

section of 1 200 mm x 600 mm. The maximum operating pressure for the

vessel was 6 bar.

r r

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61 -

Goa Clean up

ViSteam Turbine*

Fluidised bed boiler

_y

Condenser

I Gas to J.ta'k

Economiser

Ash

AirCompressor

]

i. i; i

Figure 1. Supercharged boiler combined cycle. Main application is for electric power plant larger than about 200 MWe. 20 - 30 % of total power would be provided by the gas turbine. Designs envisaged: 60 - 70 MWg gas turbine; pressure ratio 10:1 to 16:1; inlet temperature initially 800 - 850 °C; steam turbine pressure 170 - 240 bar; 538 °C superheat and reheat.

- 62 -

Figure 2. PFBC MK 1 1969 - 1973. 1. Start-up burner, 2. Bed cooling water inlet, 3. Primary cyclone, 4. Secondary cyclone, 5. Cascade, 6. Alkali measurement, 7. Coaxial gas/dust sampling probe, 8. Entry to probe, 9. Quench water sprays, 10. Outlet from venturi scrubber.

- 63 -

The primary purpose of the combustor was to obtain an indication as to

whether the combustion gases could, after cleaning in cyclones, be

expanded through a cascade of turbine blades without causing serious

erosion, or deposition that would adversely affect the aerodynamics of

the blades. A plant with the highest practicable gas flow was therefore

needed, and a combustor with a relatively large cross section was also

needed in order to obtain meaningful data for scaling up.

The test programme began in 1969, and in the period 1972 - 1981 several

other versions of the combustor were tested in the course of programmes

that were mainly sponsored by the US Department of Energy or their

predecessors, but which sometimes also involved such organisations as

Electric Power Research Institute (EPRI), General Electric, American

Electric Power, Stal-laval, British Columbia Hydro, Fluidised Com­

bustion Contractors Limited, and the Grimethorpe PFBC Project.

In a total of over 2000 hours operation information was obtained on

factors affecting, combustion and sulphur retention efficiences, N0x

emissions, heat transfer, combustor shape, and on the effect of the

coiribustion gases on the surfaces of static turbine blades and targets

downstream of them. The effects on performance of alternative methods

of load changing were also investigated.

The experience obtained has provided guidance for the design and opera­

tion of the Grimethorpe combustor. Figure 3. This has a cross-section

of 2 m x 2 m, and typically operated at pressures up to 12 bar, with

thermal inputs around 50 MW. This project was jointly financed by the

governments of the United States, the United Kingdom, and the Federal

Republic of Germany, and was operated under the auspices of the Inter­

national Energy Agency.

The test programme began in 1981 and continued until 1984 during which

period the combustor was operated for about 4000 hours.

Operation subsequently to 1984 has mainly been financed by the UK

Central Electricity Generating Board, and British Coal. The USDOE and

I

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1100Top gas outlet flange mating to gas distributor

Air inlet

Furnace hood

Pressure vessel top dome.

Freeboard cooler inlet

Support bracket Support steelwork (secti

Furnace support stee.work

Freeboard cooler bype

Pressure shell, 30 mn

Internal insulation

Position of manways

Probe holesbed

Propane burn-up surfaceburner rated at 150 rff

Internal access galleries

Furnace water wall

Propane burners for bed warm-up & ignition - rated at 300 kW Coal feed nozzles Distributor plate Furnace water wall inlet distributor Pressure shell bottom flangePressure vessel bottom dome1 Coal inlets

Ash removal pipes

TOOff

Figure 3. The Grimethorpe combustor.

- 65 -

have funded the provision of equipment respectively, for operation on

coal-water mixtures, and for investigating hot gas filtration, and have

contributed to the cost of testing these items.

In order to investigate the effect on combustion and sulphur retention

performance of operating over a wider pressure range than was feasible

in the foregoing rigs, a combustor with a cross-section of 300 x 300 mm

capable of operating at pressures up to 20 bar. Figure 4, was built at

Leatherhead at the end of 1981. The construction was funded jointly by

EPRI, AEP/Stal-Laval/Deutsche Babcock, Swedish NE, and the NCB-CURL.

]

I iij

Coal Itom 6olfn storage vessel

Coal lockhopper

Waterquench

VenturiCyclone vessel"

scrubberCoal teeder hopperTo pressure

let-down

filter Rotary valvesCombustor

Waterseparator Coal sampler

Slurry

Bed removal screw Bed lock hopper

bagtiller

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Figure 4. General layout of 20 atmosphere rig.

Tests totalling over 2 000 hours were carried out, including over 300

hours burning coal-water mixtures containing 70 % coal. The costs of

operating the combustor were mainly met by the USDOE.

During over 9 000 hours operation in the three combustors and their

variants, data were obtained for the following ranges of conditions:

Pressure

Bed temperature

Fluidising velocity

Excess air

Thermal input

Bed depth

Coals

5 - 20 bar

720 - 950 °C

0.7 - 2.5 m/s

25 - 100 1o0.5-59 MWt

1.2 - 4.2 m

Sub-bit > bituminous

10 - 44 %

12 - 36 %

Dolomite, limestone

ash content adb

volatiles "

Sorbents

The test programmes varied in their objectives. Some were planned to

methodically explore the effects on performance of systematically

changing process variables, whilst others were aimed at resolving spe­

cific design and operating issues raised by those engaged in design

studies for commercial plant.

2 SOME FEATURES OF THE COMBUSTORS AT LEATHERHEAD

The ensuing paragraphs are intended to highlight the major features of

the various combustors that were operated at Leatherhead, to point out

their merits, and some of the things that we would not recommend others

to do.

2.1 Combustors for pressures up to 6 bar

The containment for the various versions of the several combustors thathad plan cross-sections of 1 200 mm x 600 mm, or 900 mm x 600 mm, was avertical pressure vessel 6.7 m tall, and 1.8 m diameter. It was pro-

- 67 -

tected from risk of local overhezting by a water cooled liner; this was

the legacy from an earlier programme and was never a necessary provi­

sion. The rectangular combustor shafts were fabricated from carbon

steel and were lined with insulating refractory with a total thickness

of 75 mm which had an inner layer of higher abrasion resistance than

the outer layer. The combustion air sweeping down from the inlet at the

top of the pressure vessel provided a measure of cooling for the

casings.

In all but one version of the combustor heat absorption in the bed has

been by horizontal water cooled tube plattens. The configuration of the

tube plattens chosen results from the start up procedure adopted. The

problem was to raise the temperature of a bed that was cooled suffi­

ciently for operation at high pressure, before the pressure has risen

far enough for the appropriate quantity of fuel to be burned.

This problem was overcome on the early combustors by delaying admission

of the cooling water until the appropriate level of heat input had been

reached. The admission of water into tubes oeprating at bed temperature

was not conducive to long tube life, but in contrast to vertical plat­

tens, horizontal plattens kept their shape under these circumstances.

Tubes inserted into the bed for assessing corrosion behaviour were of

course cooled to the appropriate temperature all the time.

Fluidising air was fed into the combustor through closely spaced stand­

pipes welded to a heat resisting steel baseplate in the first build of

the combustor. Figure 2, and gas was burned in the air to preheat the

bed. This was a slow process and resulted in distortion of the distri­

butor plate. In subsequent builds, a carbon steel sandwich type distri­

butor plate was used, but the standpipes were still made from heat re­

sisting steel. For heating up the bed, propane was fed into the stand­

pipes from the compartment between the plates, initially into zones

that were being heated to the propane ignition temperature by above or

in-bed burners. This enabled the bed to be heated up much more rapidly

and in the event of there being an interruption to fuel feeding, bed

temperature could be maintained for a short while.

I

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68 -

Coal blended with the predetermined amount of S02 sorbent was conveyed

in air from a version of the Petrocarb pressurising and feeding system,

and was injected into the bed through up to four horizontal nozzles.

The rate of feed was controlled by the differential pressure between

the feed hopper and the combustor. It was NOT a satisfactory method and

was ultimately replaced by a rotary valve system.

Material elutriated from the combustor was separated from the com­

bustion gases in two stages of cyclones which were initially inside the

pressure casing of the combustor. This was done to save cost, but be­

cause of the limited space available and the vagaries of experimental

work, this proved to be a very unsatisfactory arrangement. Cyclones in

a separate pressure casing were ultimately used and the dust was

discharged dry through lock hoppers. For a commercial plant however it

would be attractive from the point of view of cost to contain the

cyclones in the same vessel as the combustor.

Turbine blade cascades installed downstream of the cyclones to assess

the erosion/corrosion/fouling characteristics of the combustion gases

included simple curved blades, Proteus marine turbine blades, and

assemblies supplied by Stal-Laval, and US General Electric.

Solids and gas sampling were carried out downstream of the cascade

after which the combustion gases were quenched in water, passed through

a venturi scrubber, and discharged to atmosphere through a pressure

let-down system. Water droplet erosion/corrosion could be a problem

unless the pH of the quench water was controlled.

T 1*

i r

F T

I E

2.2 Operating experiences

MK 1: During the initial koperation of the combustor it was found that

the internal cyclone for recycling material elutriated from the bed did

not function. The head of material in the discharge leg was insuffi­

cient to provide a seal and none of the devices tried were effective in

correcting this. The consequence was that fuel short circuited the bed

- 69 -

and there was a loss of combustion efficiency. When the cyclone had

been removed it was demonstrated that over 99 % combustion efficiency

could be achieved. Within a year of starting operation the combustor

was in a fit state to take part in a test programme carried out under

a contract between the US Environmental Protection Agency and the NCB.

Maintaining satisfactory coal flow and ash discharge however required

a large amount of effort.

MK 2: The main purpose of the test runs on this combustor was to

investigate the effect on combustor performance and turbine blade

fouling/corrosion of operating at a range of bed temperatures between

800 °C and 950 °C. The work was carried out under contract to the US

Office of Coal Research. Before the test began a new air distributor

was fitted and improvements were made to the fuel feeding and ash

discharge arrangements.

The results suggested that bed temperatures above about 900 °C could

lead to the formation of deposits on turbine blades that would be dif­

ficult to remove and pieces breaking away from one row of blades would

be sufficiently hard to cause damage to the blades in following rows.

The achievement of tolerable fuel feeding and ash removal performance

was only possible because of the diligence of the staff.

MK 3: The main objective for this combustor (Figure 5) was to investi­

gate the effect on performance of operating with a tube bank with simi­

lar height and tube spacing to a design envisaged by Foster Wheeler in

association with General Electric for a commercial plant operating at

10 bar. This was also to operate at a fluidising velocity of up to

2.5 m/s in contrast to the then maximum of 0.9 m/s. The tube bank com­

prised 27 horizontal plattens of 38 mm outside diameter heat resisting

steel tubing formed to give horizontal pitches of approximately 90 nm,

with vertical spacing between plattens (i.e., vertical pitches) of

about 75 mm resulting in a 2.0 m deep tube bank.

Because the operating pressure of the Leatherhead combustor was limited

to 6 bar, only a part of the tubes in the bank were cooled. Since the

r r

\

- 70 -

SECTION

Figure 5. Arrangement of fluid-bed combustor MK III. 1. Cooling water inlet, 2. Primary cyclone, 3. Combustor casing, 4. Start-up burner, 5. Air distributor, 6. Ash discharge, 7. Air inlet,8. Secondary cyclone, 9. Pressure casing, 10. Cascade, 11. Alkali sample point, 12. Vertical well coil inlet, 13. Water sprays, 14. Axial solids sample probe, 15. Discharge to pres­sure let-down, 16. Additional outlet for combustion gases at high flow rates, 17. Flow measurement, 18. Dust discharge from cyclones, 19. Coal feed nozzle, 20. Corrosion specimen, 21. Above bed baffle, 22. Above bed cooler - never used.

I... c ;;

W ' TJL X

71 -

investigations also included exploring performance at fluidising velo­

cities between 0.9 and 2.2 m/s, the proportion of cooled tubes in the

bed was varied over a wide range. The grouping of the cooled tubes over

the height of the combustor did not appear to have a significant effect

on performance. Owing to administrative delays, performance data for

operation at the highest fluidising velocity were not available at the

time the range of operating conditions for the Grimethorpe combustor

were chosen. The inferior performance achieved as regards sulphur

retention and combustion efficiency at 2 m/s compared with that at

0.9 m/s led GE under their contract with the ERDA to sponsor tests at

a fluidising velocity of around 1.4 m/s. The performance obtained was

considered to be sufficiently good for this fluidising velocity to be

chosen for the design study for the commercial plant, and for the

design of the tube bank used after the first series of rebuild of the

Grimethorpe combustor.

MK 4: Operation of this combustor. Figure 6, was a major deviation

from the development programme for the supercharged boiler combined

cycle. Curtiss-Wright had been successful in obtaining from the USDOE

a contract to develop the Air-Cycle. In this the bed temperature is

controlled by passing air instead of steam through tubes immersed in

the bed, and the hot air produced is mixed with the products of com­

bustion from the bed after these have been cleaned. Since in the air-

cycle tubes can operate at temperatures approaching the bed temper­

ature, change of load and start-up are simplified.

The surface area of in-bed tubing required however is greater than in

the steam cycle, the tubing is more expensive, and none of the esti­

mates of power costs made so far have shown the Air-Cycle to be

superior to conventional power plant.

The main objective of the test was to assess the behaviour of per­

formance of this type of tube bank and the performance of a bed with

vertical tubes and an unusual type of air distribution system. The

performance in all respects was better than we had anticipated, but

the tubes appeared to be very sensitive to imperfections in the

w ir 3

I - E . J

TJL

72 -

Propane.

i r

IT3 -

Figure 6. Arrangement of fluid bed combustor MK IV 1977 - 1978.

distribution of the cooling air. Subsequent work by Curtiss-Wright in

the USA culminated in a successful 1 000 hour test run, but good as the

results were, they could not outweigh the unfavourable economics of the

air-cycle approach.

MK 5: This version was the outcome of a design study carried out by

Americal Electric Power, Stal-laval (now ASEA-PFBC), and Fluidised

Combustion Contractors Limited, for a 176 MWe demonstration plant in

the USA. The design of the combustor was carried out by that group, and

they met the cost of construction and operation. There were two phases

of operation, one without tubes in the bed, and the second with the

tube arrangement shown in Figure 7.

This was the first PfBC to have a tapered bed. A tapered bed has many

potential merits. These include:

lower cost for the air distributor,

fewer coal nozzles,

fewer ash offtakes,

less fuel for initial heating of the bed.

The geometry of the combustor was wholly unsuccessful as regards com­

bustion efficiency and sulphur retention. The larger than usual pres­

sure pulsations resulted in even more unsatisfactory fuel feeding than

had prevailed previously, and this combined with the higher carbon led

to grievous problems with the very poor ash discharge system from the

primary cyclone.

The faults of MK 5 combustor basically added up to:

Taper angle too wide, bottom/top bed area ratio too large, tube

bank to distributor distance too large, tube bank tube spacings

wrong.

The lessons well and truly learned as regards the test rig were:

Feed rate control must not be solely dependent upon the pressure

differential between the feed hopper and the combustor.

- 74 -

Figure 7. Arrangement of fluid bed combustor MK V. 1. Cooling water inlets, 2. Primary cyclone, 3. Main bed cooling coils, 4. Start-up burners, 5. Air distributor, 6. Ash discharge leg, 7. Main air inlet, 8. Secondary cyclone, 9. Pressure casing, 10. Two-stage cascade, 11. Water jets, 12. Venturi scrubber, 13. Additional outlet, 14. Lower bed cooling coils, 15. Baffle tubes, 16. Above bed cooler, 17. Ash leg discharge 18. Coal nozzle.

r r

- 75 -

Ash discharge legs need to follow an essentially straight path,

and to be well above the acid dew point temperature before com­

mencing to burn a hydrogen containing fuel.

HK 6: The MK 6 combustor. Figure 8, was built to demonstrate that good

performance as regards combustion efficiency, sulphur retention, ash

removal by cyclones, and freedom from erosion, and fouling of turbine

blade cascades could be maintained over an operating period of 1 000

hours. The thermal input was 2.5 MW.

The lower part was manufactured by Babcock UK in accordance with de­

signs evolved at Leatherhead, and Combustion Systems Ltd., met the cost

of both the combustor and the cold model work involved in its design.

Babcock, under contract to EPRI, also provided assemblies of super­

heater tube metal specimens, tube hanger materials, and a variety of

other specimens for a major programme of corrosion investigations.

Stal-Laval, and GE provided cascades of turbine blades, and in addition

the latter provided a cyclone designed to investigate the possibility

of improving performance by electrostatic charging of the particles.

American Electric Power provided the coal and dolomite, and the USDOE

met the full cost of operating the rig. The layout of the test equip­

ment is shown in Figure 9.

Some ten years after operating the first version of the combustor we

had at last been able to provide ourselves with a reliable coal feeding

system, and the external cyclones provided relatively troublefree ash

removal.

The freedom from deterioration of the cascades, and generally good

performance achieved in this rig was a landmark in the development of

PFBC as it provided confidence as to the viability of the technology.

MK 7; The principle objective of the programme for this combustor was

to investigate the performance that might be achieved when using change

of bed level as the means for changing load. For this purpose we

76 -

\

Figure 8. Arrangement of fluid bed combustor MK VI 1979

2 6

Strea m Stream 2Cascade

To pressure let - down

To pressure let-down

Figure 9. 1 000 hour test facility.

Dust sampling (CURL)Ikor probes

PMS laser and impactor Speclron laserDust sampling (NYSERDA)and ionisation alkali monitorGas a/ialysis (O,. CO,, C 0 , SOj , NOx ) SO, analysisSpectral alkali photometer Electrostatic charging Electrostatic charge sampler Propane in jpc-tion (reheat )

i-j

l i r* Mi

hi

r i r**t

M

- 78 -

reverted to the use of a parallel sided combustor. Figure 10, largely

because the tapered combustor had been dismembered for rigorous exami­

nation following upon the 1 000 hour run.

The results of the tests showed that reduction of load to 50 - 60 * of

mcr by reducing bed level resulted in a lesser penalty on combustion

W_

1.

2.3.4.5.6.7.8. 9.

High temperature | Kamyr valve Conveying air •. T Purge nitrogen Fluidising distributor Adjustable tube Preheater Gas offtake Quench water Venturi scrubber

I T.

II-

Figure 10. MK Vll combustor with bed transfer facility.

- 79 -

efficiency and sulphur retention than when the same reduction was

achieved by reducing bed temperature, and that by combining the two

methods together with by-passing of air around the combustor, operation

at loads down to 30 - 40 % mcr should be feasible. The results indi­

cated that rates of load change of 3 - 5 % per minute should be

achievable.

The inferior performance achieved at low loads indicated that where

this type of operation formed a significant part of the operating life

of a plant, it would be necessary to consider recycling the fines sepa­

rated from the first stage of the hot gas cleanup system. It can be

deduced from experience on AFBC systems that this would result in a

significant improvement in performance, but as yet there are no defi­

nitive data for operation of a PFBC using recycle.

MK 8: The function of the tests on this combustor (Figure 11) was to

provide a link between the results achieved at Leatherhead and those

achieved by the much larger Grimethorpe combustor, which at that time

(April 1980) was hortly to begin hot commissioning. A new tube bank

installed in the MK VII combustor which had the same diameter tubes

(34 mm), and the same horizontal (89 mm) and vertical (65 mm) pitches

as the Grimethorpe tube bank, but the overall height of the bank was

about 1 m lower, and the distance between the bottom of the bank and

the air distributor was 450 mm as compared with 900 mm at Grimethorpe.

The combustor was operated at a fluidising velocity of 2 m/s and with

35 % excess air which resulted in both higher combustion and heat ab­

sorption intensities than hitherto experienced in the tests at Leather­

head. However, because of the lower pressure in the Leatherhead rig

(6 atm, as compared with 10 atm), the intensities were lower than was

to apply at Grimethorpe.

The results were some cause for alarm, as a bed temperature gradient of

60 °C between the top and the bottom of the bed was found, and this

would clearly be greater in the Grimethorpe combustor. Design work on a

new tube bank for Grimethorpe was therefore initiated, and because we

T

V

\

'•4

- 80 -

Figure 11. Arrangement of fluid bed combustor MK VIII 1980

- 81 -

had already shown that performance at a fluidising velocity of 2 m/s

was inferior to that at 1.5 m/s, it was designed for the lower veloci­

ty. This concluded operations for the 6 atm rig.

2.3 The 20 atm combustor

As mentioned previously this was designed to investigate performance at

pressures above those attainable in the Grimethorpe rig. Its design

incorporated a number of features that experience suggested to us were

desirable in future commercial combustor designs. The small cross-

section, 300 inn x 300 mm was dictated by cost considerations.

The layout of the equipment is shown in Figure 2, and some details are

shown in Figure 12. It was designed, built and commissioned in 9 months

and came into operation at the end of 1981. Before it was dismantled

and transferred to the British Coal Research Establishment at Stoke

Orchard in August 1984 it had been operated for over 2 000 hours.

The combustor consisted of a refractory-lined vessel, the fluidised bed

and freeboard together occupied a height of about 4.8 m. The tube bank

comprised up to 119, 25 mm diameter cooled and uncooled tubes in 34

horizontal rows, on 75 inn horizontal and 64 mm vertical pitches. Air

was admitted through a sparge-type distributor consisting of four

headers containing 16 stand-pipes arranged on a 70 mm square pitch.

Differences in thermal input were met by changing the ratio of cooled

to uncooled elements, and the pattern in which they were inserted.

Surplus bed material descends through the distributor and is cooled in

passing between cooling coils below it en-route to arotary valve which

controlled the rate of outflow to the lock hopper for depressurising.

The valve was replaced by a screw in a later edition.

For most of the tests dry coal was used. This was fed from a lock

hopper system at a rate controlled by a rotary valve into a pneumatic

3

113

t r

82

transport line that was already carrying the S02 sorbent. This had also

been fed through a lock hopper system. The coal nozzle was a plain pipe

with a dispersion baffle on the end.

Special provisions were made for in-bed measurements. Traversing probes

were installed at 1.1 m and 2.3 m above the distributor to enable the

Figure 12 20 atm rig: arrangement of distributor and tube bank.

- 83 -

solids and the gases (using conventional on-line instrumentation) to

be sampled at various positions across the bed. Fixed probes were also

installed in the bed to carry out measurements of bubble size and velo­

city (by capacitance measurements); oxygen potentials (using zirconia

electrolytic cells) and gas composition (by on-line mass spectrometry).

These special investigations were planned and carried out by staff from

Chalmers University of Technology (Sweden) and formed part of the pro­

gramme supoorted by the Swedish Department of Energy (NE). The major

part of the cost involved in manufacturing and isntalling the traver­

sing and fixed probes was met by Chalmers University.

For over 300 hours of the operating period, the coal was fed as a coal-

water mixture. This was produced on site, and mainly contained 70 - 73

* of 1 mm -0 coal by weight. In some tests the sorbent was mixed with

the coal. Both US and UK coal were used. With the US coal (caking),

combustion efficiency was maintained at a higher level when bed tem­

perature was reduced than applied with dry burning (Figure 13).

The investigation also included tests with two-stage air admission with

the intent of further reducing N0x emission. This made only a small

improvement but sulphur retention was greater improved. We do not think

however that two-stage air admission is practicable in a pressurised

bubbling bed combustion system.

The main conclusion from the programme was that operation at pressures

in the range 12 - 20 atm was unlikely to result in problems that would

discourage operation of commercial plant in this pressure range.

IT.]

f w 1

I E

3 FEATURES OF THE IEA GRIMETHORPE COMBUSTOR

W3.1 The facility ^

The decision to build the Grimethorpe Test Facility was taken at the

end of 1975. The objective was to have a flexible facility capable of

investigating a wide range of operating conditions on behalf of the

- 84 -

COMBUSTIONEFFIENCV

%

COMBUSTIONEFFICIENCY

%

HIGH VOLATILE COAL EXCESS AIR >30%

BED TEMPERATURE °C

i r.

MEDIUM VOLATILE COAL BED TEMPERATURE C. 850°C

EXCESS AIR %

CWM

DRY COAL

{{

0-16 atm. • • 10 atm. A - 16 atm. ▲ • 10 atm.

GAS RESIDENCE TIME IN BED C. 3 SECS

r t

\

Figure 13. Combustion efficiency: CWM V dry feeding.

- 85 -

sponsors. As mentioned previously owing to delays in funding some key

test programmes at Leatherhead, important information affecting the

choice of operating conditions was not available at the time the design

of the facility was frozen.

The overall flow diagram for the form of the rig used for the activi­

ties subsequently to the first test series is shown in Figure 14. The

most significant changes that were made to the test rig after the first

period included the installation of rotary valves for controlling coal

feed, major revision of the coal preparation equipment, and the instal­

lation of comprehensive research Instrumentation and data processing

equipment. The facility was and still is the largest PFBC installation

in existence. During the IEA phase the coal was burned for around 4 000

hours, and subsequently for over 2 000 hours in the course of which

coal-water mixtures with high solids content have been burned success­

fully for several hundred hours.

3.2 The combustor

The combustor which has a plan cross section within the refractory

lining of the water wall tubes of 1.9 m x 1.9 m is contained in a 4 m

diameter, 14 m high pressure vessel.

The first tube bank, 'A', was designed for heat loads up to 30 MW. The

tube bank was 3.5 m high and began about 0.9 m above the air distri­

butor. The tubes were on triangular pitches, 89 mm horizontal and 65 mm

vertical. Operated at 10 bar, with a bed temperature of 850 °C and a

fluidising velocity of 2.6 m/s, the vertical bed temperature gradient

was typically 170 °C. It was accordingly difficult to get meaningful

performance data.

The replacement tube bank 'C was designed for a fluidising velocity of

1.5 m/s, a bed height of 4.5 m to give a longer residence time, and

because the tube bank had now only to accept heat loads of up to 18 MW,

it had a much more open structure (Figure 15).

High temperature ducting

pi eaner Cascade

Combust oi

Main air compressor

Figure 14. Layout of Grimethorpe Test Facility.

2500

m rr

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- 87

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EEoom

CO

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‘A’ ‘C’

Figure 15. Comparison of layouts for tube banks 'A' and C through centre of the bank).

i view

T ' T 1-44. «. -J

The effect on the temperature gradient was all that could be wished

for. Figure 16, and combustion efficiency and sulphur retention were

both satisfactory, but then came the first major set-back in the whole

development programme - metal wastage.

This problem as far as PFBC was concerned waited until the end of 1982

to manifest itself. The subsequent investigations have involved major

research programmes in the UK and elsewhere.

Investigations involving cold models have played a major part towards

finding solutions to the problem. The most recent cold model is a full

size replica of the Grimethorpe combustor, and there is good reason to

believe that tests on this scale will provide a much more reliable

guide as to the likely success of a tube bank design than will smaller

scale activities.

- 88 -

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Figure 16. Comparison of bed temperature drop: tube banks 'A1 and 'C.

The results of the 1982 - 1983 model work was a new tube bank, ,C2I,

which had the same triangular pitch as 'C, but all the tubes were

finned or studded. This resulted in a major reduction in metal loss,

and it was observed as in other investigations the main loss was from

the evaporator tubes.

Model work and rig tests have shown fluidising velocity to be a major

factor, and it would be surprising if fluidising velocities higher than

1 m/s would be used in future PFBC's. The indication from tests of re­

cent developments in tube bank design are that commercially acceptable

tube life should be achievable.

wl

- 89 -

4 COMBUSTOR PERFORMANCE

Processing of data from the Grimethorpe combustor and from CURL rigs -w-for the investigations carried out up to the middle of 1984 have ^

resulted in a number of correlations (for Nomenclature see page xx),

some of which are given below:

4.1 Combustion efficiency

(1 - E') = K [{1 + X)_3(U/H)1*69(1300 - T)/T2] ....................... (1)

The value of the parameter K is dependent on the reactivity of the

fuel, typically 12.2 for a high volatile (c. 40 %) bituminous coal. A

plot of Equation (1) for operating temperature and bed gas residence

times relevant to a typical load range for a PFBC power plant is shown

in Figure 17.

It will be noted that this correlation does not include a pressure term

as such. Taken as a whole the experimental data do not show there to be

a significant effect of pressure per se over the anticipated load range

for a typical plant (which usually involves a pressure change of less

than 1.5:1).

4.2 Sulphur capture

Data from several combustors, including Grimethorpe up to mid-1984,

have been correlated by the equation:

R = 100 {1 - exp(-mc)]............................................................... (2)

where m is given by another empirical relationship

m = A" (tg)1/2 exp (-E/T)

i _e :i

T ' r 1

I I

(3)

- 90 -

/ //

Temperature :°C

------ Gee reeldence time in bed —4e•t •> •> n •> —2s•» •• it ii •• —te

Numbers adjacent to curves are percent excess air.

Figure 17. Effect of part-load parameters on combustion efficiency.

I E

T T

I l

The value of A" is mainly dependent on the sorbent characteristics and

variations between 10 and 50 have been found.

A polot of these equations for a reactive dolomite (Plum Run) where A"

= 50 (mm/s)^ and E = 4 650 K is shown in Figure 18. At full load,

the Ca/S ratio for 90 % retention is about 1.4, rising to 1.6 to 2.0 at

60 % load and to greater than 3 at loads below 40 %.

Over the above load range sulphur retention appears to be independent

of pressure and excess air.

- 91 -

Ca/S mol ratio

for 90% sulphur

retention

Residence time in bed, s Sorbent : Plum Run dolomite

Particle size : 1mm

Figure 18* Effect of pa^t-load parameters on sorbent requirements.

4.3 NO^ emissions

Data from the Grimethorpe plant (pre-June 1984, Figure 19) and the CURL

rigs can be correlated by the expression:

NO = 0.028 N (02 + 0.5) Ib/million Btu................................. (4)x

(0.6 lb per million Btu = 720 mg/m3).j

There is now much experimental evidence to suggest that, although

pressure has a significant effect over the range 1 to 6 bar, it has no

effect above about 6 bar. Likewise, no effect of bed temperature in the

range 750 to 950 °C has been observed.

]

T r i-

I E

- 92 -

■ TEST SERIES 2.1Q TEST SERIES 2.2O TEST SERIES 2.3A• TEST SERIES 2.3B

FRO STANDARD

U.S. EPA STANDARD

400-

O

cf ■200-

O, IN OFF-GAS, PERCENT BY VOLUME

oz

oElE -oz

Figure 19. N0x emission data from tests on the Grimethorpe rig.

Thus at full load (30 %> excess air), the N0x emissions with a 1.4 % N

coal would be about 240 mg/m3 rising to c. 400 mg/m3 at 60 % load when

the excess air level would typically have increased to 60 %. These

values are well within the U.S. emission standard of 720 mg/m3. If, or

when, the standard is made more rigorous than can be met without opera­

ting at excess air values that in current practice entail loss of com­

bustion efficiency, performance could be restored by recycling fines.

T r-

r r

T T

4.4 Heat transfer

Examination of the CURL data, and of the Grimethorpe data for operation

with plain tubes (Tube Bank C) has resulted in the following correla-

ti on:

Nu0.183 (Ar)

0.265

Ev0.7

W

\

(5)

- 93 -

where Nu is Nusselt Nubmer (based on a particle diameter),

Ar Archimedes Number, and

E Bed Voidage.

Convective heat transfer coefficient predicted by the above expression

agreed with measured values to within 95 % confidence limits of ± 14 %.

For the Grimethorpe data alone, the limits were ± 8 %, or about ± 20

W/m2 K. This compares with limits of ± 35 W/m2 K for the correlation

presented In the final Grimethorpe report. When using the correlation

it is important to calculate voidage and bed density by the procedures

that were used when the correlation was derived.

5 FUTURE DEVELOPMENTS

Two ways in which power generating efficiency of the PFBC combined

cycle system might be increased are:

1. by upgrading the steam system to operate at high supercritical

pressure with two stages of reheat. This could increase efficiency

by up to 2.5 percentage points,

2. by adding a pyrolysis/gasification stage ahead of the PFBC. The

gas produced in this stage would be burned in the stream of com­

bustion gases from the FBC raising the temperature to c. 1 100 °C.

This cycle, the After Burner Cycle (Figure 20) combined with the

advanced steam system has the potential to Increase power generating

efficiency to around 45 %.

The main uncertainty is in the design of the combustion chamber. Even

if the gases from both stages are cleaned in filters the gases will

contain some dust. Designing a combustion chamber that will remain, free

from any deposits in the mixing zones will tax the ingenuity of the

most accomplished team. Lack of success will mean disaster for the

blades of the high temperature turbine.

I l

T

- 94 -

800-850°CReheat chamber1000-

1100°C CombustionGasification

Stack

Steam j I turbine | I power plant I | and heat j i recovery ^ I____system |

Coal 8 dolomite

Wasteheatrecovery

. -A. _

r .i.

Figure 20. After-burner cycle scheme i.

6 CONCLUDING REMARKS

7 7J:'. J.

Consequent upon around 10 000 hours operation in the UK alone, there is

now a very considerable amount of data on the operation and performance

of PFBC systems, and on the associated activities of hot gas cleaning,

fuel and sorbent feeding, and ash removal.

I lIt can be concluded that:

High combustion efficiency, high sulphur retention and low N0x

emissions can be obtained when burning high-medium volatile coals

over a wide range of operating conditions.

The system is likely to have adequate load following capabilities ; *■

for most utility power plant applications.

The design bases for fuel feeding and ash removal systems have

been established.

- 95 -

Sufficient know-how has been obtained to enable tube banks having

acceptable lives as regards metal wastage to be designed with a

reasonable degree of confidence.

The results of exposing cascades of turbine blades or similar

devices to the combustion gases have been sufficiently promising

to have confidence that adequate protection of turbine blades from

particulate erosion will be provided by two stages of cyclones.

Thanks to the very considerable efforts of ASEA PfBC there is the

prospect of having several combined cycle power plants in operation

within the next three years.

Major contributions to the development of the technology have been made

at Grimethorpe, but despite all the foregoing endeavours and the in­

vestment of many tens of millions of pounds, and other currencies, no

one has yet run an industrial gas turbine on the combustion gases from

a PFBC. This factor alone dictates the need for major back-up activi­

ties, e.g. on hot-gas cleaning and on more detailed examination of the

effects on the dust loading and constituents of the combustion gases on

turbine blades.

There is a powerful justification even now for operating a gas turbine.

The unique facilities and the proven and highly accomplished team at

Grimethorpe make this a logical choice of venue for a major back-up

programme.

There is also need:

to develop an understanding of the mechanisms involved in tube

metal wastage, and experience has also shown the need to have

available large ‘cold1 rigs in which models of tube banks can be

tested,

to establish reliable laboratory procedures for predicting the

sintering behaviour of material fed to fluidised beds.

r _T: j

T T 1-ii.__j

I l

w

- 96 -

to ascertain the reasons for there still being a considerable

scatter in the data on sulphur retention,

to provide the background for reducing even further N0x emission,

e.g. by operating at much lower excess air levels, and by using

recycle to restore high combustion efficiency,

to develope reliable and safely applicable 'on-line' techniques

for detecting significant changes in dust loading and size con­

sistency of the gas stream ahead of the gas turbine. It is essen­

tial to have means available to detect malfunction of hot-gas

clean-up equipment before serious turbine operation begins.

Many of the foregoing activities involve operation of large scale plant

but there is still considerable scope for work on small units. In addi­

tion to contributing to the development of the technology and measuring

techniques, these have a vital function in providing basic training for

PfBC engineers for the considerable number of plants that we shall hope

to see in the future.

r' r

f

ACKNOWLEGEMENTS

Firstly, I would like to thank the organisers for Inviting me to this

meeting.

I would also like to acknowledge the major contribution that Scandi­

navia has made to the promotion and development of the technology. The

former began early in 1970 when Henrik Harboe of Stal-Laval having seen

the activities at Leatherhead set about promoting interest in PFBC

technology at every possibly opportunity.

Thanks are also due to all those who have seen the wisdom in contri­

buting to the development of PFBC technology, particularly in the UK.

These include U.S. Department of Energy, EPRI, Stal-Laval, Swedish NE,

American Electric Power, General Electric, Combustion Systems Ltd.,

Wf

- 97 -

the Governments of U.K., U.S.A., and F.R.G. in their support of

Grimethorpe, Combustion Systems Ltd. and Babcokc U.K. and U.S.A.

In paying tribute to those who have striven to get the technology to

its present state of development, often against non-technical rather

than technical problems, I would also wish to remember those who are

sadly no longer available to see its success, in particular, Douglas

Elliott and George Sparham.

f J

NOMENCLATURE i :i

A" "Sorbent reactivity index" in Equation (3), (irm/s)1/2 A Total submerged heat transfer areaAr Archimedes Number = g Zpr (s-r)/u2C Ca/S molar ratioCb Specific heat of bed materialeg Specific heat of gasD Particle size (e.g. surface mean dia.), mm in equ (3),and m in equ (5)E' Fractional combustion efficiency in Equation (1)E Constant related to the apparent activation energy for

the sulphur capture reactions in Equation (3), K Ev Fractional bed voidage (typically 0.7) in Equation (5) H Bed height, mh Convective heat transfer coefficient, W/m^Kho Overall heat transfer coefficient from bed to water or steamK Constant; representing fuel reactivity,in Equation (1)k Mean gas conductivity in film, W/raK% Bed weightMg Gas mass flow ratem A "sulphur capture parameter", dimensionlessN Nitrogen content of coal, wt. % d.a.f.Nu Nusselt Number & hD/k02 Oxygen content of the dry gas, mol % p' Partial pressure of oxygen, barR Sulphur capture, %r Mean gas density in film, kg/m^ «s Mean particle density (usually taken to be 2500 kg/m^T Absolute bed temperature, Kt Gas residence time in off-gas system, stg Gas residence time in bed, sU Fluidising velocity, m/su Mean viscosity in film, kg/msX Fractional excess air level.

T f l

-u..------ J

I l

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SOME REFERENCES

Leatherhead Combustors ^ TFTMk 1/Mk 2BCURA Laboratories, Fluidised Combustion of Coal at High Pressures, AIChE Symposium Series, Vol. 68, No. 126, p.225.Mk 3NRDC Contract No. 14-32-001-1511. "Pressurised Fluidised Bed Cambustion”, R & D Report No. 85, Interim No. 1, 1974.Mk 6Coal Utilisation Research Laboratory, "investigations on the Leatherhead Pressurised Facility". Proceedings of the 6th International Conference on Fluidised Bed Combustion, April 1980, Atlanta, Georgia.Mk 7Coal Utilisation Research Laboratory. "Final report on tEst 3, 1980, in the Leatherhead fluidised bed facility". Report No. IE-14129-5 to the U.S. Department of Energy, March 1981.Mk 8NCB (IEA Grimethorpe) Limited. "Report on the Leatherhead Link Tests." Report No. GEF/U/81/4, December 1981. Also published by NTIS as Report No. DOE/MEPC-83-63.20 Atmosphere RigGoal Utilisation Research Laboratory. "Performance of a snail combustor at pressures up to 20 atm". Paper to the 7th International Conference on Fluidized Bed Cambustion, Philadelphia, October 1982.Grimethorpe Rig"Grimethorpe Pressurised Fluidised Bed Combustion Project: Overall Project Review". IEA Grimethorpe GEF/U/84/21, Mayl985.

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PRESENT STATUS AND FURTHER DEVELOPMENT OF PFBC COMBINED-CYCLE PLANT

Krishna K. Pillai Manager Market Development ASEA PFBC ABS-612 20 Finspong, Sweden

Abstract

?FBC plants are under construction in Sweden, Spain and the US. This paper summarises these projects and describes the technology they are based upon. The PFBC concept is described together with the advances which have brought this technology into the start of its commercial phase.

The role of the Component Test Facility (CTF) in the development of pressurised fluidised bed combustion (PFBC) technology is described together with an update of the further development in progress. An Important feature of PFBC technology is not only that current "Clean Coal" requirements can be met effectively but also that clear development paths can be identified for the future. The 10-15Z saving on fuel consumption available today can be reduced by a further 15-25Z over the next 25 years. The potential available for the Improvement of emissions performance means that current regulations can be met and continue to be met as more and more stringent requirements are placed on power generation systems.

Whereas efficiency and environmental requirements determine the entry of a new technology into the market place and also provide the driving force for development efforts, plant availability and investment costs are the criteria which will determine its acceptance.

In retrofit applications, the combined cycle provided by the PFBC leads to significant efficiency improvements together with an

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uprating of power production. This Is combined with the ability to use high-sulphur or low-grade coals In an environmentally acceptable manner. The low footprint area required for this pressurised process also makes constrained sites feasible.

1 STATUS OF CURRENT PFBC PROJECTS

1.1 General

PFBC plants are under construction in the US, Spain and Sweden. These plants will be taken Into operation through 1990. Tables 1, 2 and 3 provide a summary of these projects.

Table 1. PFBC Projects - Summary.

SITE ESCATRON TIDD VARTAN

Utility Endesa OPCo SEPSupplier ASEA PFBC + BWE ASEA Babcock PFBC ASEA PFBCPurpose PFBC Demo PFBC Demo CogenerationOutput 79 MWe 70 MWe 130 MWe +

210 MWtUnit lxP200 lxP200 2xP200Steam turbine Existing Existing New turbineStart-up date Mid 1990 Mid 1990 51 Late -89

52 Early ’90

1.2 Sweden

Stockholm Energi, the main producer of district heat for the city of Stockholm, with an annual heat-production of about 4,000 GWh, has decided to injtall a PFBC cogeneration plant at VMrtan in the centre of Stockholm.

Like most cities in Sweden, Stockholm is, to a great extent, served by a district heating system. This system is rapidly expanding and is targetted to supply most of the city's heat demand in the 1990s.

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The largest heating plant is VUrtan with a capacity for supplying nearly 1,000 MW heat, based on oil fired boilers and heat pumps.

Table 2. PFBC Projects - Technical Data.

PROJECT Escatron Tidd VSrtan

Coal type Lignite Ohio PolishCoal-S 7Z 4% 1%Ash 36Z 10% 15%Coal feed Dry Paste PasteSorbent feed Dry Dry Mixed as pasteSorbent Limestone Dolomite Limestone, DolomiteFeed points 16 6 6CyclonesSteam data:

9x2 7x2 7x2

Pressure 95 bar 90 bar 140 barTemperature 510°C 496°C 540°C

Excess air 15% 25% 25%Fluidising velocity 0.9 m/s 0.9 m/s 0.9 m/sBed height 3.5 m 3.5 m 3.5 mPressure 12 bar 12 bar 12 barCoal feed rate 18.0 kg/s 7.2 kg/s 2x8.0 kg/sSorbent feed rate 7.0 kg/s 2.3 kg/s 2x0.75 kg/sSteam flow 60.0 kg/s 56.0 kg/s 114.0 kg/s

Table 3. PFBC Projects - Emissions Data.

PROJECT Escatron Tidd Vartan

Sulphurremoval 90% 90% c.85%Emission <50 mg/MJ

«0.12 Ib/MBtu)Ca/S ratio NOx

1.8 1.6 c.1.3

emissions 150 mg/MJ (250 mg/MJ) 70 mg/MJ(0.35 Ib/MBtu) 0.5 Ib/MBtu (0.15 Ib/MBtu)

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In Sweden there is an extensive program to reduce oil dependence and to utilize all possible heat demands for the combined production of heat and electricity. Thus the introduction of a coal-fired cogeneration unit at Vartan has been considered for a long time. However, for environmental reasons, a decision has been difficult to reach.

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The decision was finally taken to use PFBC technology for Vartan based on Its economic and environmental advantages compared to other alternatives.

The authorities Impose very stringent emission standards on this city plant. It is also subject to special restrictions concerning dust and noise as there are residential areas directly adjacent to the plant.

The PFBC concept lends itself very well to this site where a minimum of space is available. Compactness is a feature of the pressurized combustion system.

The change-over to coal-fired PFBC boilers having low specific emissions of SC^ and NO^ gives Stockholm Energi a higher production capacity with no additional environmental impact.

The new cogeneration installation consists of two PFBC modules of 200 MW thermal input each (P200)t and is designed for base load operation.

ASEA PFBC has turnkey responsibility for the VMrtan plant. Site construction is in progress and most major equipment has beer, ordered. The first unit will be started up in late 1989.

1.3 USA

American Electric Power (AEP) has decided to install one PFBC module of 200 MW thermal input (P200) for the repowering of its Tidd Power Plant at Brilliant, Ohio, as a demonstration plant for PFBC combined-cycle technology.

AEP is one of the largest electric utilities and coal consumers in the United States. They consume over 40 millions tons of coal annually to produce some 85% of their electricity generation.

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AEP's commitment to clean coal technologies which generate electricity in an economic and environmentally compatible manner has led to Interest in the development of PFBC combined-cycle technology, especially for the use of vast high-sulphur coal reserves In the states of Ohio and Indiana.

AEP and ASEA PFBC (then STAL-LAVAL) began a study of PFBC in late 1976. The encouraging results of this study led to further cooperation between the two companies. Combustion tests at Leatherhead, England and further test programs at ASEA PFBC's Component Test Facility (CTF) in Sweden confirmed the technical and economic advantages of PFBC technology over conventional pulverized coal plants with flue gas desulphurization systems and even over other coal technologies.

The ultimate goal of the Tidd PFBC demonstration plant is to establish this technology as a viable alternative for generating electricity using coal.

Both the U.S. Department of Energy and the State of Ohio Coal Development Office have selected this program for funding in their respective Clean Coal Technology Programs.

In the United States, ASEA PFBC cooperates with the Babcock & Wilcox Company. In late 1985, B&W and ASEA PFBC formed a business partner­ship, ASEA Babcock PFBC, to commercialize PFBC technology in the United States and Canada. This company is the main contractor for Tidd P200.

Site work is due to start in April 1988. The pressure vessel for the PFBC is now virtually complete and the boiler is being fabricated. All major equipment has been or is in the process of being ordered. Start-up is scheduled for summer 1990.

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1.4 Spain

Empresa Naclonal de Electrlcidad S.A. (ENDESA) is the Spanish stateowned utility producing about 25% of Spain's electricity mainly fromcoal and lignite. The black lignite reserves in the Aragon provincetotal more than 1000 M tonnes with known recoverable reserves ofmore than 400 M tonnes. These coals contain 4-8% sulphur, 25-45% ashand about 20% moisture. The annual consumption was 5 M tonnes in1985 and it is Intended to be increased to 10 M tonnes per annum

*corresponding to about 2000 MWe.

After a thorough investigation of different technologies, ENDESA has chosen the PFBC system of ASEA PFBC design as the most promising technology for burning this type of coal especially considering its potential for future larger plants of 350 MWe size.

ENDESA has, therefore, decided to install one PFBC module of 200 MW thermal input (P200) for repowering the existing unit No. 4 in the Escatron Power Plant as a commercial demonstration plant for Spanish black lignite with high ash and high sulphur content. This is the first step in commercializing PFBC technology in Spain.

The PFBC unit will be delivered in cooperation between ASEA PFBC AB and Babcock & Wilcox Espanola S.A. About 60% will be manufactured in Spain.

Site work is due to start in February 1988 and procurement of equipment has started. This unit will start-up in spring 1990.

1.5 The Gas Turbine GT 35P

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All the four units ordered so far will utilise the 17 MW^ GT 35 gas Wturbine specially modified for PFBC use. The first GT 35P machine has now been constructed and is being installed on a specially designed full-scale, oil-fired test stand in ASEA STAL's workshops

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in Flnspong, Sweden. The comprehensive test program for this machine will begin in December 1987. All GT 35P machines will be factorytested upto full load before being transported to site. w ^

2 THE CURRENT PFBC CONCEPT

2.1 Concept

Today, PFBC represents a coal-fired combined cycle process which incorporates a pressurised bubbling fluidised bed and a gas turbine operating at a conservative turbine entry temperature. The deep (c. 4 m) fluidised bed together with the relatively low fluidising velocity (1 m/s) gives excellent sulphur capture and combustion efficiency data without having to resort to recycle arrangements. Recycling of fines is only called for when dealing with very unreactive fuels (anthracite or petroleum coke for example) or when particularly low values of minimum load are required (c. 20% mcr).

The basic cycle is illustrated in figure 1. Some design features worth noting are:o Separation of temperature and pressure barriers o A two-shaft gas turbine with a small intercooler between the

LP and HP compressors to limit pressure vessel temperature o Mild steel as material for the pressure vessel o All hot parts surrounded by higher pressure, compressed airo Absence of moving parts within the pressure vesselo Absence of refractory linings through the gas patho Hot gas cleaning (cyclones) and load changing system within

the same pressure vessel as the fluidised bed o A ruggedised. Industrial gas turbineo Bed temperature below 880°Co Gas turbine entry temperature of 830°C

o Integral gas-air ducting and isolation valves between the combustor and the gas turbine

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PRESSURIZED FLUIDISED BED BOILER

Figure 1. Pressurized fluidized bed boiler, Asea PFBC.

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The combustor layout Is Illustrated In figure 2 and shows some of these features.

2.2 Coal Feeding

Two methods of coal feeding are available. For coals with ash contents of less than about 25Z which have a relatively high heating value, pumping of a coal-water-paste is used. The paste has a total water content of between 20 and 25Z. At this level of water content, net efficiency is virtually unchanged compared to dry coal feeding since:(a) part of the evaporative loss is recovered as additional mass

flow through the gas turbine,(b) coal drying is eliminated,(c) lock hopper pressurisatlon and inertisation is eliminated.

For coals with high ash contents and/or low heating values, the efficiency penalty became unacceptable when feeding as a paste and pneumatic feeding via a lock hopper is used instead.

If a low-sulphur coal is used (< 1.5% S) then the sorbent material material (limestone or dolomite) can be mixed directly into the coal-water-paste. For higher-sulphur coals, pneumatic feeding is used. (See figure 3)

2.3 Steam Generation

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The gas turbine exhaust gas is used for feed water heating but represents only about 10-15Z of the heat transferred to the steam cycle. The remainder, including all steam generation, superheating and reheating (if applicable), comes from the in-bed tube bundle. Intensive development of tube bundle design has taken place in three areas:(a) Design of bed and tube bundle geometry together with

fluidisation studies to minimise the erosion potential within the fluidised bed.

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Combustor C 25 for PFBC 200

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Figure 2. The layout of Combustor C 25 for PFBC 200.

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Figure 3. Selection of coal feeding method.

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(b) Testing and choice of tube materials, and(c) Tube bundle design and construction to enable rapid service or

tube-changes to be carried out.

This development has led to design values of tube bundle life of over 50,000 hours for all tubes. In addition the design philosophy allows rapid repair or. If necessary, rapid replacement of entire tube sections. Materials under Investigation promise over 100,000 hrs life.

The once-through boiler concept is normally used though a drumboller could also be used.

At the steam conditions currently being used (160 bar, 565°C/5650C)

corrosion does not limit tube bank design. Present materials are suitable for perhaps upto 600°C. However at these higher

temperatures and consequently high steam pressures, the high heat fluxes in a fluid bed lead to material limitations from the view­point of thermal stresses.

2.4 Gas Cleaning

We judge that hot gas cleaning Is not yet commercially mature and thus, for the present concept, we use a two-step gas cleaning process. The first step - with two stages of high efficiency cyclones - performs the gas turbine protection function. For clean air requirements however, a back-end filter at conventional pressures and temperatures is used.

2.5 Gas Turbine

The available gas cleaning techniques today together with the ruggedisatlon of our gas turbines for PFBC use give a design life for the first stage rotor blades in excess of 20,000 hours. Here too the design problem has been attacked from two converging directions.

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Much effort has been expended in obtaining a truly modularised design such that rapid replacement of modules is possible, with the module removed being then refurbished to provide the new spare.

2.6 Load Changing

Bed level alteration at constant bed temperature is used for load changing. Bed storage vessels are placed within the pressure vessel and bed material is transported pneumatically between the storage vessel and the fluidised bed. When bed material is removed, in-bed tubes are exposed in the freeboard. This leads to both steam generation and the gas turbine entry temperature being reduced. The two-shaft gas turbine with its free-wheeling low pressure shaft gives a constant volume flow over much of the load range. This in turn leads to fluidising velocity, cyclone velocity and excess air remaining constant over much of the load range.

Load changing rates are around 4% per minute and minimum load is typically 30% mcr. A cold-start takes 5-6 hours while a warm start can be achieved in around 2 hours.

3 CURRENT PFBC PLANT SIZES AND THEIR PERFORMANCE

PFBC module sizes are determined by the chosen gas turbine which, in turn determine the air flow delivered. Table 4 indicates the module sizes available based on the gas turbines adapted for PFBC application.

Table 4. PFBC Modules.

DESIGNATION P200 P800

Gas Turbine GT 35P GT 140PNet Power MWe 80 350Gross Power MWe 84 366GT Power MWe 17 80ST Power MWe 67 286Net Efficiency % 42.5 44Heat input MWt (LHV) 188 795

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One or more PFBC modules can be coupled to a single steam turbine.As shown in Table 5 the current configurations span the range from 80 MWe upto 700 MWe. Table 4 also gives the plant efficiencies. It should be noted that the larger module exhibits a higher efficiency due to the absence of the gas turbine gear, the slightly higher pressure ratio (16:1 instead of 12:1) and the economies of scale for auxiliary power. Both module sizes exhibit a small efficiency increase when coupled in pairs to a larger steam turbine.

Table 5. Current PFBC Plant Sizes and Efficiency.

Plant Size No. of PFBC Steam Gas Net(net) MWe modules turbine turbines efficiency

MWe MWe % LHV

80 lxP200 67 1x17 42.5162 2xP200 135 2x17 42.9243 3xP200 202 3x17 43.0

350 IxPSOO 286 1x80 43.8705 2xP800 573 2x80 44.1

Currently coals ranging from lignite to low-volatile bituminous can be accepted. Anthracite is acceptable but needs some fines recycle if high combustion efficiency is to be maintained (> 9SZ). Excess air levels are at around 20% (15% for lignite). Carbon monoxide levels are usually negligible (< 20 ppm).

Sulphur capture levels of upto 98% can be achieved if sufficient sorbent can be fed. With limestones of "average" reactivity a Ca/S ratio of around 2.0 is needed for 90% retention. With dolomites the corresponding Ca/S ratio is about 1.5. Sulphur levels of upto 8% can be accepted provided that feed systems and ash removal systems can be designed to cope.

Ash levels of 40-50% (depending upon how much sorbent must be fed too) are acceptable.

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NO^ emissions are typically around 100 mg/MJ at these excess air levels, and, with special reduction procedures (but not catalytic converters) can be reduced to around 60 mg/MJ. Note that with PFBC, unlike atmospheric fluidised beds, NO^, S0x and CO are very weakly linked. Thus 60 mg/MJ NO can be achieved simultaneously with around 50 mg/MJ SO^ (or lower) and less than 10 mg/MJ of CO.

The ash produced contains a mixture of partially sulphated sorbent and coal ash. Virtually no free lime (CaO) exists and sulphites or sulphides are also negligible. After setting with water the heavy metals in the coal ash are non-leachable but of course sulphate content is high. Strengths are comparable to concrete and may even exceed those of concrete if small quantities of a binder (cement) are added. One disadvantage when using dolomite as sorbent is that the slow hydration of MgO causes the strength to deteriorate. Limestone based ash is much more stable. Development efforts have concentrated on, and succeeded in, testing and identifying potential uses of this material. The material has been found suitable:(a) as road building material(b) for manufacture of synthetic gravel(c) for manufacture of "formed" bricks or tiles(d) as general building material, and(e) as roofing or flooring material for tips containing hazardous

material.

An entrepreneurial phase now remains to find sales outlets for this material. It thus seems highly probable that ash disposal costs - which are currently being used in the economic evaluation of PFBC plants - will give way first to a zero cost disposal (free to user who will transport material away from the site) scenario and then to a credit for the sale of such material.

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4 HIGH FERFOBMANCE AND SECOND GENERATION PFBC'S

4.1 High Performance PFBC's

With no change to the current concept and with much the same gas turbine entry temperature as present (830-870°C), significant

improvements can be made to PFBC plant performance. Plants ordered over the next 3 or 4 years will exhibit these improvements to an increasing degree.

Since superheat and reheat surface requirements in a PFBC are only between 10 and 20Z of that required in a conventional pulverised boiler, the techno-economic optimal steam conditions for a PFBC are different to those of a boiler relying mainly on convective heat transfer surface. In general, the optimum lies at higher temperatures and lower pressures than for a conventional boiler. The steam pressure value is partly limited by thermal stresses arising from the high but evenly distributed heat fluxes in the PFBC. This in turn leads to optimum reheat temperatures somewhat higher than main steam temperatures.

The materials available today are suitable for upto 565°C from the

erosion, corrosion and strength viewpoints. Materials already developed (using powder metallurgy) and currently being tested will soon allow steam temperatures to be raised to around 600°C. The

efficiency of the current PFBC concept can be Increased by about 1 percentage point by this means alone.

For simplicity the gas turbine exhaust gases are used only for feed water heating. Some efficiency gains are possible with the added complexity of a heat recovery steam generator. However this is unlikely to be economic till higher gas turbine exit temperatures are attained.

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Emissions performance is also improving. The effectiveness of the NO^ reduction techniques currently being used Involve:(a) modifying combustion patterns within the bed> and(b) the use of in-bed or gas path additive injection (though

without any catalyst)and are not, as yet, well understood. Nevertheless, NO^ emissions currently can be held to around 60 mg/MJ (independantly of SO^ or CO emissions) and results from our test plant show that much lower emissions are possible. The next PFBC ordered (l.e. within 2 years) will be designed for NO^ emissions of less than 40-50 mg/MJ.

While S0^ emissions themselves pose no significant limitations, there is a strong incentive to reduce sorbent consumption to a minimum. The use of limestone rather than dolomites in PFBC is relatively recent, but it is clear that the original fears that high pressure operation would inhibit calcination and consequently sulphur capture (based on theoretical analyses in the US) are mlsfounded. Within 3-5 years therefore the High Performance PFBC (HP-PFBC) will need a Ca/S mol ratio of around 1.5-1.6 with limestone for 90Z sulphur capture. Optimisation of size distribution and recycle of streams containing unused calcium are the main lines of development here.

For dust control it is anticipated that the HP-PFBC will still use cyclones for turbine protection and a bag filter for clean air requirements. The expected tolerable dust emission levels of 5-10 mg/MJ will probably exclude the use of electrostatic precipitators for final clean-up, especially if regulations regarding sub-micron particles begin to come into force.

A summary of the development potential for PFBC is given in Table 6.

4.2 Second Generation PFBC's

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Continued development of PFBC then requires that some means be found for increasing gas turbine entry temperature. Since the bed temperature is limited to be under 900°C this implies the use of an

Table 6 Development Potential with PFBC Combined Cycle Plant

Possible Order Date

Plant Description Steam Data Gas Turbine Data

ExpectedEmissions

Efficiency Net (LHV) !

Today "Conventional" PFBC 180 bar 5650C/565°C830°C TBT Cyclone Gas Cleaning

NOx: 60 mg/MJSOx: 90Z at Ca/S-2 Dust: 5 mg/MJCO: 10 mg/MJ

44.1

1989-92 "Conventional" PFBC 190 bar565°C/600°C850°C TBT Cyclone Gaa Cleaning -

HOx: 50SOx: 90Z at 1.8 Duat^ 5CO: 10

44.6

1990-93 "Conventional" PFBC High Performance Version

180 bar 600°C/600OC870°C TBT Cyclone Gaa Cleaning

NOx: 40SOx: 90Z at 1.6 Dust: 5CO: 10

45.0

1992-96 Hybrid PFBC + Pyrolyser f8r temp.topping

c.210 bar 600oC/625°C950°C TBTHot gaa cleaning

NOx: 30SOx: 90Z at 1.3 Dust: 5CO: 10

46.5

1995-98 Hybrid PFBC + Pyrolyser for temp.topping

c.220 bar 625oC/650°C1050°C TBTHot gaa cleaning

NOx: 25SOx: 90Z at 1.3 Dust: 5CO: 10

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1998-2000 Hybrid PFBC + Pyrolyser for temp.topping

c.240 bar 650oC/650°C1150°C TBT Advanced hot gaa cleaning

NOx: 20SOx: 90Z at 1.2 Dust: 5CO: 10

48.8

2010 Hybrid PFBC + air blown gasifier for temp.topping

c.3Q0 bar650 C/700°C1500°C TBT Very advanced hot gas cleaning

NOx:<20SOx: 90Z«><1.2 Dust: 5CO: 5

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after-burning process. Here well-known processes for pyrolysis or partial coal gasification could come into play, to produce the gas for turbine temperature topping. The PFBC must be retained for combustion of fresh coal and the residues from the pyrolysis or partial gasification process. These hybrid schemes have the big advantage that the PFBC Is available not only for steam generation but as the sink for partially utilised coal and for heat losses from the partial gasification stage.

Pyrolysis, depending upon coal type, permits gas turbine entry temperatures to be increased to between 1100 and 1200°C. To raise

temperature further would need the use of steam, though an air-blown process is sufficient, in a partial gasification stage. Here, of course, fuel gas and PFBC exhaust gas would still be at temperatures below 900°C but would need rigorous cleaning before the temperature

topping process.

Hot gas cleaning is therefore necessary for this development and must be suitable for both oxidising as well as reducing atmospheres. The level of cleaning required is intimately connected to the final gas turbine entry temperature and a staged development as shown In Table 6 is possible.

Gas turbine development beyond today's state-of-the-art will be required once temperatures exceed about 1200°C. It should be noted

however, that the hybrid-PFBC always needs a turbine temperature some 300-400°C lower than a pure gasification process for power

generation to attain the same efficiency. (See figure 40

Emissions performance will be largely determined by the PFBC and the after-burning process for NO^ and by the pyrolyser/PFBC for SO^.

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50

45

40

PFBC 180/565/565

PC 240/565/565 PC 180/565/565

35

30 ■— 600

J_ _ _ I__ L700 800 900

1

Gas turbine Inlet tempera-

J__________ ture, °C

1000 1100 1200

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Hgure 4. Efficiency comparison for power generation systems.

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5 CONCLUSIONS

Without good plant availability, performance can be worthless. A basic requirement while working on performance development is therefore that engineering development proceeds in parallel to ensure and then improve plant availability. The techniques developed during the Swedish nuclear program are being used to identify critical plant systems and components and these are the focus of efforts to simplify and Improve.

From the performance viewpoint, PFBC, as a newly commercial technology, offers not only the boost needed today but also clear lines of development for the future. Fuel consumption savings of 10-152 are available now and a further 15-252 is available in the future. The emissions potential means that a "Clean Coal Technology" today maintains its ability to meet the more and more stringent requirements anticipated over the next 20-30 years.

BIBLIOGRAPHY

1. Pillai K. K., WickstrSm B., and Tjellander G., The Influence of Gas Turbine Technology on the P200 Power Plant, in Proceedings 2nd Biennial PFBC Power Plants Untility Conference, Milwaukee 1986.

2. Menendez-Peres E. et al, Escatron - A PFBC Combined Cycle Demonstration Plant for High Ash and Sulphur Coals, in Proceedings 2nd Biennial PFBC Power Plants Utility Conference, Milwaukee 1986.

3. Almqvist P., The PFBC Plant in Vartan, Stockholm, in Proceedings 9th International Conference on FBC, Boston 1987.

4. Mudd M J and Bauer D A, Tidd PFBC Demonstration Plant, in Proceedings 9th International Conference on FBC, Boston 1987.

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PRESSURIZED FLUIDIZED BED COMBUSTION OF PEAT

Andras Horvath, M.Sc.(Eng)

Helsinki University of Technology

Laboratory of Energy Economics and Power Plant Engineering

02150 Espoo, Finland

Abstract

Combustion and emission properties of peat (a fuel of high volatile and low ash content)

were investigated under the special conditions of pressurized fluidized bed combustion.

At the Otaniemi PFBC/G test rig, during a total of 150 hours of test runs crushed peat

pellets of about 20 % moisture were burnt. Natural silica sand was used as the bed

material. The bed height was 0.6... 1.2 m depending on operating pressure and thermal

input. The pressure levels of operation were 5,7 and 10 bar (abs.). The amount of excess

air was varied between 10 and 50 %. The flue gas composition was monitored

continuously after the two stage cyclone collector, using pressure reduction and double

sampling system. Gas samples were also taken from the bed, freeboard and downstream

of the cyclones. Bed material and fly-ash samples were analysed in respect of fuel bum out.

Combustion efficiency was over 99 %. The losses caused by unbumt carbon in fly-ash

were very low. Carbon monoxide emission was always below 90 mg/MJ, effected by the

bed temperature. The emitted amount of sulphur dioxide was smaller than 160 mg/MJ

(peat has. a low sulphur content of 0.2 %wt, d.b.). Sulphur trioxide was found in the flue

gases and due to that rigorous low temperature corrosion appeared. NOx-emission was in

the range of 10 to 100 mg/MJ strongly influenced by excess air. The effect of pressure on

NO-formation appeared to be moderate. Compared to atmospheric fluidized bed burning

of peat, slightly better combustion and emission performance were experienced in the

pressurized fluidized bed combustor.

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1 INTRODUCTION

Combined cycle power generation, using pressurized fluidized bed combustion (PFBC),

has the potential of increasing power generation efficiency and at the same time fulfilling

environmental requirements, especially on S02 and NOx emissions. There is quite a lot of

informadon available on the pressurized fluidized bed burning of coal, as a result of

intensive research activity all over the world. In Finland peat is the only indigenous fuel,

thus it is burnt widely in conventional PF, AFB and CFB boilers. Within the joint project

of Helsinki University of Technology, Technical Research Center of Finland and Imatran

Voima Oy, a PFBC/G test rig was built in Otaniemi, Espoo. Peat is the fuel of main

interest during the first phase of the research project. The test rig has been in operation

for 150 hours in the autumn of 1987. The investigation focused on the combustion and

emission properties of peat corresponding to pressurized fluidized bed burning

conditions.

2 THE FACILITY AND MEASURING SYSTEMS

The Otaniemi PFBC/G test rig has a 1.2 m high bed area of 0.15 m diameter and a 2.5 m

high freeboard area of 0.25 m diameter. The electrically preheated air flows through the

nozzle type air distributor in to the water cooled (vertical cooling coils) fluidized bed. The

flue gases leaving the freeboard pass through a two stage cyclone collector and heat

exchanger to the pressure let down valve (Figure 1). Fuel is fed into the bed. The system

pressure can be increased up to 10 bar, the thermal input can reach the 130 kW level.

Figure 1. The flow diagram of the PFBC/G test rig.

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The measuring system includes gas and solid material sampling devices, different types

of gas analysers, and chemical analysis. The flue gas composition is monitored

continuously after the cyclones, using pressure reduction, filtration and double sampling

system. 02 concentration is measured by Zr02-analyser, while C02, CO and S02

concentrations are analysed by infrared techniques. NOx emisssion is measured by

chemiluminescence gas analyser. CxHy, CH4, CO and NH3 are determined by a single

beam variable path IR-analyser. Gas samples are taken from the bed and freeboard (0.2 m

and 1.8 m above the air distributor), using air cooled probes and from the sample NOx,

CxHy, CH4, CO and NH3 concentrations are determined. Figure 2 shows the position of

the sampling ports and Figure 3 the arrangement of gas analysers.

3

I _ I ‘ 1

PAHTCLE SAM PLUG

’ PRESSURE t LETDOWN

l E

Figure 2. Sampling ports at the PFBC/G test rig.

Before and after the cyclones S03 (H2S04) concentration was measured using

isopropanol, barium chloranilate crystal bed and light absorption method. Bed material is

sampled 0.3 m above the air distributor through a water cooled probe, into nitrogen

atmosphere. The samples are analysed for ash, volatile and nitrogen content. The ash

content and oxide composition of the cyclone fly-ash was also analysed.

124 -

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■I 0

StaMAOOO NOxJtO

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LET DOWN

BEOSXMniia CEMiCFLIBI

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Figure 3. Arrangement of gas analysers at the PFBC/G test rig.

3 PROPERTIES OF PEAT AND BED MATERIAL

Peat is a fuel of high volatile (about 69 %, d.b.) and low ash content (about 4 %, d.b.). It

has a high moisture ccmtent when harvested. During the experiments at the PFBC test rig

crushed peat pellets of about 20 % moisture were burnt The mass mean particle size of

die fuel is 4.83 mm (33 % larger than 6 mm and 10 % smaller than 1 mm). The proximate

and ultimate analyses of peat arc presented in Table 1. Regarding emissions it can be

noted, that the sulphur content of peat is lower (about 0.2 %wt, d.b.) and the nitrogen

content is somewhat higher (1.4 %wt, d.b.) than in coal.

Natural silica sand was used as bed material. The mean particle size of the bed material

was 0.73 mm before the test and it decreased to 0.33 mm (32 % smaller than 0.3 mm)

during the tests, because of attrition and the effect of fine ash particles.

- 125 -

. Table 1. Properties of peat.

proximate analysis (dry base)%

ash composition %

- volatile matter 69.1 Si02 50.3- ash 4.2 AI2O3 15.3-fixed carbon 26.7 Fe^ 12.0

CaO 7.9ultimate analysis (dry base) MgO 2.3- carbon 55.1 K20 1.8

- hydrogen 6.2 NajO 0.98-nitrogen 1.4 TiOz 0.27- sulphur 0.19 P2O5 2.4- oxygen (by difference) 32.91 SO3 2.1

heating value (on dry base)

moisture when used (mean) effective heating value mean particle size

20.25 MJ/kg

20.5 %15.6 MJ/kg4.85 mm(35 % larger than 6 mm; 10 % smaller than 1 mm)

i: r ]

4 EXPERIMENTAL

4.1 Experimental propram

The range of experimental conditions aimed to cover the area tested by other researchers

using coal as fuel. The pressure levels were 5, 7 and 10 bar (absolute). The excess air

amount was changed from 10 to 50 %. The fluidizing velocity varied between 0.6 and1.2 m/s. The thermal input was regulated between 60 and 130 kW. The bed height was

adjusted (from 0.6 to 1.2 m) to the thermal input and presure level, applying a bed

temperature range between 800 and 900 °C. The in-bed gas residence time altered

between 0.5 and 1.5 sec and the gas residence time in the freeboard between 4 and 9 sec.

Experimental data on combustor performance were obtained on combustion efficiency,

fuel bum out, CO, NOx, SO2 and SO3 emissions as well as other subjects such as

particle emission, cyclone performance, bed hydrodynamics, heat transfer, which are not

reported in this paper.

From the 150 hours of operation 120 hours were evaluated, divided into 15 test runs

(continous operation at stabilized conditions).

126 -

4.2 Experimental results

4.2.1 Combustion efficiency, char bum out, CO-emission

The combustion efficiency was generally higher than 99 %. The maximum value of combustion efficiency achieved was 99.6 %, when bed temperature was above 850 °C

and the gas residence time in the reactor was longer than 10 sec. The high combustion

efficiency is due to the high volatile content of peat. About 2 % of the total combustion

losses can be atributed to the unbumt content of fly-ash which is a very low value. Combustion losses were caused mainly by the carbon monoxide emission. It was found

that the combustion efficiency (T|cff %) could be correlated by the following empirical

equation, using multiparameter linear regression:

neff = 96.01+ 0.004 tb + 0.165 x - 0.021 02 (i2 = 0.844)

I L

The influence of bed temperature (tb ,°C ) - the strongest parameter in that correlation -

upon combustion efficiency is shown in Figure 4. The increasing residence time (x, sec)

in the bed increases combustion efficiency. The effect of oxygen concentration (02, %)

seems to be a contradictory, because of the combined effect of bed temperature and

excess air. After reduction of number of parameters the bed temperature and residence

time remained as determining parameters. Further measurements are needed for more

reliable regression analysis of combustion efficiency.

100.0

99.8

99.6comb eff. %

99.4

99.2

99.0800 810 820 830 840 850 860 870 880 890 900

bed temperature, °C

□Li

♦

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Figure 4. Combustion efficiency as a function of bed temperature.

127 -

From the bed material samples, taken 0.3 m above the fuel feeding point, unbumt char

particles were separated and analysed. At this point about 95.. .97 % of the volatile matter

has already left the fuel particles, but only 20.. .45 % of the fixed carbon has burnt off.

Analysing the effect of operation parameters it was found, that bed temperature and the

oxygen concentration had the strongest effects on the fixed carbon burning rate

(mfc,l/sec), since the effect of pressure was negligible.

As mentioned earlier the major part of combustion losses was due to carbon monoxide

formation and emission. The bed temperature (tb ,°C) had the most significant influence

on the CO concentration . The CO concentration (ppm)* as a function of bed temperature

is best described by the following polynomial equation, at an operating pressure of 7 bar :

CO = 5010.7 - 10.65 tb + 0.01 (i2 = 0.61)

Figure 5 shows the effect of bed temperature on CO concentration. The CO emission

seemed to decrease when the presure increased. Generally the CO-emission was below

200 ppm* (90 mg/MJ) under stabilized operating conditions.

bed temperature, °C

Figure S CO-concentration as a function of bed temperature.

]

\

♦The concentrations are calculated assuming 6 % O2 in dry gas.

128 -

4.2.2 NOx-emission

The emission of NOx was in the range of 20 to 130 ppm* (10 to 100 mg/MJ calculated as

N02). Less than 13 % of the fuel bound nitrogen appeared as NOx, assuming that all the

NOx was originated from the fuel. Such low conversion is typical when burning peat in a

fluidized bed. The effects of operation parameters on NOx formation were analyzed, with

respect to pressure (p, bar) and oxygen concentration (02, %).

The following expression was found to describe their influence:

NOx = -9.548 + 27.424 02 - 2.627 p (i2 = 0.958 )

NOx-emission decreases when the system pressure increases, but the effect of pressure is

moderate. As has been reported (1., 2.), the main infuencing parameter on

NOx-formation was the excess air level (02 concentration of the flue gas), as shown in

Figure 6 and characterized by the following linear equation of correlation:

NOx = -27.937 + 27.03 02 (i^^^SO)

0.0 0.5 1.0 1.5 2.0 2.5 3.0 3.5 4.0 4.5 5.0 5.5 6.0 Oxygen concentration, %

Figure 6. NOx-emission as a function of oxygen concentration.

i r

T r

♦The concentrations are calculated assuming 6 % 02 in dry gas.

129 -

The effect of oxygen concentration on NO-formation can be explained by the in-situ

reduction of NO in the fluidized bed by reactive pyrolysis products like CxHy, CH4 and

CO. Using a high excess air level those compounds bum off and therefore NO-emission

increases. In the case of peat pyrolysis significant amounts of ammonia (NH3) appears .

as an important intermediate compound, reducing nitric oxide. These gas components

were measured in the fluidized bed, freeboard, and after the cyclones (Figures 7 and 8).

after cyclones

top of reactor

freeboard

bed

1 10 100 1000 10000 concentration, ppm

Figure 7. Distribution of gas components, p=7 bar, 02=1.9 %.

after cyclones

top of reactor

freeboard

bed

1 10 100 1000 10000 concentration, ppm

Figure 8. Distribution of gas components, p=7 bar, 02=5.9 %.

i f :i

V f 1JL __i

VI

130 -

The gas distributions in Figures 7 and 8 are measured at the same pressure (7 bar) and at

different excess air levels. Each point of the distribution curves represents concentrations

at the local oxygen concentration. The concentration of both NO and reductive

compounds decrease downstream from the feeding point, having the highest

concentrations near the feeding point. About 40.. .50 % of fuel nitrogen was found to be

bound to the char particles separated from the bed material sample (sampled 0.3 m from

the feeding point).

4.2.3 SO2 and SO3 measurements

Peat has a low sulphur content (below 0.2 %wt. d.b.), thus the level ofsulphur dioxide

emission, even without using sorbents, does not exceed the limits of environmental

regulations. When burning peat in the pressurized fluidized bed the SC>2*emission was

below 200 ppm* (160 mg/MJ). The operating pressure had no effect on the

S (^-emission, but at the bed temperature above 840 °C, sulphur dioxide emission

slightly increases with temperature. At the temperature and pressure levels used in

pressurized fluidized bed combustors, S02 can be oxidized to SO3. During the trials a

strong evidence of S03 (H2SO4) presence was observed in the form of rigorous

low-temperature corrosion. Sulphur trioxide was measured in the freeboard, before and

after the cyclones. The SO3 concentration was found to be in the range of 4...6 ppm,

when system pressure was 10 bar. The effect of operating parameters on sulphur trioxide

formation will be studied later. Counter measures were taken to prevent low-temperature corrosion, keeping the flue gas temperature above 200 °C (above the acid dewpoint)

before pressure letdown.

t r■*- *

I l

5 CONCLUSIONS

Some of the results of the 120 hour peat combustion experiments at the Otaniemi

PFBC/G test rig can be summarized as follows:

- Combustion efficiency was above 99 %. The losses due to unbumt carbon in the fly-ash

were very low. Carbon monoxide emission was always below 90 mg/MJ, effected by the

bed temperature.

♦The concentrations are calculated assuming 6 % 02 in dry gas.

Wi

•SI

\

131

- NOx-emissions were in the range of 10 to 100 mg/MJ, strongly influenced by excess

air, since the effect of pressure on NO-formation appeared to be moderate.

- The emitted amount of sulphur dioxide was smaller than 160 mg/MJ (peat has low

sulphur content). Sulphur trioxide was found in the flue gases and low temperature

corrosion appeared on the flue gas duct.- Compared to atmospheric fluidized bed burning of peat slightly better combustion and

emission performance was experienced in the pressurized fluidized bed combustor.

References:

1. Hoy, HJi., Roberts, A.G.& Scott, RL. Operation of a small combustor on dry coal

and on cwm at pressure up to 20 atm. Proc. of 8th Conf. on FBC. Houston 1985

2. Grimethorpe pressurized fuidized bed combustion project. Overall project review.

CEF/U84/21, 19853. Horvath, A. Jahkola, A.& Virtanen, M. Combustion and emission of the pyrolysis

products of peat in fluidized bed. Finnish "Flame Days" Conference, Otaniemi 1986.

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IT-A

11

- 133 -

NEW COAL TECHNOLOGIES

Rainer Muller, Dr.-Eng., Deputy Director

Siemens AG, KWU Group

D-8520 Erlangen, F.R.G.

Abstract

Combined cycle power plants with integrated coal gasifica­tion (IGCC) are a good alternative for coal-based power ge­neration in the nineties. Coal gasification can reduce emis­sions (SO^, N0x, dust) to significantly lower levels than other technologies and can also contribute to minimize waste disposal problems. Besides environmental reasons achievable degrees of efficiency and economic data let expect an advan­tage for IGCC systems in the future.

There are a large number of advanced coal gasification pro­cesses at various stages of development. Each of the processes has its own distinctive characteristics, preferred feed coals, applications, and operating experiences. Most of them still require successful operation of large-scale units before confirmed test results are available.

Based on power-plant-specific requirement criteria and on available information a comparing assessment of selected gasification systems was performed, and results of IGCC system studies were summarized in the form of simplified energy balances. Calculations were performed for three types of gasifiers in a large-scale IGCC power plant, assuming identical boundary conditions for all systems to rule out any influence which might be use of different types of coal and configu­rations .

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134 -1 INTRODUCTION

The environmental problems caused by thermal power plants stem above all from the pollutants liberated in the course of combustion and their dispersion in the atmosphere. Where­as a start was made with effective dust separation some time ago in the Sixties with the installation of electro­static precipitators, retention technologies for sulfur or, as the next step, for nitrogen oxides have only been developed in recent years.

The retrofitting of existing power plants ist going ahead in full swing following the requirements of new ordinances that are coming into force. The emphasis of the engineering and economic efforts is on desulfurization and denitrogeni- zation plant to remove pollutants from flue gas.

FT 1—A. . j

2 COAL TECHNOLOGIES

In principle, there are a number preventing the pollutants bonded leased from the vent stack (Fig.

of possible methods for with coal from being re- 1) .

2.1 Coal Beneficiation

Following the path on the way to electricity generation, the first stage is coal treatment. The coal is first milled, screened and then treated chemically or physically. The ob­jective of this process is to separate rock and minerals from the coal and furthermore, to reduce the sulfur content. While chemical treatment, which entails elaborate decomposi­tion processes, hitherto has only rarely found application, physical methods of beneficiation have become industrial practice. These methods exploit differences in density, surface tension, magnetic or electric properties of the coal consti­tuents for the purpose of separation. They permit considerable reductions of the ash content and of the inorganically bound, that is pyritic sulfur content.

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- 135 -

Coal -Beneficiation -Conversion -Combustion -Flue gas cleaning

Chemical

Leachingprocesses ®0

Oxidativeprocesses ®0

Microbiological processes ®

Gasification and gas purification• Fixed bed•Fluidized bed •Entrained flow w•Molten iron bath

Liquefaction• Direct hydro­

genation• Indirect hydro^

genation

Dry additive processes ®

Fluidized bed combustion ®®

Substoichiometric combustion ®

r ^ iHue gas desulfurization •*- Jprocesses (DeSOx)• Wet ®• Dry ®Flue gas denitrification processes (DeNOg)•Catalytic ®•Noneatalytic ®Simultaneous _ _ iprocesses 1 .. I .

Physical

Grading

Sizing ®0• Wet processes• Dry processes

Dewatering

Fine grinding and mixing with •Water (CWM)•Oil (COM) ®• Methanol (CMM) @

_____________

Burnerconstruction ®Hue gasrecycling ®Combustion chamber design ®

Electrostatic precipitator O

Gravityseparator ^

Fabric filter O

Primary measures4------------------------------------------------- -1—--------------------------------------------------H Secondary measures y

Reduction of ® Sulfur ® NOx O Ash, Dust

Fig. i. Measures for the Reduction of Pollutants Emissions in Power Production from Coal.

2.2 Coal Conversion

Gasification produces a combustible gas which can be scrubbed to obtain almost any purity if processes familiar to the chemi­cal industry are used. Liquefaction by the addition of hydrogen produces-a liquid fuel which also features high purity.

In order to permit the transport, storage and combustion of coal as if it were a liquid, coal slurries have been developed which consist of 3/A coal milled to ultra-fine grade. The re­mainder is made up by liquids such as water, oil, methanol or tensides.

After undergoing treatment processes, this new fuel has a reduced ash and sulfur content.

WI

- 136 -

2.3 Coal Combustion

During coal combustion itself, it is possible to bond sulfur within the ash in the form of calcium sulfite or sulfate by adding limestone, for example. The same principle is applied successfully in fluidized-bed combustion. Different kinds of this technology are summarized in Fig. 2.

Since N0x are only to a small extent formed from the nitrogen bonded with the fuel, attention had to centre above all on the prevention of the formation of nitrogen oxides from at­mospheric nitrogen at high temperatures. In this context, the extremely low combustion temperatures of about 850°C used in fluidized-bed combustion are of benefit.

Conventional combustion techniques can be significantly im­proved in respect of N0x emissions by employing substoichio­metric combustion, which can be achieved by means of suitable burner designs (e.g. 2-stage burner) or burning chambers.

FBC principle

FBC boiler schematic

Bubblinggf1

CA

irculatingFBC jajwmn

BubUiog

Main features: • Combustion of poor fuels and waste materials possible• Sulphur rstentlon In the combustor by moans

of Hmestono (no necessity of FCO)• Low NOx omission duo to low combustion temperature

Special features: • Simple plant concept

• Wear of heating surfaces

• Sophisticated plant concept

• Lower NOx emission due to staged combustion

• Better part-load bohaviour

• Compact components• Shop fabrication• Development problems - Hot gas cleaning -Gas turbine

for dustladen and corrosive gases

Comparison of technical data:VelocityHeat release Combustion efficiencyN0X emissionCa/S molar ratio (90 % S retention)

up to 2.5 m/s1 to2 NlWg/m2

90 to 95 %

200 to 400 ppm2.5 to 3.5

up to 7.0 m/s4 to 6 MWft/m2

90 to 99 %

100 to 300 ppm1.5

up to 2.5 m/s1 to20 MWft/m2

90 to 99 %

100 to 200 ppm1 to 1.5

Fig. 2. Compartson of Fluidized-Bed Combustion (FBC) Principles

137 -

2.4 Flue Gas Purification

The final link in the coal-to-power conversion chain is flue gas cleaning. This is the most suitable of the methods that can be implemented at present in response to the exacting legal environmental requirements imposed on existing power plants. With other methods either the pollutant separation efficiency is inadequate or technical and economic requirements of power plant operation cannot be met.

Is there then no choice but first to dilute the pollutants in coal to 10.000 times their volume in order then to collect them again laboriously in the flue gas?

In this respect coal conversion can be considered a genuine alternative since the progress that has been made in the development of coal gasification systems permits the prediction that this technology will be ready for commercial use in power plants by the end of the retrofitting period in West Germany - that is, by 1993.

2.5 Coal Gasification

Integration of coal gasification systems into combined cycle power plants, the flow chart of which is shown in Fig. 3, results in specific requirements which can be summarized in the form of assessment criteria (Fig. 4). What is needed is a gasification system which can be used with a wide variety of coals, contains tried and tested key components and is flexible in operation. Less sophisticated technical equipment should be sufficient for gas and by-product processing and cost-effective operation of the gasifier should be possible while maintaining a high degree of efficiency.

To achieve a high plant efficiency, the gasification pres­sure has to be matched to the gasturbine inlet pressure. Furthermore, the produced fuel gas must fulfill the purity

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Steam turbinepower plant

WHB_____1 Flua Gai

Fig. 3. Simplified Flow Sheet of a Combined CyclePower Plant with Pressurized Coal Gasificalieation.

f 1

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No limitation by • Coal type• Particle alze• Ash content• Caking behavior

[ Gasifier technique )Solved problems concerning

[ Environment

Extensive reduction of

f EconomicsResulting from

• Coal feed• Ash removal• Heat recovery• Gasifier pressure!* <»•* turbin* pmtura)• Availability• Transient behavior d0Wn’

• Scale up9 Manufacturing experience

• Dust• Sulfur, NOx• By-products, auxiliary materials• Waste water• Space requirement• Overall height

• Low Investment cost• Low O 4- M cost• High efficiency

Fig. Requirements for Gasification ProcessesIntegrated within Combined Cycle Power Plants.

M M M

. 1^ 1

139 -

requirements of the gas turbine to prevent blade erosion and corrosion. At the same time the allowable emission limits must be observed, if possible with an adequate margin since legal requirements may become even more stringent in future depending on the state of the art of emission controls.

2.5.1 Advanced Coal Gasification Systems and Assessment thereof

A distinction is generally made between three basic types of coal gasifiers, i.e. fixed-bed, fluidized-bed and entrained- flow gasifiers. Another type available is the molten-iron-bath gasification system which cannot directly be assigned to one of the "classical" basic types. The schematic diagrams apper­taining to the different types of coal gasification processes are shown in Fig. 5.

r r ]

Fixed bed800 -1000 "C, 10 -100 bar

Coal

Aah

Fluidized bed800 -1000 °C. 10 - 25 bar

Gas

Ash

Entrained flow1500 -1900 °C. 25 - 40 bar

Stag

Molten iron bath- 1500 °C, 1 -3bar

Gaa

Slag♦

Sulfur

Processes Processes Processes Processes

Fig. 5. Basic Types of Coal Gasification Processes.

T r 1

- 140 -

2.5.1-1 Fixed-bed gasifier

Lump coal is use, in the fixed-bed gasifier. In countercurrent flow to the rising gas, it is initially dried, carburized and finally gasified and partly burned, respectively. Depen­ding on the method used, the ash is removed either in a solid or molten state. The technique with dry ash removal is known as Lurgi Pressurized Gasification and has been used for the production of synthesis gas and fuel gas for many years.At the 170-MW-plant in Lunen (FRG), the world's first com­bined-cycle power plant with coal gasification, this type of gasifier using an air/steam mixture as gasifying agent has been installed /6/. Further development of the Lurgi gasifier led to the BGC/Lurgi slagging gasifier of whichseve ral sizes have been tested at We stf ie Id (Scotland) n / .It i s characterized by the use of a wider variety of c oals,inc r eased operational flexibility, a high carbon conve rsionrate , reduced steam requirement and a low CO^ content i nthe raw gas /8/. To date, an oxygen/ steam mixture has beenused exclusively as gasifying agent by th e slagging ga sifier

The KILnGAS process developed in the USA is also a fix ed-bedsystem. In East Alton (Illinois), a demonstration plant is in operation /9/. Gasification takes place in a ported rotary kiln; since air is used as gasifying agent, the carbon cannot be converted completely and this has adverse effects on thermal efficiency. If the KILnGAS gasifier is integrated into an IGCC power plant, the low operating pressure means that the produced gas has to be further compressed to the gas turbine inlet pressure. Since the gasification temperature is rela­tively low, tars and oils are produced as by-products as is the case in all systems with partial carburization of the coal. It is possible to return them to the gasifier but the required waste water treatment then becomes considerably more complex and the necessary equipment has to be more sophisticated than for the fluidized-bed and entrained-flow gasifier.

L.L

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- 141 -

2.5. 1.2 Fluidized-bed gasifier

In the fluidized-bed gasifier, fine-grained and hardly caking coal is used which is converted in a fluidized-bed using an oxygen/steam or air/steam mixture as the gasifying agent.The gasification temperature ist limited to the range below the ash softening point. Lignite and highly reactive hard coal are particularly suitable for this type of gasifier.This system was operated at atmospheric pressure on an indu­strial scale as Winkler gasifier, preferably for the coals mentioned above. Further development of this process towards the pressurized High-Temperature Winkler (HTW) gasifier led to improved carbon conversion, gas quality and specific power.A commercial HTW plant for the production of synthesis gas has been put in operation in Berrenrath (FRG) about a year ago. An oxygen/steam mixture is used as gasifying agent.Further development activities are planned which will raise the gasifying pressure to the level required by the gas turbine so that it can be used in combined cycle plants /10/.

The fluidized-bed gasification process prevents the production of tars and oils as by-products, however, incomplete carbon conversion results in the accumulation of residual char, the chemically-bound energy of which cannot be utilized in the gas turbine process. Attempts are beeing made to solve this problem in the KRW gasifier being tested at the pilot plant in Waltz Mill (Pennsylvania) by a process providing for the return of entrained char and separate steam injection into the upper fluidized-bed zone. This should permit gasifi­cation of a large portion of the accumulated residual char even when hard coal is used. A prerequisite for this, however, is a very high separation efficiency of the recycle cyclones provided for the entrained char. Both oxygen and air can be used as a gasifying agent / 11 /, / 12 /.

142 -

Similar goals are being pursued with the U-Gas process which also has an ash agglomerating zone. A pilot plant is in opera­tion in Southwest Chicago / 13 /.

The CGT experimental plant for hard-coal gasification in Hiinxe (FRG) uses a two-stage process in which residual char from the fluidized-bed is utilized in an entrained-flow stage / 15 /. It is intended to develop the system further especially for application in power plants using air and steam as gasifying agents. The entrained-flow section operated with oxygen is not to be used in future.

2.5.1.3 Entrained flow gasifier

In the entrained-flow gasifier, very finegrained coal ist converted using oxygen and a small quantity of steam at a high reaction temperature into a by-product-free gas mainly containing CO and Most of the ash is removed in a molten state. This process has been used for years on a large scale for the production of synthesis gas at atmospheric pressure under the name of Koppers-Totzek process. Further develop­ment of the process towards pressurized operation led to the PRENFLO gasifier via the Shell/-Koppers gasifier which was tested in Hamburg-Harburg (FRG). An experimental plant is in operation in Furstenhausen (FRG) / 16 /. The development work conducted independently by Shell led to the construction of a demonstration plant.in Houston (Texas) /17/. This system is characterized by high carbon conversion rate and dry coal • dust feeding. Furthermore, all types of coal can be used irrespective of their caking characteristics, their inerts and sulfur contents. High ash contents, however, cause an increase in oxygen consumption which has a negative effect on thermal efficiency and the cost-effectiveness of the overall process.

The GSP gasifier is another entrained-flow system with dry coal feeding. A large-scale experimental facility using

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143 -

pulverized lignite is in operation at the VEB Gas Combine "Schwarze Pumpe" (GDR) / 18 /.

In the Texaco gasifier, coal slurry is entrained which requires less sophisticated systems and control equipment. The attain­able proportion of coal in suspension, however, is limited.In addition, the proportion of water in the slurry leads to higher oxygen consumption and a lower heating value of the produced gas. The Texaco process is the only advanced hard coal gasification process being used for the production of synthesis gas on an industrial scale in several plants.In the "Cool Water" plant near Daggett (California), theuse of this gasifier type for fuel gas production for a 100 MWIGCC power plant is currently being tested / 20 /.

Another slurry-fed entrained-flow system is the DOW gasifier of which several sizes were tested in Plaquemine (Louisiana).A commercial gasifier for the production of fuel gas for an existing combined cycle power plant is currently being commissioned.

In the DOW system, the hot raw gases are cooled in a first stage by injecting coal slurry and this reduces oxygen con­sumption and improves gas quality / 21 /, / 22 /.

The VEW gasifier is the only entrained-flow system with par­tial gasification using air. This process was specially de­veloped for use in combined-cycle plants. The residual char is burned in a separate boiler. A demonstration plant of the VEW gasifier is in operation at Gersteinwerk (FRG) /23/.

2.5.1.A Molten iron bath gasifier

The molten iron bath gasifier was a further development of steel converter technology. Coal is fed into the molten iron bath together with limestone and oxygen. The carbon contained

r ... c _ ]

y r

in the iron is converted into CO and ash and sulfur arechemically bonded in the slag. If this process is developed further towards a process operating continuously at elevated pressure, its main field of application will be in the combined steel and electricity production in connection with new steel production technologies / 2k /.

Fig. 6 contains a table of advanced gasification plants under construction or in operation.

Process Company Leeatien Coal Capacity (M) ■tatiM Product

BGOLur* British Gas Westfield. Soottend sso MOperaMon Fuel Gee. Dsdric Power

KILnGAS AMs Chalmers East Alton, Moots 600 In Operation Fust Gee tor Boder

HTW RhoWschs BraunhoMsworite Ornenmth. FRG 720 In Operation Chernies/ rseddocfc

HTW KemiraOy Oulu. FinrSend 690 Under Cenakrudion Chemieel FeedMoch

KRW KRW Energy SyWems Wahz MB. Pennsylvania 90 In Operation Fuel Gas

KRW ... China 720 Under Canetrudion ...U4es MfeuM of Gas Technology Southwssl Chicago 90 In Operation FustG*

U-Gm ChMbonnogo «• Franco Mahngarhe, France 200 Under Design Fust Gee. Chemical Fsedst.

COT B^ooek HQnse.FRG •6 In Operation Fuel Gas

gaiflnorf Flam

PRENFLO KnniKen»n Fflmtenhauaen, FRG 46 In Operation Fuel Get

Shei Shol CM Company Deer Parik. Texaa 290 Undor Construction Fuel Get

GSP VEBGasCambinoSchwarz# Pianpe

GOR 720 hi Operation Fuel Gas. Chemieel

Teiaoo Tennsoote Easmen King^iBit. Tiwiumi eo tnOparaHsn Chemical FdedModt

Tesaoo Cost WWsr Pn(scl Osggn. CeMnnde sto taOpsiMian Electric Power

Teuco Ubo hnilriss UbeCfey. Japan I960 biOpenataa Chemtoei FsedMocfc

T— RuhrehomWRuhriuMs Osf und Gas FRG

720 to Operation Chemical Fsaddodt

Tesaoo ... Lu Nan. China 960 Under Construction Chemical Feadslcdt

Tesaoo NynAs Energy Chsniicals Camples (NEX)

Nyntohamn, Sereden 4700 Under Design Chemical Feedslodt.Fuel Gaa

VEW VamWgis EtmatTlWrOr Osrms>iesriL.FRQ 240 ■nOpermton Fuel Qaa tor Bator

MattsoJtta

MP MNoa Uriea. Swedsn 240 In Operation Fuel Gaa

KMdtnerIliiimRsnshflns

Suhbacii • ReaarWerg, FRG

•0 Under ComtrutSion Fuel Gaa

Fig. 6. Advanced Gasification Systems Application In Commercial and Demonstration Plants.

145 -

2.5.2 Efficiency and Energy Balance of IGCC Systems

Since different processes are used for the gasification systems presented, the process temperature, gasifying agent requirement and the quality of the produced gas vary. From the energy point of view, processes which make it possible to use a considerable proportion of the coal energy input in the gas turbine in the form of chemically bound energy should preferably be envisaged for integration into combi­ned cycle plants. For high-temperature processes such as entrained-flow gasification, the production of high-pressuresteam during raw gas heat recovery is an essential factor

/in the reduction of exergy losses. The quantity of oxygen and air required for gasification influences the (electrical) auxiliary power requirement of the overall plant via the required compressor drive power.

The simplified flow diagrams of the IGCC power plants are shown in Figs. 7 to 10 for the chosen types of gasifier. Compared to other systems, the combined cycle power plant equipped with the PRENFLO gasifier (Fig. 9) has certain special features in terms of gasifying agent supply and the concept for N0x reduction. The air separation plant operated at elevated pressure is supplied with air taken from the gas turbine air compressor. In addition to the oxygen used for gasification, the produced nitrogen is also utilized.A large proportion of the nitrogen is added to the clean gas; this reduces N0x emissions in the gas turbine and in­creases power due to the increased mass flow. The other systems, however, are designed for independent compression of separation air and saturation of the clean gas with water or steam injection to reduce NO in the gas turbine.

I

I

\

]46 -

Fig. 7 . BGC / Lurgf-Slagger-Based IGCC Power Plant.

Lignte S2S Tonnes/h

Grinding

270 MW

ASU C432 MW

Qroes Power OMpul : 702 MWNet Power Outxit : 628 MW

I I

w r

Fig. 8. IGCC Power Plant with Fluidized Bed Gasitier ("HTW from RHEINBRAUN / UHDE).

- 147 -

Waste Heat Fuel Gas

Grinding

Gas Purificabor

Waste

HI - Combustion

Gas Turbine

Fig. 9. PRENFLOBased IGCC Power Plant.

i . r _ ]

T r i

Fig. 10. Texaco-Based IGCC Power Plant.

148 -

A comparison of the currently achievable efficiences with IGCC power plants reveals advantages for systems equipped with PRENFLO, Texaco and BGC/Lurgi gasifiers. For this reason, calculations were performed for these three types of gasifiers in a large-scale power plant, assuming identical boundary conditions for all systems to rule out any influence which might be use of different types of coal and configurations.The concept for gasifying agent supply and N0x reduction already described for the PRENFLO gasifier (Fig. 9) was used in all systems. Gas turbines attaining ISO inlet temperatures of 1220°C (2230°F) were used to ascertain the efficiency improvement potential resulting from further gas turbine development / 31 /.

The results are shown in Fig. 11, again in the form of a simplified energy balance. In addition to the gas and steam turbine output and the efficiencies, the energy quantities supplied to the gas turbine via raw gas and flue gas heat are also shown.

A

Parcant el anargy input

AReheated clean gaa

with N2 lor gaa turbine

95.1' SB,3

I

90.5 75,7 > VZZL

=gh: : i2.o

ffi:.

392$ 3^0;

| Sensible heat

□ IGCC with BGC/Uirgl

IGCC with PRENFLO

•Xl IGCC with TEXACO

Heat lor ataam turbine Raw gaa heat (RGH)

GT flue gaa heat (FGH)

50.4 49.1

Grose power Gas turbine (GT)

Steam turbine (ST)

Auxiliary Nat power

45,2 47.1 45.9

ST ST^15.7 20.8. ’21,0;

GT GT ; 6t;29.5 26.3.

AP3,6

AP4.3

mNP np *;

41.6 42.8*•••

<AP) (HP)

AP5.6

np" l •: Tip

Percentages are related tom Coi| • HHVCatl (German Enadorl-coal)

Fig. 11. Energy Balance of Large-Scale IGCC Power Plants with Different Gasification Systems(Gas Turbine ISO-Inlet Temperature 1220 °C = 2230 °F,Steam Turbine with Reheating).

149 -

It is apparent that, compared to other systems, most of the coal energy input is contained in the clean gas of the BGC/Lurgi slagging gassifier as chemically bound energy.This means that it has the highest gas turbine power fraction. In the Texaco gasifier, the lowest percentage of chemically bound energy is supplied to the gas turbine with the fuel gas since a larger amount of input coal has to be burnt due to the slurry-feeding. The PRENFLO gasifier designed for dry coal feeding lies between the two systems mentioned above as far as fuel gas energy and gas turbine output are concerned

Since entrained-flow gasifiers operate at high temperatures, efficient utilization of the raw gas heat for steam generation is of particular importance. This is shown by the relatively high steam turbine output fractions in the case of the PRENFLO and Texaco gasifiers. The BGC/Lurgi gasifier operates at lower temperatures so that less raw gas heat is released due to the process conditions. However, since most of it is removed in the scrubbing cooler, it constitutes a source of losses in the steam process.

The auxiliary power requirement of the overall plant i strongly influenced by the required drive power for th and nitrogen compressors. For this reason, difference the oxygen requirement directly affect the auxiliary p requirement. This gives the PRENFLO and BGC/Lurgi gasi advantages over the Texaco system.

se oxygen s in ower f iers

The resulting net efficiencies lie between 40 and 43 % and increase starting with Texaco followed by BGC/Lurgi and the PRENFLO gasifier.

Fig.using

2 shows the Sankey the PRENFLO system

diagram for an as an example.

IGCC power plant

I I ]

150 -

Fig. 12. Sankey Diagram of a Large - Scale IGCC Power Plant with PRENFLO Gasifier(Gas Turbine ISO Inlet Temperature 1220 ° C = 2230 °F, Steam Turbine with Reheating).

2.5.3 Comparison with other Power Plant Processes.

The net efficiencies achievable in IGCC power plants are compared with those of other power plant processes in Fig. 13. In addition to pulverized-coal-fired steam power plants with flue-gas desulfurization (FGD) and DENOX, combined-cycle processes with pressurized fluidized-bed combustion (PFBC) and natural-gas-fired combined cycle plants are included in the comparison.

The net efficiencies of the combined gas and steam turbine processes improve considerably compared to those achieved with pure steam processes (Band A). For systems with fluidized- bed combustion (Band B), the maximum process temperature and thus the gas turbine inlet temperature are limited to below 900 °C (1650 °F). IGCC power plants (Band C), however, offer the possibility of fully utilizing further gas turbine

- 15.1 -

development towards higher inlet temperatures. Conversionlosses and the increased auxiliary power requirement associatedwith coal gasification, however, cause a drop in efficiency- in comparison with natural-gas-fired combined-cycle plants(Band D) - by 5 to 10 percentage points. The range is fairly ^ ^wide since the efficiencies attainable in IGCC power plantsnot only depend on the gasification process used and itsauxiliary power requirement, but also on the configurationand the chosen gas purification system.

Net efficiency (HHV-related)

Pulverized coal-fired power plants with FGD and DENOX

Combined cycle power plants with natural gas firing

•PRENFLOCombined cycle power plants with pressurized

fluldlzed-bed combustionPRENFLO

BGC/Lurgi

•Texaco

Texaco1Combined cycle

power plants with Integrated coal

gasification (IGCC)

Gas turbine Inlet temperature (ISO)

—I------------------------------ 1------------------------------1------------------------------ 1-----------—"1600 1800 2000 2200 t’FJ

Fig. 13. Comparison of Station Efficiencies for Different Types of Combined Cycle Processes and Pulverized Hard-Coal-Fired Power Plants.

i r

t r

I l

Coal gasification power plants with two or more gasification trains and gas turbines also attain high efficiency during part-load operation. They are therefore not only suitable for operation in the base-load range, but also in the inter­mediate-load range which is the most cost-effective operating range for hard-coal-fired power plants in the Federal Republic of Germany.

152 -

2.5.A Environmental and Economic Aspects

Combined-cycle plants with coal gasification will become a valid alternative to conventional technologies only if they prove advantageous from the point of view of cost or emissions in addition to the higher efficiencies attained. For this rea­son, these two aspects are briefly discussed in the following.

Fig. 14 shows the supply and disposal flows, and the emission values of a conventional hard-coal-fired power plant with FGD and DENOX and of an IGCC power plant. The power plants with coal gasification are not only advantageous as far as SO^, N0x and CO2 emissions are concerned, but also with respect to the disposal of solid residues.

In a conventional power plant, a considerable amount of gypsum is produced due to the use of lime for FGD purposes; utili­zation of the gypsum will become increasingly difficult in the future. The IGCC power plant, on the other hand, requires no additional absorbent as the sulfur is bonded in the form of

with regenerable scrubbing fluids. Marketable elemental sulfur produced from this can be sold in a concentrated form.

For the use of hard coal, Fig. 15 provides a comparison of the cost structures of pulverized-coal-fired power plants with FGD and DENOX and IGCC power plants. It is particularly noticeable that the actual power generating equipment accounts for less than half of the total costs of an IGCC plant, whereas it accounts for 3/4 of the total investment in a pulverized- coal- fired steam power plant. The reason for this is the extremely low specific cost of the gas turbine in comparison with the steam turbine cycle. Total investment is expected to be about the same for both power plant types within the

W„JL.

I T

Y Y

\

153 -

Coal (6366)

n159 wagons J I

‘O

Umasiona (217)

_i gr^-i090 090 £90 090 090 090

0

SO, :22,4

NO. :1U iz_Hi±jUo c

Convantlonal Power Plant ( ^ -36%)

0 0 0

100 m

Coal (5376)

£t>C30C3 134 waoons

0dally amcunta(metric tons)

P~[> so2 :«

o “NO, :6.6

IQCC-Powar Plant ( t) -43%) *)0 0 0 0C3CJCTC3 C3C3CTC3 C3 CTCT C3 C3CJC3C3

Ash (637)

coco earneammoaoa oo oa ooOQ oo 00 ooGypsum (315)

C3C3CT C3 aQ

C02 SLJ W IwJ Fine DustC3C0C3 W

CD

Sutlerco, s

(17000) =| C3CJC3

Sag (488)

(63) (14400) =

*) based on PRENFLO-GasificationFig. 14. Comparison of Supply and Disposal Flow in Hard Coal Power Plants

(700 MW(et); 73 % C; 1,2 % S; 10 % Ash; HHV w 26500 kj / kg; IGCC: 98 % Desulfurization; Conventional Power Plant: 400 mg S02 / m3 Dry Hue Gas).

[%] |100 -

80 -

60 -

40 -

20 -

Electrostatic precipitatorDENOXFGD25%

Steam generator Civil works

34%

Mechanical equipment Electrical equipment

Others41 %

Steam turbine 12%

Wet dust removal 6 %Desulfurization

7%Gasifier9%

Waste heat boiler 11 %

Air separation 14%

doal preparation8%

Combined cycle unit45%

Pulverized coal-fired power plant with FGD and DENOX

IGCC power plant

Fig. 15 . Cost Structures of Pulverized Hard-Coal-Fired Power Plants with FGD and DENOX Compared to Combined Cycle Power Plants with Integrated Coal Gasification (IGCC).

]

IT.]

t r

154 -

normal range of cost variation. On account of the fact that an IGCC power plant requires 20 % less coal, its electricity supply is cheaper than that of a conventional power plant of th-' same capacity by the same order of magnitude. And the amount to be received for the marketable chemical product sulfur has not yet been considered.

2.5.5 Future Prospects

Just as with any other new technology, it will take time following the development and test phase for power plants with coal gasification to reach commercial maturity. Fig. 16 illustrates a development schedule based on conditions in Ger­many. It can be seen that about 10 or 15 years lie between the testing phase cf the components and commissioning of a large-scale commercial plant. Widespread application of IGCC systems cannot be expected before this period has elapsed.

r r

I T

T T

Erection and Operation of Commercial Plant Phase 4 ]

Erection and Test of Demonstration Plant Phased ]

.XXXXXXXXXXXV1)^xxxxxxxxxxxx*

Development of Components and Overall Plant

Phase 2

ProjectDefinition Phase 1

r ri—i—i—i—i—i—i—i—i—i—i—i—n—i—i—i—i—i—i—i - -

1980 1985 1990 w 1995 2000Year

Fig. i6 . IGCC Power Plant • Development Schedule.

- 155 -

The fact, that so many different gasification systems are currently being developed means that it will be necessary in the nineties when confirmed test results are available to start a selection process among the systems suitable for application in power plants. If the power-plant-specific requirements defined in Fig. k are used as selection cri­teria, it is expected that only a small number of processes will be suitable for commercial application.

3 CONCLUSIONS

In principle, all advanced types of gasifier operating under pressure are suitable for use in a power plant. The power- plant- specific requirements, however, are fulfilled in different ways by the various systems. In addition, a com­parison of the basic types of gasifiers has shown that they can be assigned to special fields of application in which they have specific advantages over other systems. For example, fixed-bed gasifiers are particularly suitable for plant con­cepts providing physical separation of the gasifying unit and the power plant, and fluidized-bed gasifiers are more suitable for highly reactive types of coal. Entrained-flow gasifiers are universally applicable and particularly suit­able for concepts with a high degree of thermal integration.

Integration of coal gasification systems into combined gas and steam turbine power plants is technically feasible and provides a high potential for improvement in efficiency as against conventional coal-fired power plants. In addition, IGCC power plants are economically and ecologically superior to conventional technologies.

Although further development work is necessary for certain components such as gasifier, heat recovery boiler and high- temperature gas turbine, the results obtained to date give

- 156 -

confidence that this technology will be ready for commercial application in the nineties. This IGCC technology also allows better adaptation of power generation to the increase in power consumption by way of phased construction and the avoi­dance of excess capacities which is of particular interest from the economic point of view. In addition, gasification is suitable far retrofitting of existing steam or combined gas and steam turbine power plants to IGCC systems.

A ACKNOWLEDGEMENT

The work described in this paper was performed by Siemens AG, KWU Group, sponsored by the German Federal Ministry for Research and Technology.

REFERENCES

/I/ Frewer, H.: Strukturwandel in der Technik fossilbeheiz-ter Kraftwerke in der Bundesrepublik Deutschland.VGB Kraftwerkstechnik 1986, 66(1), 303 - 326.

/2/ Breucker, H., and Stadie, L.: SteinkohlebefeuerteDampferzeuger fur Kraftwerke mit hohen Dampfzustanden. VGB Kraftwerkstechnik 1983, 63(1), 29 - 37.

73/ .Ewald, J., Haase, H., and Zorner, W.: DampfturbinenfUr Kraftwerke mit hohen Dampfzustanden. VGB Kraftwerks­technik 1983, 63(1), 37 - 45.

/A/ Giovanni, D. V., and Armor, A. F.: Die Verbesserungdes thermischen Wirkungsgrades konventioneller Kraft­werke. VGB Kraftwerkstechnik 1982, 62(2), 78 - 89.

75/ Schiffers, U., and Stocksmeier, H.: Kombiniertes Gas-Dampfturbinenkraftwerk mit Kohlevergasung, Projektde- finition (Phase 1).BMFT-Report-FB-T86-036, February 1986.

T lpr

i r

t r

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/6/ Dorstewitz, U., and Wien, H.: KDV-Anlage Liinen.Report TB 201 of the VGB Conference "Kohlevergasung in der Energietechnik".

77/ Davies, H. S., Lacey, J. A., Scott,J.E., Herbert, P.K.,and Vierrath, H. E.: Application of the British Gas/Lurgi Slagging Gasifier for Combined Cycle Power Generation. Symposium on Coal Gasification and Synthetic Fuels for Power Generation, San Francisco, April 16, 1985.

78/ Thomas, K. M.: Some Scientific Aspects of the Develop­ment of Slagging Gasifier. NATO ASI Ser., Ser. E, 105 (Carbon Coal Gasification), A39-53.

79/ Garside, P. G.: KILnGAS Process for Producing CleanGas from High-Sulfur Coal; Major Technical Accomplish­ment and Plans. Energy Progress 1986, 6 (1), 5-11.

710/ Fabianek, G., Adloch, W, and Wischnewski,R.: Standdes Rheinbraun HTW-Verfahrens. Haus der Technik Con­ference "Die Projektierung von Kohleveredelungsanlagen und deren Veredelungsbetriebe." Essen, November 6-7, 1986.

711/ Haldipur, G. B., Cherish, P., Lewandowski, D. A.,and Pavel, J.: Hot Gas Cleanup Using the KRW Gasifier. Fifth Annual EPRI Contractor's Conference on Coal Gasification. Palo Alto, California, October 30-31,1985-

712/ Salvador, L. A.: The Appalachian Project: Clean Power from Coal. West Virginia Conference on Clean Coal Technology. Morgantown, June 19, 1986.

713/ Patel, J. G., and Schora, F. C.: The U-Gas-Process:Clean Fuel Gas from Coal. Urja 1986, 19 (A), 283-91-

Palat, P., Schora, F. C., and Patel, J. G-: U-Gas Process CdF. Urja 1984, 17(3), 187-94.

7 14/

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/ 1 5/

/ 16/

/ 17/

/ 18/

/1 9/

/20/

/2 1/

/22/

/23/

Carbon Gas Technology GmbH: CGT-Kohlevergasung.Company Brochure.

Pohl, H. C.: Commissioning of the PRENFLO Test Unit.Sixth Annual EPRI Contractors's Conference on Coal Gasification. Palo Alto, California, October 15-16, 1986.

Heitz, W.L., and Nager, M.: Status of the Shell Gasi­fication Process (SCGP). Fifth Annual EPRI Contractor's Conference on Coal Gasification. Palo Alto, California, October 30-31, 1985.

Schingnitz, M., Brandt, H., Berger, F., Gohler, P. and Kretschmer, H.: GSP-Verfahren: Druckvergasung von Braunkohlenstaub. Erdol und Kohle, Erdgas, Petrochemie vereinigt mit Brennstoff-Chemie 1987, A0(2), 76-80.

Ruhrkohle Oel und Gas: Business Report 1985.

Spencer, D. F., Alpert, S. B., and Gilman, H. H.:Cool Water: Demonstration of Clean and Efficient New Technology. Science 1986, 232 (2), 609-12.

Borneman, G. A., and Sundstrom, D. G.: Commerciali­zation of the DOW Gasification Process. American Chemical Society, Div. Fuel. Chem. 1986, 31(2), 270-73.

Fisackerly, R. H., and Sundstrom, D. G.: The DOW Syngas Project Overview and Status Report.

Weinzierl, K.: Combined Cycle Process with Integrated Coal Gasification; VEW-Coal Conversion Process. Sympo­sium Advanced Coal Power Technology. Malmo, November 6-7, 1985.

I2kl

/25/

726/

727/

728/

729/

730/

731/

Henrich, G., Pook, H. and Waldhecker, H.D.: Kohle­vergasung im Eisenbad nach dem MIP-Proze/3. Energie 1986, 38(10), 39-41.

Simbeck, D. R., and Dickenson, R.L.: Integrated Gasifi­cation, Combined Cycle for Acid Rain Control. Chemical Engineering Progress 1986, 82(10), 28-33.

Coal Gasification Systems: A Guide to Status, Appli­cation and Economics. EPRI-Report AP-3109, June 1983.

Cost and Performance for Commercial Applications of Texaco-Based Gasification-Combined-Cycle Plants. EPRI-Report AP-3A86, April 1984.

Evaluation of the British Gas Corporation/Lurgi Slagging Gasifier in Gasification-Combined-Cycle Power Generation. EPRI-Report AP-3980, March 1985-

Cost and Performance of Kellogg Rust Westinghouse-Based Gasification-Combined-Cycle Plants. EPRI-Report AP-4018, June 1985-

Petersen, G. T.: KILnGAS Coal Gasification Process- Status. Coal Gasification and Synthetic Fuels for Power Generation Conference. San Francisco, California, April 14 - 16, 1985.

Muller R., Karg J.: Assessment of Coal Gasification Systems for Combined Cycle Power Plants. Jt.ASME/IEEE Power Generation Conference Miami Beach, Florida,October 4-8, 1987. The American Society of Mechanical Engineers, New York, October 1987.

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THE RHEINBRAUN HTW-GASIFICATION TECHNOLOGY FOR COMBINED CYCLE POWER

GENERATION

Eberhard Nitschke, Dr.

UHDE GmbH

Dortmund, FRG

Abstract

One of the most advanced coal gasification processes is the Rheinbraun

HTW-fluidized bed gasification. The process has been developed by

Rheinische Braunkohlenwerke AG, Cologne, in close cooperation with Uhde

GmbH as the engineering partner.

In 1978 a pilot plant (capacity up to 30 t/d) was put into operation.

The influence of different gasification parameters on the performance

of the plant was studied. Also Finnish peat and Australian lignite with

a high sulphur and salt content were tested. Based on the results a

commercial scale demonstration plant with a capacity of 720 t/d dried

lignite was built and is successful in operation since 1986.

Together with Rheinbraun and Kraftwerk-Union Uhde has performed as

study for an IGCC-plant (600 MW), based on HTW-process. The concept has

great advantages with respect to efficiency. Also emissions from the

power station can be drastically reduced.

1 INTRODUCTION

The Rheinische Braunkohlenwerke AG (Rheinbraun) which has economically

mineable brown coal deposits of more than 35 000 million tons, is also

engaged in the development of new coal beneficiation processes. Besides

coal liquefaction, the coal gasification is the most important and

promising route.

162 -

The experience gained in the fifties from the Winkler gasification

process at Union Kraftstoff a Rheinbraun subsidiary in Wesseling was

the basis for the development of the high-temperature Winkler (HTW)

process for the production of synthesis gas or low-BTU gas.

Rheinbraun signed an agreement with the engineering company Uhde GmbH,

Dortmund, in order to cooperate in the development of the Rheinbraun

HTW process. Uhde is the engineering partner of Rheinbraun and, conse­

quently, it was responsible for the engineering, erection and start-up

of the Wachtberg pilot plant as well as for the commercial-scale Ber-

renrath plant. Uhde is also the contractor for the peat gasification

plant in Finland.

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2 DEVELOPMENT OF THE HTW PROCESS

The HTW process has been developed on the basis of the wellknown

Winkler process in which coal or coke is gasified with the aid of air

or oxygen/steam mixtures in a fluidized bed. The first step of the

development aimed at the production of a reducing gas for the direct

reduction of iron ore "in an in-line heat". Therefore a high gasifi­

cation temperature - which is now used to designate this process - had

priority over a high gasification pressure which was initially planned

to be 5 bar.

During the technological development, the interest concentrated upon

the production of gas for chemical synthesis and in order to obtain a

high gasifier performance, it was decided to use a process pressure of

10 bar.

A further step is currently under consideration, i.e. the use of the

process potentials in a combined gas and steam turbine process with a

pressure of over 20 bar.

The development work on the HTW process, therefore, mainly concentrates

on the improvement of the gasifier output with the aid of higher

T T

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- 163 -

operating pressures and on the solution of all problems related to

process engineering and operation.

3 THERMODYNAMIC AND REACTION-KINETIC FUNDAMENTALS

Reactive carbon carriers such as wood, peat, brown coals, lignites and

subbituminous coals have a high content of volatile matter, i.e. in the

presense of heat they tend to decompose rapidly to form gaseous and

liquid fractions and char. In a Winkler-type fluidized-bed gasifier,

the so-called carbonization or devolatilization process can be

controlled in such a way that gaseous products and reactive char are

only obtained as intermediates which further react in the gasifier.

Since the a.m. carbonization or devolatilization is an endothermal pro­

cess, reaction heat has to be introduced by burning part of the inter­

mediates. Thus, each feedstock used in the fluidized bed of a Winkler

gasifier involves a specific temperature at which - with a constant bed

depth - a steady state is reached between the required quantity of

feedstock and the heat consumed by its endothermal devolatilization or

gasification, on the other hand, and the heat obtained by burning the

feedstock or its intermediates, on the other hand. Table 1 shows that a

higher volatile content leads to a lower temperature in the fluidized

bed. Said specific temperature is also crucial for the composition of

the gas from the fluidized bed. Feedstock with a low degree of coalifi-

cation such as wood yields also liquid carbonization products which may

cause problems in the downstream gas treatment units.

In order to prevent such problems, it is necessary to raise the tem­

perature in the so-called freeboard above the fluidized bed by in­

jecting additional oxygen or air to such an extent that unwanted by­

products are converted into useful gases. This is possible because the

only substances that are present in this zone are small quantities of

higher hydrocarbons and the entrained degassed char particles. They are

contained in the gas stream in relatively low concentrations so that

the endothermal reactions are impeded.

164 -

Table 1. Comparison of performance data from different feedstocks used in the HTW process development unit.

wood peatRhenish

brown coalLignite

subbit uminoui

coal

degree of coallication 0-0.07 0.18 - 0.36 045 - 0.48 0.52 058 - 059

votetle content % t 82-87 61-73 49 - 55 46 32 - 35

char reactivity with COg at 900*Cand 30% conversion

cm3/g.s 10- 15 5- 10 4- 8 3 - 4 05 - 0.7

gasification temperaturein the IMdteed bed

•C 550 - 650 700 - 750 750 - 800 760- 810 800 - 850

specific input t. maf/m2.h 1.12 0.69 0.73 0.52 0.35

specific synthesis gas yield

m3(STP)/t/naf 1020 1170 1370 1590 1650

specific synthesis gas output

m3<STP)/m2.h 1U0 1040 1000 630 650

gasification pressure: 1 bar energy loss: 10% carbon conversion: 95%

r r-

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A mean content of volatile matter permits an increase in the heating

value of the gas by reducing the oxygen quantity fed to the freeboard.

This prevents the conversion of a major portion of the primary

methane produced in the fluidized bed into CO + H at high temperatures.

This method is particularly effective at higher pressures as shown in

Figures 1 and 2. When, for instance, a pressure of 20 bar is used and when the temperature drops from 950 to 850 °C, the methane content in

the raw gas increases from about 4 % to 10 % (calculated in terms of

thermodynamics).

W

165 -

gas composition (%voi)

80---y

___ 1____60---- - +--------+ -

--1- --------- 1

C_ J- - - - - - - 1- - - - -20-------4

temperature (°C)

Figure 1. Carbon/steam equilibrium at 1 bar total pressure.

gas composition (Sbvol)

- -I----------- 1-----------\----------- 1------60-------------r

---------- r_

1000 1200temperature (°C)

Figure 2. Carbon/steam equilibrium at 20 bar total pressure.

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4 THE HTW PROCESS DEVELOPMENT UNIT

In the seventies, the development of the HTW process was starting by

building a process development unit at the Aachen Technical University

with a solids input capacity of about 50 kg/h. In this plant operating

under atmospheric conditions, the influence of an increase in temper­

ature on the performance parameters was investigated, i.e. carbon con­

version, gas yield and specific gas output. This plant is still being

used for testing the suitability of different feedstocks for the HTW

process. Table 1 reflects the main results of said tests which show

that a higher rate of coalification - from wood to subbituminous coal -

causes a significant decrease in the solids input due to lower reacti­

vity, while the yield of useable gases such as synthesis gas signifi­

cantly increases due to the lower oxygen content and the higher heating

value. The slight decline in the synthesis gas output, however, can be

partly set off by a higher oxygen supply and, consequently, a temper­

ature rise in the freeboard.

The verification of such performance data in continuous operation also

depends on the quantity of inorganic matter such as ash in the

respective feedstock and the ash fusibility.

Table 2 reflects to feedstocks with different ash composition. Type I

is an "acid" ash with a particularly favourable fusibility as, for

instance, in the case of hard or subbituminous coal. Type III is a

feedstock with a poor ash fusibility. This applies, for instance, to

some American lignites. This "basic" ash, however, is also found in

coals with a significantly better ash fusibility, as is the case with

Rhenish brown coal which contains more alkaline earth oxides but has a

lower alkali content (Type II). In general, wood has a very low ash

content.

Since the conversion temperatures are relatively low in the fluidized

bed and also in the freeboard as mentioned above, there will be no

serious trouble due to agglomeration of inorganic matter.

Table 2. Ash melting behaviour and ash composition.

ash melting behaviour(reducing atmosphere)

I 11 III

initial deformation point 1520 1265 1180

hemispherical point °C 1570 1315 1210

flow point 1580 1375 1230

ash compositionsio2 49,3 74 29,7

ai2o3 39,0 2.7 9.1

Fe203 3.4 144 14.5

CaO 2,3 42,4 15,9

MgO Xwt 2,0 9,3 5.1

Ti02 1.3 0.4

Na20 1,1 1.3 6,8

KjO 0,3 0.3 0.6

Mn304 ^O.l - 0.6SOj 14 22,2 17.9

In the case of a subbituminous coal with lower reactivity, however,

comparatively high temperatures are required in the gasifier in order

to obtain sufficient yields. These temperatures may even be as high as

the ash sintering temperature or the initial deformation point causing

ash agglomeration and/or gasifier choking, particularly in oxygen or

air injection zones where the reaction heat cannot rapidly dissipate

due to endothermal reactions. Consequently, peak temperatures may

occur. To avoid these problems, subbituminous coal with a high ash

fusion point should only be gasified.

:68 -

5 THE HTW PILOT PLANT

After the tests in the a.m. process development unit, a pilot plant for

the gasification of Rhenish dry brown coal using the HTW process went

on stream in 1978 on the site of the Rheinbraun brown-coal upgrading

plant in Wachtberg near Cologne (Figures 3 and 4). The engineering for

this plant, i.e. design, construction and start-up, were performed by

Uhde GmbH, Dortmund. The objective of the test operation was to

optimize the synthesis gas production.

The feedstock is dry brown coal with a moisture content of 18 % which

is also used for briquetting. Depending on the test target, the

moisture content of the coal can be reduced to 10 - 12 % by means of a

pneumatic conveyor for drying. In the downstream lock-hopper system,

the coal is subjected to the gasification pressure and fed by a screw

conveyor to the gasifier which has a refractory lining.

gasification agents oxygen/steam or ai pressure up to 10 bar

throughput (browncoal ) up to 1800 kg/h. raw gas output up to 4000 inVh.temperature up to 1100 °C .

I _C

Figure 3. Flow sheet HTW pilot plant.

The gasifier Is divided Into two rones. The gas leaves the fluidlred

bed where the solids concentration is relatively high, entraining a

considerable amount of char particles. In the second zone, the so-

called freeboard or post gasification zone where the solids density is

lower, additional oxygen and steam are injected in order to ensure a

carbon conversion at temperature of approx. 1 000 °C and in order to

adjust the gas composition such that the specification is met, parti­

cularly with regard to the lowest methane content that can be obtained.

The raw gas flows from the gasifier outlet via a short refractory-1ined

pipe directly into the first cyclone in which the coarser dust

170 -

particles are separated and recycled to the gasifier. In the second

cyclone, the major part of the remaining fine dust particles is removed

from the raw gas and - as is the case with the ash withdrawn from the

gasifier bottom - discharged via a cooling screw conveyor and depressu­

rized in a lock-hopper system. The purified raw gas is then cooled in a

waste heat boiler and burnt in an adjacent power plant.

During the first years, certain problems had to be solved in the HTW

pilot plant. The plant availability was significantly reduced by ash

agglomeration, particularly in the section of the oxygen injection

nozzle. Meanwhile, a high plant availability has been reached by means

of an adequate nozzle arrangement in the gasifier and a new configura­

tion of the raw gas outlet to the cyclones, the dust recycling system,

waste heat boiler, water scrubber and bottom product (residual char)

discharge. The longest period of continuous operation was seven weeks.

In late June 1985, the plant was shut down because the operating expe­

rience and the process engineering data regarding the investigated

10-bar mode were conclusive. Since the start-up, the plant had been in

operation for almost 38 000 hours, i.e. 27 400 hours under gasification

conditions. During this period, more than 21 100 t of Rhenish brown

coal plus approx. 200 t of Finnish peat were processed within the scope

of various test programs.

In December 1986 and January 1987, additional tests were performed with

South Australian brown coal having a high sulphur and salt content (see

chapter 7.2).

Table 3 shows a comparison of the operating results of the HTW pilot

plant and the conventional Winkler gasifiers using an oxygen/steam

mixture for gasification. As can be seen, the carbon conversion rate

increased from 91 % to 96 % and, consequently, the synthesis gas yield

per ton of brown coal was Improved accordingly.

A typical raw gas contains the following: CO 33 % by vol., H2 26 % by

vol., CH,, 3 % by vol., C02 19 % by vol. and H20 19 % by vol. On a wet

I I

T T

I E

J /1

Table 3. Gasification with oxygen and steam.

Winklergasifier HTW

gasifying conditionstemperature 0 C 950 1000pressure bar 1.2 10

gasifying agentsm3 (STP)/kg*bCm,ioxygen 0.42 0.40

steam kg/kg’bc*., 0.18 0.33

spec, yield (CO»H2) m3 (STP)/t*bcm., 1460 maul 580

spec, output (CO + H2) m3 (STP)/ m2 • h 2120 •naK.7 745

C-conversion % 91 max. 9 6

be =brown coal

and C02 free basis it contains CO 53 % by vol., H2 42 % by vol., and

CH4 5 % by vol.

The synthesis gas output of almost 1 600 m3 (STP/t dry brown coal

(calculated moisture- and ash-free maf) corresponds to 95 % of a

maximum synthesis gas output of about 1 650 m3 (STP) C0+H2/t brown

coalmaf that can be reached thermodynamically under comparable gasi­

fication conditions.

Since the pressure was increased to 10 bar, it was also possible to

treble the specific synthesis gas output, i.e. more than 7 700 m3 (STP)

C0+H2 per hour and m2 gasifier cross-section compared with the

atmospheric Winkler process.

The dry and C02-free raw gas has a lower heating value of approx.

13 000 kJ/m3 (STP). When changing over to higher pressures of 20 to

W-X ]

i t : i

T T 1-JL __ I

172 -

25 bar, higher methane contents can be expected and the heating value

may be as high as about 14 000 kJ/m3(STP) according to a preliminary

estimate.

The unproblematic low-ash Rhenish coals with a low ash content of up to

approx. 5 wti as well as high-ash coals with an ash content of up to

20 wt% were processed. (Both figures are related to dry coal with a

10 % moisture content.)

Figure 5 shows various overall test results and the wide range of spe­

cific coal throughput parameters that have been obtained. The figure

also reflects that the specific synthesis gas yield of about ± 9 % over

a wide range does not depend on the coal throughput and, consequently,

the specific synthesis gas output (as the product of specific synthesis

gas yield with coal throughput) rises in a linear manner. The scatter

of the measuring data is due to the following factors: inevitable

spec, synthesis gas output

(m3(STP)/in2.h)

02 supply: 0.38 to 0.42 (m3(STP)/kg dry br.c.mat)

■ ■ low-ash briquetting coal • * high-ash steam coal

spec, synthesis gas yield (m3(STP)/t dry br.c.maU

coal input, mal (t/m2-h)

f T

Figure 5. HTW pilot plant - synthesis gas yield and synthesis gas out­put as a function of coal input.

- 173 -

measuring errors in analytics and mass flow measurement and varying

distribution of the gasification agent injected into the gasifier, on

one hand, and normal variations of the reactivity of the utilized brown

coal, on the other hand. In the Rhenish brown-coal district, 14 diffe­

rent so-called litho-based brown coals have been found so far. While

their reactivity varies considerably depending on the gasification

agents used in the HTW process such as steam and carbon dioxide, their

distribution related to the coal is almost stable when large equipment

is employed in opencast mining.

Figure 5 also shows that the mean output figures for the processing of

high-ash coals are lower than those for low-ash coals but are still

within the a.m. range. Considering the fact that, when processing high-

ash coals, an inert material portion of up to 15 wt% (mostly silica

sands) has also to be heated and that, together with the bottom dis­

charge of this inorganic material, carbon losses cannot be avoided in

the process, the result can be regarded as very satisfactory.

Figure 6 shows a comparison of the synthesis gas output initially ex­

pected in the case of a pressure increase to 10 bar and the test re­

sults. It was possible to exceed the expected figure of 6 000 m3 consi­

derably: almost 8 000 m3(STP)/m2 were reached. Compared with the atmo­

spheric Winkler gasifier, this increase corresponds to an uprating

index of about 0.6, previously 0.5. From these positive results it is

concluded that a further increase in pressure will also improve the

output. In this connection, the fluid mechanics have to be considered

as well. The above specific synthesis gas outputs were reached at

superficial velocities of about 1.6 m/s in the freeboard. Extrapolation

of these conditions to higher pressures indicates even higher perform­

ance parameters than can be expected, i.e. in excess of the pressure

uprating index of 0.6. Exact figures, however, can only be obtained by

testing.

The objective of the HTW process development is to use higher pres­

sures. Therefore, it has been decided to use the former pilot plant for

hydrogasification of brown coal at the Rheinbraun subsidiary in Wesse­

ling as new HTW pilot plant.

?.W ]

I T ]

T r 1

wi,

'•*1

174 -

Jnisooo-

13 0 CO-

12 OOO-

11 000-

V. 365.02

7000-

6000-

5000-

3 000- = thermodyn calculations< possible range of

synthesis gas output2000-

gasification pressure (bar)

Figure 6. HTW process. Test results and extrapolation.

The object of the new pilot plant is, among others, to gain experience

with the operation of the HTW gasifier at an elevated pressure up to

25 bar required for IGCC technology. In addition, the potential of the

HTW process is to be investigated with this pilot plant. This refers to

its probable range of possible feedstocks and also to an improvement in

economic viability by raising its efficiency and lowering plant costs.

In particular, an increase of the fluidizing gas velocity up to a cir­

culating fluidized-bed mode is to be tested - a logical further step of

development since the HTW gasifier has already been operating success­

fully using dust recirculation to the fluidized bed.

\

1 /o

As the fluidized-bed pilot plant an existing and currently mothballed

pilot plant and its infrastructural installations can be used, the

necessary refitting of which is scheduled to be completed by mid-1989.

During the last weeks the first orders for the new plant were placed.

The refitted pilot plant will have a throughput of 6.5 t of dry lignite

per hour. The design allows for gasification both with 02/steam and

air/steam. In addition to gasifier testing, it is planned to investi­

gate also the problems of a hot gas clean-up system.

6 THE HTW DEMONSTRATION PLANT

On the basis of the HTW pilot plant experience, Rheinbraun has built a

HTW demonstration plant for converting dry brown coal into synthesis

gas. The engineering was performed by Uhde GmbH, Dortmund. This demon­

stration plant is located on the site of the Rheinbraun coal upgrading

facilities in Berrenrath near Cologne and includes one HTW gasifier and

all downstream units for further gas treatment required to produce

syngas.

The demonstration plant represents the last step of the HTW process

development towards the construction of a commercial-scale synthesis

gas unit. It is required to demonstrate the complete process line from

the coal to the final product and offers a new and reliable technology

for future applications.

The synthesis gas from this plant is piped to the Union Kraftstoff

plant in Wesseling where it is used in the existing methanol synthesis

gas system in order to replace the synthesis gas still produced from

residual oil.

With regard to the coal preparation, process steam/power plant and

off-sites for waste water disposal, the demonstration plant is inte­

grated into the existing infrastructure of the Berrenrath complex of

Rheinbraun (Figures 7 and 8).

176 -

The flow sheet shows that a particularly favourable gas treatment train

was chosen for the process line in order to reduce energy losses: a hot

water scrubbing system has been installed directly behind the raw gas

cooler and has two simultaneous functions, i.e. solids removal and gas

saturation with steam. These process steps could be linked because the

catalyst used in the CO-shift conversion is sulphur-resistant and

operates at relatively low steam-to-gas ratios and low pressures.

Said water scrubbing system has three stages: After passing the hot

water quench, the gas is further purified by removing finest dust

particles in a Venturi scrubber and a water scrubbing tower. The gas

enters the shift conversion at a temperature of about 220 °C and is

adjusted to a C0/H2 ratio of 1:2.4 required for the methanol synthesis.

When the gas has cooled, it is compressed from about 8 to 37 bar and

then piped to a one-step non-selective H2S/C02 scrubber. In this so-

called Recti sol unit, the sulphur compounds are almost completely re­

moved at temperatures below 0 °C using methanol as a solvent. The C02

content is simultaneously reduced to 3 % in order to meet the methanol

synthesis requirements.

The acid gas separated in the H2/C02 scrubber is fed to the sulphur

recovery unit. Due to the low sulphur content in the coal, the total

sulphur content in the acid gas is as low as approximately 1 %. The

installed Stretford plant is capable of regenerating such low sulphur

contents at conversion rates of > 99 % without an upstream concentra­

tion step. Thus the sulphur emissions are extremely low. The H2S is

converted into salable elementary sulphur. A part stream of the puri­

fied C02 is compressed and used for pressurizing the coal, ash and dust

lock-hopper systems.

The coal consumption of the HTW demonstration plant amounts to approx.

30 t/h. This corresponds to about 37 000 m3/h methanol synthesis gas to

specification. The erection of the plant was completed in autumn 1985.

From March 1986 to the end of October 1987, the HTW demonstration plant

was run for about 6 100 hours in seven periods of operation. About

t r

- 177 -

Waste heat boilerGasifier, lock hopper

Coalhopper Dust

refflovalsystems W< ter

si rubbe

Shiftconversi

Ash lock hopper

Synthesis gas

Sulphur

Sulphur recovery

COg/HgS scrubbingWaste water

Ash, dust

1

III

Figure 7. HTW pilot plant - synthesis gas yield and synthesis gas out­put as a function of coal input.

t r—C- —U.. - -J

Figure 8. Demonstration plant Berrenrath.

4834

178 -

130 000 t of dry brown coal were converted. Since the beginning of June

1986, nearly 138 million m3 synthesis gas have been produced.

Table 4 compares the most important test results with the design. The

C-conversion design data of 91 % have been exceeded. Further positive

results are: the oxygen consumption is lower and the specific synthesis

gas yield is higher than the expected figures. This is due to the im­

proved heat balance of the demonstration plant compared with the pilot

plant and to the extended post gasification zone. The lower oxygen

consumption is also substantiated by the design/as-measured comparison

of the raw gas analysis. Owing to the low C02 formation, the product

gas content (CO + H2) in the raw gas is as high as about 75 % by vol.

while 70 % was expected.

The a.m. operating data show that the gasifier reaches or even exceeds

the design figures. Nevertheless, the experience from said plant in the

field of process engineering such as nozzle charging of gasifying

agents, will lead to improved operating conditions. Technical improve­

ments are still required, particularly in the field of solids hand­

ling.

Table 4. Results of the Demoplant.

r t

DESIGN ACTUAL

Synthesis gas CO+Hj m3/h 35 500 35 500

Coal (dry) t/h 30.5 28.0

Raw gas composition VoIX

h2 31.3 36.0

CO 39.2 39.6

COj 25.7 20.1

ch4 2.7 3.6

h2s 0.16 0.05

Carbon conversion X 91 96

Oxygen consumption m3/t coal maf 430 400

- 179 -

The plant units in which the raw gas is converted into synthesis gas

suitable for the methanol synthesis loop have been working more or less

troublefree after solving the usual start-up problems.

The design values for the gas throughput of the sulphur recovery unit

(Stretford process) have not yet been reached. In order to optimize the

Stretford plant operation, a test programme has been prepared and part­

ly been carried out. In close co-operation with engineering firms, we

have elaborated adequate means to improve the performance and said

measures will be implemented on a step-to-step basis.

7 APPLICATION OF THE RHEINBRAUN HTW PROCESS

As has already been mentioned in the previous chapter, the commis­

sioning of the Berrenrath demonstration plant is a significant stage in

the development of the Rheinbraun HTW process. The operational results

of the Wachtberg pilot plant were confirmed to a large extent. Various

interesting process applications will be available.

7.1 Synthesis gas

The primary aim of the process development scheme was to generate syn­

thesis gas for the production of methanol, ammonia, etc. The gas gene­

rated in the Berrenrath demonstration plant is used as feedstock for a

methanol synthesis plant.

It is also possible to use a synthesis gas unit for the production of

lignite-based ammonia. Figure 9 shows the configuration of such an

ammonia plant with a capacity of 1 350 t/d. In this case, the lignite

is first ground, dried to a moisture content of 12 % and then fed to

two parallel gasifiers. A third gasifier is required as standby unit.

Saturated steam of 120 bar is generated in the raw gas cooler. In a

water scrubber, the dust is removed and the gas is simultaenously satu­

rated with water. The major part of the carbon monoxide contained in

the gas is converted into hydrogen in a two-stage CO-shift conversion

]

I .C .]

T f

Wl

180 -

Raw lignite 126 t/h (32.8 % moisture)

NitrogenHP-Steam

Dry lignite

SteamWastewater

Sulfur

Steam

Purge

Hydrogen

Hydrogen

Partialoxidation PSA-unit I

2*1 train

HjS/COj- Wash 1 train

PSA-unit II

Fluidized bed boiler

Waste water treatment

CO-Shift I1 train

CO-shilt II

Claus unit

Air-separat. 1 train

Gasification 2 * 1 train

Drying & grinding

Ammoniasynthesis

Purge gas to FBB 1 350 t/d Ammonia

Figure 9. Ammonia from lignite.

I I

T f

unit using a sulphur-resistant catalyst. In the PSA unit, highly puri­

fied hydrogen is separated and, after admixing nitrogen, it is piped to

the ammonia synthesis plant operating at a pressure of 130 bar.T 7

A significant feature of this configuration is the treatment of the

purge gas which leaves PSA unit I and has a methane content of about

40 *. This methane-rich gas is recompressed and fed to a partial oxida­

tion unit. The product gas produced in this unit is piped to CO-shift

181

conversion unit II and PSA unit II. This permits a considerable in­

crease in the hydrogen recovery rate. Another significant feature is

that the carbon-laden ash is fed to the fluidized-bed combustion

section where it is incinerated with additional lignite in order to

produce high-pressure steam required for driving the large compressors.

The additional electric energy is rather low (approx. 5.7 MWh/h) and is

only required for small plant consumers.

Tests with Finnish peat had also been performed in the Wachtberg pilot

plant. Since the test results were good, Uhde was awarded a contract in

early 1986 by Kemira Oy for the design and construction of a gasifica­

tion plant using Finnish peat as feedstock. The technical data of this

plant are indicated below:

Feedstock rate: 30 t/h peat (dry)

Gasification pressure: 13.5 bar max.

Gasification temperature: 900 °C

Synthesis gas yield: 1.110 kmol C0+H2/h.

This synthesis gas yield permits a production of 240 t/d ammonia.

The peat is first moulded in extruders to form large pellets. The dried

pellets are ground in order to obtain the grain size distribution re­

quired for the gasification process. The synthesis gas produced in this

unit is treated in the existing downstream synthesis gas scrubbing

section (Ami sol. Recti sol, liquid nitrogen scrubber) and fed to the

ammonia synthesis unit. The gasifier dimensions are similar to those of

the gasifier in the Berrenrath demonstration plant.

When considering the Berrenrath demonstration plant as reference unit,

it was necessary to determine the scale-up factor for the envisaged

ammonia synthesis unit with a capacity of 1 400 t/d (which is the

standard size at present) or a methanol plant with a capacity of 2 500

t/d. Commercial facilities using natural gas are already in operation.

Uhde has also designed and constructed such plants. Several gasifiers

are required for this plant size, i.e. three gasifiers for a 2 500 tpd

rw \_ . i

I l

- 182 -

methanol plant. Compared with the Berrenrath gasifier capacity, each of

said gasifiers has double the throughput which can be obtained without

major dimensional modifications to the Berrenrath gasifier. The gasifi­

cation pressure is increased and permits a higher specific gasifier

throughput. In view of the sophisticated technology of the coal gasifi­

cation unit (compared with natural-gas steam reforming), provision

should be made for a reserve capacity to improve the on-stream factor,

for Instance by means of a stand-by gasifier.

7.2 Combined-cycle power plants

The classical configuration of the power plants using fossil fuels is

based on the combustion technology. The combustion heat is utilized for

the generation of steam driving a steam turbine coupled to the gene­

rator. For thermodynamic reasons, the efficiency of such a plant is not

very high. When fuel gas is burned and the flue gas is fed to a gas

turbine directly coupled to the generator, the energy conversion takes

place at a higher temperature and will improve the thermodynamic condi­

tions provided the waste heat of the gas leaving the turbine is re­

covered. The combination of gas turbine and steam turbine is called

combined-cycle power plant. Natural gas-fired combined-cycle power

plants are being successfully operated in many parts of the world.

When coal is used as feedstock, it has first to be gasified. The over­

all efficiency of such an integrated gasification combined-cycle power

plant (IGCC plant) is considerably higher than that of a conventional

solid-fuel power plant, i.e. in excess of 40 %. The efficiency largely

depends

- on the selected type of gas turbine and the inlet temperature,

- the selected type of coal gasification,

- the waste heat recovery system,

- the mode of gas purification.

Uhde and Rheinbraun, in cooperation with Kraftwerk Union, have deve­

loped a concept for a 600 MW combined-cycle power plant on the basis of

the HTW process.

i :i

II

V rL 1

\

- 183 -

Within the scope of this stu<iyf two basic alternatives have been in­

vestigated in detail, i.e. gasification with air or with oxygen. These

results indicate that, in both cases, the overall investment and the

costs for operating the power plant are very similar. This calculation

is based on an assumed location in the Federal Republic of Germany and

the conditions in this country. Hence, the costs may vary if the basic

conditions are different.

Figure 10 shows a block diagram of the plant using air as gasification

agent. The coal is dried to a residual moisture content of 12 %, using

1

i _t :j

T T 1

F

Figure 10. Block flow diagram of a HTW based 600 MW IGCC plant.

- 184 -

back-pressure steam. The dried coal is fed to the gasifier. The gasifi­

cation process takes place in six parallel gasification lines, while an

additional gasifier is required as stand-by unit. In case of an oxygen

blown gasifier four parallel lines would be needed.

The hot raw gas is first fed to a cyclone in which the coarser dust

particles are removed. Then it is cooled in a waste heat boiler, high-

pressure steam being generated simultaneously. In a down-stream heat

exchanger, the raw gas is cooled further with the aid of the purified

gas. The remaining fine dust particles are removed in a wet scrubber.

After passing a further cooling step, the gas is fed to a gas scrubber

in which hydrogen sulphide is separated, thus obtaining a low sulphur

content without difficulty. The acid gas is treated in a Claus plant

which comprises downstream tail gas treatment facilities and a final

incinerator.

The purified gas is reheated in a heat exchanger with the raw gas and

then fed to the combustion chamber of the gas turbine, i.e. two 200 MW

gas turbines with an ISO inlet temperature of 1 200 °C will be instal­

led. The thermal potential of the turbine exhaust gas is utilized in

waste heat boilers generating high-pressure and low-pressure steam. The

high-pressure steam from the raw-gas waste-heat boilers is superheated

in these boilers. Additional high-pressure steam is generated in

boilers with a circulating fluidized-bed combustion system using the

residual char from HTW gasification. The steam turbine capacity amounts

to 300 MW. Thus, the overall power output of the HTW combined-cycle

power plant is 700 MW, the net value being 630 MW.

The maximum overall efficiency (LHV-basis) of this plant configuration

is 46 *. If the gas turbine temperature decreases to 1 000 OC, the

efficiency drops to about 42 %. Compared with a conventional power

plant, the fuel saved at an efficiency rate of 46 % amounts to 35 %

(see Figure 11).

The combined-cycle power plant with integrated coal gasification, the

concept of which is illustrated here, uses proven process technologies

T f

185 -

Ligniteconsumption

GT.LT.(ISO)

1 120°C 1220°C

GT.LT.(ISO)

I UQ't 1 220°C100X

Conventional IGCC oxygen IGCC airpower plant gasification gasification

(LHV) 3SX 43.S 45.6 44.4 46.7

(HHV) 30 X 36.8 38.6 37.6 39.6

Basis: 600 MW HTW gasification Integrated combined cycle system Location: FRG Feedstock: Rhenish lignite

Figure 11. Comparison of power generation efficiencies.

for the gas treatment and gas/steam turbine processes. The technology

of hot-gas desulphurization and hot-dust removal from the hot gas may

become an alternative in the field of gas treatment when these process

steps have been developed further and proved to be reliable. A hot gas

clean-up would further improve the overall efficiency.

In 1986, Uhde signed a contract for an engineering study of a 2 x 300

MW combined-cycle power plant in South Australia using the HTW process

and South Australian brown coal. Preliminary tests have been performed

in the Aachen test facilities with South Australian brown coal, which

has a high salt and sulphur content, and they have shown that this type

- 186 -

of coal is suitable for the HTW process. Further larger-scale tests

have been performed in the HTW pilot plant in Wachtberg. The test

results indicate that it is possible to gasify also problematic coals

without serious problems. At present, a feasibility study is being

elaborated on the basis of the above experience.

Figure 12 shows a model of the envisaged combind-cycle power plant in

South Australia.

Figure 12. Combined cycle power station. South Australia (artist view).

8 ENVIRONMENTAL ASPECTS

Over the years, we have become ever more aware of the fact that some­

thing has to be done to effectively reduce the pollution of the atmo­

sphere. Emission from power stations is of primary interest in this

case. There are stringent requirements in the USA, Japan and the in­

dustrial countries of Western Europe (especially in the Federal Re­

public of Germany) to limit the emissions of sulphur dioxide, nitrogen

oxides, and dust. For a conventional power station this means that a

187 -

desulphurization unit and a DeNox unit have to be added and highly

efficient dust filters have to be installed. The investment costs of

these facilities nowadays constitute 25 % of the total investment costs

for a power station.

In the case of a combined-cycle power station, gaseous waste and

effluents, e.g. the waste water from the gasification plant, must be

taken into consideration. In addition, there is solid waste in the form

of ash and dust.

8.1 Gaseous waste

In a combined-cycle power station, gas purification constitutes an

essential part of gas treatment and is relatively simple. For compa­

rison, here are some figures which make this clear.

In a conventional 600 MW lignite-fired power station, the amount of

flue gas obtained is 2.4 million, mn3 per hour, whereas in a combined-

cycle power station, as outlined before, the amount of raw gas to be

purified is approx. 360 000 mn3 per hour, this being less than 20 % of

the flue gas. This does not take into consideration the fact that, in

the combined-cycle power station, the gas to be treated is under pres­

sure. If the comparison is related to the volume, the ratio is 100:1.

It should be mentioned that the sulphur compounds contained in the gas

are not present in the form of S02 but in the form of hydrogen sul­

phide, which is much easier to remove from the gas and which can be

converted directly to elemental sulphur in a Claus unit. The sulphur

dioxide obtained in the flue gas treatment processes presently being

used in conventional power stations is converted to gypsum (calcium

sulphate). However, it is obvious that this will no longer be possible

in the near future due to disposal problems, considering the amounts of

gypsum obtained.

The formation of nitrogen oxides in the flue gas of a power station can

be traced to two sources. Firstly, nitrogen oxides are formed during

the combustion of nitrogen-bearing compounds contained in the coal and,

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secondly, nitrogen oxides are formed thermally from nitrogen and oxygen

at the high combustion temperature. In a combined-cycle power station,

the nitrogen compounds contained in the coal are degraded during gasi­

fication or completely removed from the gas in the gas treatment unit.

The temperature prevailing in the combustion chamber limits the thermal

formation of nitrogen oxides.

Figure 13 shows a comparison of the emissions from a conventional power

station and a combined-cycle power station. Attention is drawn to the

fact that the conventional power station is equipped with a gas treat­

ment unit as required by the present statutory environmental regula­

tions in the Federal Republic of Germany. As shown in the figure, the

expected emission from the combined-cycle power station is considerably

below the statutory emission limits.

In this connection, the C02 emission must also be mentioned, which, due

to the higher efficiency, is reduced considerably. C02 emission can

cause a dramatic long-term change in climate.

'V'V A STATUTORY LIMITS (FRG) FOR CONVENTIONAL LIGNITE-FIRED POWER “ STATIONS

A COMBINED CYCLE POWER STATION

1i~'“

Figure 13. Emissions for lignite-based power stations.

8.2 Effluents

Waste water is obtained in the coal gasification process. Figure 14

shows the basic configuration of a wet purification unit as used in the

Berrenrath plant mentioned earlier. The treatment consists of three

steps, i.e. quench, Venturi scrubber and scrubbing column. This con­

figuration permits effective removal of the dust from the gas. The

water stream containing solids is fed to the downstream separator

(filter or settling tank). The treated water from the separator can be

returned to the cycle. Depending on the water-soluble constitutents of

the coal, a certain amount of waste water must, however, be withdrawn

from the cycle in order to prevent salt enrichment in the cycle, as this might cause extensive corrosion. The quantity to be withdrawn is

thus primarily determined by the salt content of the coal. It goes

without saying that the waste water stream also contains gases which

are dissolved in the water, in particular C02, H2S and NH3, but also

water-soluble organic constituents whose point of condensation is below

the temperature of the water cycle. The dissoTved gases are removed 'n

Scrubber Clean gas

QuenchFresh waterVenturi

Raw gasSeparator

Sludge

Dissolved gas

Stripper

Waste waterRecycle

Figure 14. Dust removal and water scrubbing system.

190 -

a down-stream stripper and fed to the sour gas treatment unit. A sepa­

rate treatment step may be required in order to recover the valuable

substances, such as ammonia, in concentrated form from the waste water.

The water stream leaving the stripper is of significance with a view to

environmental pollution. It contains traces of inorganic and organic

substances. The salt content is determined to a large extent by the

composition of the mineral constitutents of the coal. The content of

organic substances depends on the gasification process used. In the HTW

process as well, where the gasification temperature is relatively low

(900 - 1 000 °C), a considerable decrease in hydrocarbons is achieved

by the special two-step process in the gasifier. The content of aroma­

tic hydrocarbons is in the ppm range only. The waste water can thus be

discharged directly into the general waste water system without any

further treatment or conveyed to a biological waste water plant, de­

pending on the applicable local regulations.

In this connection, it must be mentioned that the cooling water re­

quirement of a combined-cycle power station is considerably lower than

that of a conventional power station. This is not only a consequence of

the higher efficiency, which indicates that less heat is dissipated to

the environment, i.e. to the atmosphere and to the cooling water. The

saying in cooling water is also due to the fact that only part of the

total power output is made up by the steam turbines, which again re­

sults in a decrease of the cooling water consumption. Referred to the

entire plant, about 50 * less cooling water is required.

8.3 Solid waste

The solids leaving the power station consist of ash and dust. In the

HTW process, which operates at temperatures below the ash melting

point, there is a risk that, particularly in the case of highly basic

ash, part of the sulphur contained in the coal is combined in the form

of sulphide as a result of the reducing conditions in the gasifier.

There is, of course, the risk that these sulphides may be released

whenever there is any contact with water. The concept therefore

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provides for final combustion of the ash and dust. Apart from the addi­

tional energy that is generated, sulphide compounds are converted into

sulphates in the combustion process in the presence of an excess of

oxygen. The ash obtained after combustion can then be dumped under the

same conditions as the ash from a corresponding lignite-fired power

station.

In sumnfery, it can be said that environmental pollution is drastically

reduced by a combined-cycle power station as compared with all other

types of fossil-fuel-fired power station.

9 COMBINED-CYCLE POWER STATION WITH CO-PRODUCTION

The main distadvantage of electric power generation is the fact that

electricity cannot be stored. Compared with a chemical plant, the load

factor of a power station is relatively low. This particularly applies

to power stations which operate in the medium-load range. In this case,

there are interesting possibilities for varying the concept of a comb­

ined-cycle power station.

These variants are based on the assumption that the gasification

section including the gas treatment unit continues to operate at maxi­

mum load and that the generated gas is used for the production of, for

instance, ammonia or methanol. However, it then becomes necessary to

provide a further gas treatment stage in order to adapt the generated

raw gas to the conditions prevailing in the synthesis unit. Moreover,

it is necessary to design the gasification section for a certain excess

capacity in order to permit continuous operation of the chemical plant

even during maximum load operation of the power station, thus reducing

the problem of capacity adaptation.

Conversion to methanol is of special interest since it can easily be

stored and burnt as additional fuel in the turbine if required for

peak-load operation. In this special case, no specific demands are made

on the methanol produced and the synthesis unit can be kept relatively

simple. Distillation of the methanol is not necessary.

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10 ECONOMY

A number of investigations performed by us have shown that the invest­

ment costs for a combined-cycle power station are identical with those

for a conventional power station. In view of the higher efficiency, of a

combined-cycle power station, it is to be expected that the operating

costs will be appreciatively lower. This, of course, depends on the.

energy production costs, i.e. the price per kWh will therefore be

lower. An exact calculation will only be possible when comparing alter­

natives for the same location.

11 SUMMARY AND PERSPECTIVES

By the development of the Rheinbraun HTW process in three steps

- PDU-plant at Aachen,

- pilot plant in Wachtberg,

- demonstration plant at Brrenrath,

a large field of investigation was covered:

- comprehensive investigation of process principles was performed,

- a number of different feedstocks was tested,

- the influence of various process parameters on the gasifier per­

formance was investigated over several years,

- commercial experience from the operation of the Berrenrath plant

was gained.

The results obtained will give a reliable and proven basis for design

of large plants. The process can be used for different applications:

The prod'iction of synthesis gas to be used

- in the chemical industry or

- for the production of fuel gas for a combined cycle power

station.

The concept of such power station with integrated coal gasification

does show not only a high efficiency but also drastical reduced emis­

sions.

Wp are sure that this concept will be the basis for plants to be built

in the future.

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THE PRESSURIZED BED PEAT GASIFICATION PLANT OF KEMIRA OY IN OULU

Kalevi Kuorelahti, Manager of Energy Unit Kemira Oy Oulu Works Box 171SF-90101 Oulu, Finland

Abstract

Kemira, the largest chemical enterprise in Finland, is building a pressurized bed peat gasification for producing synthesis gas. The gas will be used for ammonia and other chemicals. The ammonia production will be 80,000 tons/year. The investment costs amount to 225 million FIM. The production is to be started in spring 1988.

1 HOW TO GASIFY PEAT

For the time being Kemira operates in Oulu a minor ammonia plant based on the gasification of heavy fuel oil. The ammonia production in Oulu was started in 1952 on the basis of coal gasification. In order to establish an ammonia production resting on domestic raw material a research program aiming at the gasification of peat was started in the middle of the 1970's. Peat is the only realistic domestic raw material for synthesis gas in Finland.

1.1 Gasification Process

In chemical and physical properties peat resembles lignite. That's why the experimental work was started at a research institute for lignite gasification in Aachen. However, the treatment of peat proved to be quite problematic.

The volume weight of peat is very small ranging from 0.2 kg/m^ to 0.45 kg/m^. The moisture content is to 40 - 60 % and the type, composition

and ash content of peat vary very much. Moreover, the Finnish peat contains fibres, wood and stones. Due to these factors peat is difficult to treat with the feeding equipment of a gasifier.

194 -

In order to study the feasibility of peat for gasification and to build a suitable gasifier for this purpose, tests with crushed peat pellets were made with the German HTW-gasifier in Berrenrath near Cologne as well as with the Swedish Mino-process in Sweden. The American IGT-process was tested using dried and crushed sod peat. On the basis of these test runs the HTW fluidized bed gasifier was chosen.

1.2 Peat pretreatment studies

Because either milled peat or sod peat dried by sun and drawn from the peat production were not without some processing suitable to be fed into the gasifier, the study of alternative pretreatment processes was started in cooperation with The Finnish Technical Research Institute.

The drying, pelletizing and crushing of milled peat resulted in good crumbs as to physical properties, but the treatment had an injurious effect on the reactivity. Moreover, the pelletizing process is not very easy to perform on a large scale.

The crushing and drying of sod peat proved to be a more interesting alternative. The reactivity of the crumbs is clearly better, even if the physical properties do not match crushed pellets. The moisture content of sod peat is 30 - 40 %, that of milled peat instead 45 - 55 %. This is an important advantage in view of the drying.

That's why the study was concentrated on the crushing technology and drying equipment as well as on the research of sod peat properties in order to master these pretreatment processes as well as possible from the very beginning.

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2 PEAT GASIFICATION PLANT

The gasification plant being under construction includes as intergrated individual processes: peat transfer, screening, crushing, drying, pressurized HTW fluidized bed gasification, soot removal, raw gas compression and three-stage gas purification. The existing Pyroflow boiler plant serves for energy supply and waste incinerator for the gasification.

195 -

2.1 Conveyors and screening equipment

Sod peat is transported to the plant by lorries. The loads are dipped into unloading pockets, from which peat is passed by belt conveyors to the screening. A two-deck disc screen separates the oversized wood material as well as the fine peat powder, which are led into the power plant to be burnt. The screened sod peat is carried into the 2,000 m^

storage silo of the gasifier.

2.2 Crushing

From the storage silo sod peat is fed at a rate of 42 t/h to the crushing. Crushing takes place in a three-screw biomass crusher with a bottom screen. The granulometric composition of the material coming from the crushing is 0.5 - 10 mm. After that the crushed product having a moisture content of 30 - 40 % is passed into the feeding silo of the drier.

2.3 Drying

Peat must be dried to a moisture content of 15 %. This is made by a counter-pressure steam drier designed by Modo Chemetics. Crushed peat is fed by means of a lock hopper and a screw into the drier, where superheated steam is flowing at a pressure of 4 bar. Steam with peat flows through a set of heat exchangers. Additional steam at a pressure of 11 bar is introduced into the heat exchangers. This evaporates further the water in peat and superheats the steam. The dried peat is then separated by a cyclone and passed into the two feeding silos of the gasifier. From the cyclone part of steam is recycled to the inlet of the drier, the rest to the heat exchangers to be condensed by condensing waters of the plant. The resulting steam of 2,5 bar can be used in the steam system, which makes the plant thermal economy quite profitable. Uncondensed smell gases resulting from drying are burnt at the power plapt.

2.4 Gasification

Dried peat is fed through two lock hopper feeding lines and screw feeders into the fluidized bed reactor having a pressure of 10 bar. Oxygen and 35 bar water vapour serving as reaction and fluidization gas are introduced into the reactor through nozzle systems.

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The result is a gas mixture containing 70 - 90 % of utility gas (H2 + CO) and various contaminants, e.g. CO2, CH4, H2S, NH3, HCN, CgHg, H2

etc. Ash is removed from the bottom of the reactor and with two dust cyclones from the gas flow. Raw gas is cooled with gas coolers generating steam of 2.5 and 35 bar. The carbon residual in gasifier ash is burnt in the dust recycle reactor of the power plant.

2.5 Soot removal

Soot is removed from raw gas by a scrubber. Precipitate is separated by a centrifuge, the ammonia dissolved in water distilled and passed into the power plant to be burnt in the dust recycle reactor.

2.6 Compression and conversion

The raw gas is pressurized to 35 bar by means of a compressor and passed to the CO-conversion, where CO reacts with water vapour to carbon dioxide (CO2) and hydrogen (H2).

2.7 Amisol wash

The amisol wash removes carbon dioxide (CO2) and hydrogen sulphide (H2S). The hydrogen sulphide gas is led into the dust recycle reactor to be burnt. Part of the produced sulphur dioxide (SO2) is absorbed in lime and the resulting gypsum mixed up with ash. Moreover, the amisol wash produces benzene, which is burnt as well.

2.8 Methanol and nitrogen washes

Methane (CH4) and carbon monoxide (CO) are removed by methanol and nitrogen washes. The result is pure hydrogen gas (H2) for the synthesis. The discharge gas is used as fuel gas at the power plant.

2.9 Power plant

The power plant has a 55 MW steam boiler fired by the dust recycle reactor and a 55 MW gas boiler, which generate steam and counter pressure electricity for the needs of the plant.

The emissions of gasification, e.g. ash, ammonia, hydrogen sulphide, hydro cyanic acid, benzene, naphthalene and smell gases, which may be iangerous to the environment, are burnt in the dust recycle reactor.

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The resulting sulphur dioxide is partly absorbed in lime and mixed up with ash. The discharge gas containing methane and carbon monoxide is burnt in the gas boiler.

Moreover, the water for the district heating is used for cooling in the gasification unit, which improves the thermal efficiency of the plant.

3 SUMMARY

The gasification plant being under construction will give us interesting operation experiences about most components of a complexed cycle power station, i.e.:

- Treatment of fuel- Drying of biomass- Pressurized oxygen operated fluidized bed gasification- Cold gas purification- Steam boiler fired by a dust recycle reactor- Steam turbine process- Burning of remaining coke of gasification- S0X, and C0X removal as well as burning of smell gases

In fact, the only missing individual process is the gas turbine process.

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TECHNO-ECONOMIC EVALUATION OF DIFFERENT GASIFICATION COMBINED CYCLE PROCESSES

YrJB Solantausta, M.Sc., Senior Research Scientist Technical Research Centre o-f Finland 02150 Espoo, Finland

Abstract

The economic and technical potential o-f Integrated Gasi-fication Combined Cycle (IGCC) with peat was studied in order to

— compare presently available IGCC to conventional power production, and

- to assess the -future potential of a developed IGCC.

The features typical of peat-based IGCC were assessed. The effects of the most significant process variables on the overall efficiency were studied. Three process concepts were compiled and their material and energy balances were calculated.

Investment and operating costs were estimated, and power production costs for condensing power at 150 MW were compared. The production price of electricity with the present IGCC is competitive with that of conventional power plants. However, its efficiency is higher, 43 X compared to 37 7. (LHV-based). The emissions'from the IGCC are only a fraction of those from a pulverized fuel boiler. The future IGCC concept proved to be very promising. The efficiency would be as high as 50 X and the production cost of electricity close to FIM 150/kWh, compared to FIM IBO/kWh for the conventional technology.

- 200 -

1 INTRODUCTION

The steady increase in the electricity consumption of Finland necessitates decisions concerning the construction of several condensing power plants. At the same time more attention is paid to the environmental effects of the plants. Although one peat based condensing power plant is under construction in Haapavesi, it is generally accepted that without the energy tax reform peat could not compete with coal as a fuel for condensing power production /!/.

During this decade power production from coal with Integrated Gasification Combined Cycle (IGCC) has been maturing to become an increasingly promising technology for the future /2/, /3/. While synthesis gas production from peat has been developing towards a commercial reality /4/, a peat based IGCC is also becoming a technical alternative. A cooperation work to study pressurized power production processes experimentally was started in Finland during 1986, when IVO Oy, HTKK and VTT, together with the Ministry of Trade and Industry, initiated an experimental program.

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However, no economic studies are available concerning the competitiveness of a peat-fired IGCC compared to the more conventional schemes. Power production cost estimates areneeded to compare the new technology to the old one, the W

L.technical and economic potential of peat IGCC has to be ^assessed, and finally experimental work has to be supportedby economic studies. This paper summarizes the results ofthe first part of a techno-economic study that is underwayat VTT. Power production from peat with condensing IGCC isiiscussed in this paper.

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2 THE APPLICATION OF IBCC FOR PEAT

2.1 Bamis for tha assessment

After the process concepts were compiled and the status of technology was assessed, material and energy balances were calculated with a spreadsheet program. The effects of some process parameters were studied (chapter 2.3). The investment and operating costs were estimated based on the data given by major unit process licensors and companies utilizing the present technology. Minor details had to be estimated from published literature, especially /!/. The aim was to compile the power production cost so that they could be compared to the numbers of a previous study of Ekono Oy /!/. The accuracy of the investment for the case utilizing oxygen blown gasifier (present technology) is estimated to be within • 20 X. There is more uncertainty in the third concept employing future technology.

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2.2 Process concepts

Three process concepts were compiled. The first one employs an oxygen blown HTW gasifier with cold gas clean-up (water quench with further physical gas scrubbing) and a gas turbine with a firing temperature of 1100 *C. This concept is referred as IGCC-present. The second case utilizes an air blown fluid-bed gasifier with water quench and chemical gas scrubbing (IGCC-air). The third concept is based on an air blown fluid-bed gasifier with in bed desulfurization and hot gas clean-up. The firing temperature was 1200 *C (IGCC-future). The block flow diagram of the first concept is presented in figure 1.

202 -

SEPARATION

PEATDRVMG

OXYGEN U [COMPRESSOR

GASnCATION

20 2)s 20 22 24WHO

QUENCH GAS SULFURCLEANMG RECOVERY SULPUR

30 MAIR

COMPRESSOROAS

TURBME

HRSG

~l----

FLUE

STEAMTURBME

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Figure 1. The block flow diagram of the peat-fired IGCCW r

?.3 The effect of eome parameters on the efficiency of the IGCC

The most important single factor typical of peat delivered to the plant is the relatively high moisture content» usually about 50 X. The effect of the moisture content on the plant overall thermal efficiency is shown in figure 2 (IGCC-present).

Gasification pressure is one of the key process parameters. Figure 3 depicts the efficiency of the IGCC-present for gasification pressures ranging 1.0-3.5 MPa. A similar presentation can be used together with other factors (gas turbine characteristics, investment and operating costs etc.) to find the optimum gasification pressure. In the present case it should be around 2.0 MPa.

203 -

MOISTirnC CONTENT OF PEAT * O HHV BASED * LHV BASED

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Figure 2. The ef-fect of the moisture content on the plant overall thermal efficiency. r r i- -A

0.1?

0.17 -

0.16

0.14

PRESSURE OF GASIFICATION MPa

w

. The effect of the gasification pressure on the plant internal power consumption.

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Figure 3

- 204 -

The pressurized fuel gas can be efficiently washed in commercial unit processes to exclude nearly all sulfur compounds. The firing of the fuel can be controlled in the turbine to suppress the generation of thermal NO„. Since the gas turbine does not tolerate solids, the flue gases are extremely clean and the emissions of the plant are order of magnitude less than with conventional power production. This is shown in figure 4.

The efficiency of the power plant can be increased by investing on heat transfer equipment. One example of such measures is presented in figure 5. The neat exchanger a^ea is depicted as a ■‘unction of the preheat temoe-'at’.re of the fuel gas (HX in figure i>. The same increase ." tne preheat temperature will increase the efficiency of the plant from 43.6 to 43.9 7.. Again the optimum temperature can be found bv taking into account, investment and fuel costs.

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1

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rr~l PULVERIZED COMB. ____ hTTI BC»Fi.UE ®AS Cl.EAN.BSlttCC-PBESENT

-'igure 4. Comparison between emissions of a conventionalpeat power plant, the same plant with DeNO„ and flue gas desulfurization, and IGCC-present.

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PREHEAT TFMFERATURE OF THE FUEL &AS OC HAS ♦ 6790 kmol/h (.NTP)

Figure 5. The e-f-fect o-f the -fuel gas preheater temperature on the heat exchanger area.

One o-f the key -features of the gas turbine technology under development is the allowed firing temperature. It is generally accepted that the firing temperature of 1100 *C should be available by the beginning of next decade. The effect of the temperature on the plant efficiency is shown in figure 6. For the future concept in this study the firing temperature of 1200 *C was selected.

3 MATERIAL AND ENERGY BALANCES OF THE CONCEPTS

A summary of balances of the concepts is presented in table1. The overall efficiency of power production with IGCC is estimated to range from 42 to 90 X based on lower heats of combustion. This can be compared to the efficiency of 37 X of a pulverized boiler plant.

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43.5

43.0

42 0

41 .9

41.0--------

110010GC

GT PIPING TfMPEBA*nip£ 0C

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Figuf# h. The e-f-fect o-f the -firing temperature on the plant e-f-f iciency. r r

Table 1. A summary o-f process balances.

GASIFICATION OXYGEN AIR AIRBLOMN BLOMN BLOMN

GAS CLEANING COLD COLD HOTIN

PEAT t/h 160 160 160LHV MM 391 391 391HHV MM 462 482 482

OUT

P0MER MM 168 164 193

EFFICIENCYLHV X 43.1 42.2 49.5HHV X 34.7 34.0 39.9

207 -

vnergy flows of two of the concepts <IGCC-air and IGCC future) are compared in figure 7. The hot gas clean up (which is carried out at 650 *C) considerably increases the power produced by the gas turbine and thus the efficiency T"

_u.of the plant.

A summary of the energy flows of a plant consuming 160 t/h of 50 % moist peat is presented in table 2.

I

DRYING

100.0100.0

51.4

POWER

STEAMTURBINE

STEAMTURBINE

GASIFI­CATION

WASTE HEAT BOILER

GASTURBINE

DRYING

HRSG

GASCLEANING

GASIFI­CATION

WASTE HEAT BOILER

GASTURBINE

HRSG

T r

PRESENT FUTURE

Figure 7. Relative energy flows of process concepts.

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Table 2. A summary o-f the energy -flows of a IGCC-present. ^ Ambient temperature 15 *C, reference temperature

25 *C.

CHEM.ENERGY(LHV)

SENSIBLEHEAT

POWER TOTAL

IN MWPEAT 391 -1 389OXYGEN PLANT AIR -1 -1GT AIR -3 -3INTERNAL POWER CONS. 24 24

TOTAL 391 -5 24 409OUT MW

GT POWER 12B 128STEAM TURBINE POWER 64 64FLUE GAS 45 45STEAM CONDENSING 100 100LOSSES IN GASIFICATION 12 12OTHER LOSSES 61 61

TOTAL 218 192 409

T

4 PRODUCTION COST OF ELECTRICITY

The production cost of electricity was estimated based on operating and investment cost estimates. Capital costs were valued with the annuity method. As stated earlier, the power production cost was assessed to be comparable with those published by Ekono Oy /!/. Therefore the rate of interest employed was 5 X and service life 25 years. For the economic comparison the capacity of the plants was scaled down to the capacity of the above study /!/, that is 150 MW* of condensing power. The capacity factor employed was 74 X corresponding to an annual peak operating time of 6500 h. The price of peat was 39 mk/flWh.

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T*bl» 3 is a summary o-f the power production costs of the three concepts. Investment tax credit was taken into account when the total capital requirement was defined.

Table 3. A summary of the power production costs.

IGCC-PRESENT

IGCC—AIR IGCC-FUTURE

PRODUCTION MW. 150 150 150PEAT CONSUMPTION t/h 143.3 146.2 124.5TOTAL CAPITALREQUIREMENT mk/MWh 4770 4760 3820

COSTSFIXED

Mmk/a35.2 35.2 33.2

VARIABLE 85.3 87.1 74.7CAPITAL 50.8 50.8 40.6

AVERAGE COST OF PRODUCTION

POWERmk/MWh 176 177 152

Power production costs with IGCC future and present cases are compared with conventional pulverized boiler technology in figure 8. The hot gas clean-up scheme offers a substantial decrease in production costs: both the capital and fuel costs are estimated to be considerably less than for both the IGCC-present and the conventional plant. A further conclusion is that the IGCC-present already competes well with the technology presently used in condensing power plants.

The future potential of IGCC is studied in figures 9 and 10. Figure 9 depicts how much the investment cost of IGCC-future is allowed to increase before power production with IGCC would be more expensive than with conventional technology. An increase of more than 50 % would yield that situation.

210 -

- '/ // /s// // //'-t'/y/A

w \\ N \\

IOCC-FUTUREl&CC-PBESENTPULVERIZED COMBUSTION

FUEL COST ] FIXED COST

ZD CAPITAL COST OTHER VARIABLE

Figure 8. Production costs of electricity condensing power jpplants 150 MW». Interest rate 5 %, service life 25 a, peat 39 mi./MWh.

s?it■

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4270 mk/kW

170 -

160 •

193 •

ICCC-PLANT INVESTMENT « OF BASE VALUE O IOCC—FUTURE ♦ PULVERIZED COMB.

• IOO«>3BOO mk/kW

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figure 9. IGCC-plant investment estimate, effect of variation on power production cost.

- 211 -

Figur* 10 dapicts th» ••ffact o-f th» 16CC •f-ficiancy on powar production cost. The estimated efficiency of close to SO X of the IGCC-future concept could decrease to below 40 X before the production cost would exceed those of conventional technology.

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EFFICIENCY OF THE ICCC-FUTUBE CLHVJ «* O PULVERIZED COMB. + ICCC

Figure lO.Efficiency of the future IGCC, effect on power production cost.

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5 ACKNOWLEDGMENT

The author would like to thank Ms. Virve Tulenheimo and Mr. Esa Kurkela for their help in preparing this work.

212 -

REFERENCES

1. Kotimaisten polttoaineiden kiIpailukyky 1986. Kauppa- Ja taol1isuusministeriB, anergiaosasto, sarja Bs62. Helsinki 1986.

2. Clark, W. at al. Cool Water: Per-formance & economics. AIChE National Meeting, Boston, MA, August 24-27, 1986.

3. Leibson, 1. Comparative Economics for Alternative Fuels and Power Technologies, Energy Progress 7(1987)2,63-71.

4. Koljonen, J. & SBderwall, T. Turve synteesikaasun raaka-aineena. Kemia-Kemi 14(1987)10, p.829-832.

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HOT GAS CLEANUP: THE STATE-OF-THE-ART AND PROSPECTS FOR THE FUTURE

Roland Clift, Prof.

University of Surrey

Department of Chemical and Process Engineering

Guildford, Surrey, U.K.

Abstract

In addition to sulphur compounds, turbine expansion gases in PFBC and

gasification cycles contain particulates, alkali salts and HC1. Gas

cleaning is most economic if carried out hot. During periods of steady

operation, cyclones can reduce particle loading in PFBC plant to a

level acceptable to a gas turbine. However, periodic upsets in cyclone

performance could destroy a turbine. Moreover, cyclone performance in

gasification plant is less well investigated. Therefore, in either

application, some form of tertiary gas cleanup is needed. Wet scrubbing

is not feasible, and sonic agglomeration is questionable. Electrostatic

precipitation is potentially usable but unproven for PfBC, and very

doubtful for gasification. Metal and ceramic fabric filters are not

sufficiently durable. Rigid ceramic barrier filters appear to be the

best developed option for tertiary cleaning, either as candles or

crossflow monoliths. Granular bed filters need further evaluation, but

offer potential advantages of steady operation and compact design, with

the possibility of collecting gases and particulates in one device. The

extent to which alkali salts represent a problem depends on chlorine

levels in the coal: high-chlorine coals are subject to more alkali

release, and the alkalis are harder to remove by "gettering".

1 INTRODUCTION

This review paper presents an overview of the various possible techno­

logies for cleaning gases derived from coal by combustion or gasifica­

tion. Most of the processes and cycles currently under development

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involve expanding the gases through a turbine. In Pressurised Fluidised

Bed Combustion cycles, turbine entry conditions are typically 10 to 20

bar and 1 100 to 1 150 K (825 to 875 °C). Gasification cycles are more

variable in their turbine entry conditions, but 10 to 20 bar and 1 170

to 1 270 K (900 to 1 000 °C) are typical /!/. In view of the Increased

overall efficiencies available if turbine inlet temperatures can be

increased, there is an incentive to go to higher temperatures still.

This is possible with some gasification cycles, and may become feasible

with the introduction of new materials, such as ceramic composites, for

construction of turbomachinery. In any of these processes, there is a

clear economic advantage in cleaning the gases hot rather than cooling,

cleaning, and reheating /2/.

Four general classes of harmful contaminants are present:

a) particulates: fly ash in a combustion process; coal char with

various ash and tar content (here termed "gasification fines" in

gasification;

b) sulphur compounds: oxides and sulphates for a combustion cycle,

sulphides in gasification;

c) alkali compounds: predominantly sulphates in combustion and

chlorides in gasification/3/;

d) hydrogen chloride: probably significant in gasification, and in

combustion of fuels high in chlorine /4/.

This review is concerned primarily with particulates. Sulphur compounds

are a matter more of environmental than of process concern, and are not

covered here. The problems associated with alkalis and hydrogen

chloride are closely associated /4/, and will be considered briefly.

In addition to the difference in turbine inlet temperatures, there are

other significant differences between the gas cleaning problems which

arise in combustion and gasification processes. PfBC fly ash is typi­

cally highly cohesive and moderately resistive, whereas gasifier fines

are highly conductive and generally uncohesive /5/. These differences

ensure that the two types of particle behave quite differently in any

collection process. Moreover, the four classes of contaminant Interact.

r r

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- 215 -

Interparticle forces are affected by sulphur and alkali compounds.

Hydrogen chloride concentration has a strong influence on the pro­

portion of the alkalis present as condensed sulphates /3, 4/, and thus

influences both the cohesion of the particles and the extent to which

the alkalis are removed by collecting the particles and any aerosol

present. For these reasons, "modelH studies using redispersed particles

or simulated gases can be misleading, and tests of gas cleaning equip­

ment on pilot- or full-scale plant are essential.

It is also worth remembering that the viscosity of a gas mixture in­

creases with temperature but, to a first approximation, is independent

of pressure. Because we are concerned with removal of fine particles,

Reynolds numbers are very low and so fluid-particle drag is determined

by viscosity not gas density /6/. Therefore, in laboratory or pilot

testing of particle arresting devices, it is more important to match

the process temperature than the pressure.

Rubow et al. /7/ have carried out a detailed technical and economic

comparison of ten particulate collection devices for PFBC. This review

draws on that study, but is concerned with fundamental and technical

considerations rather than economics, with some different gas cleaning

devices, and with gasification as well as combustion. 2

I

I T. ]

2 INERTIAL SEPARATORS

2.1 Reverse flow cyclones

Cyclones are the commonest form of inertial device for recovering

particles from gases. In general, the performance of a cyclone is sa­

tisfactory for coarse particles but deteriorates very rapidly when

particle size decreases below about 10 jim /48/. Because they are com­

pact, simple to construct and operate, and generally capable of hand­

ling high throughputs, cyclones are the most suitable inertial separa­

tors for use at high temperatures and pressures. They are therefore

natural contenders for the first stage or stages of particle collection

216 -

and have been investigated for this type of duty since early work on

coal-fired turbines for locomotive propulsion /see 9/.

In general, at low solid loadings the efficiency of collection (n) of a

particle of diameter dp and pp by a cyclone of diameter D follows, from

simple dimensional analysis /10/, as:

n

d2p Qf"-

pD3 pD(1)

where Q is the volumetric throughput of gas of density p^ and visco­

sity p. The functional relationship in equation (1) depends on the geo­

metrical design of the cyclone. The first independent variable is a

form of particle Stokes number (St) while the second can be regarded as

the cyclone Reynolds number (Re); thus St is a dimensionless measure of

particle inertia, while Re relates to the gas flow within the cyclone.

The more significant variable is St. Because of the increase of p with

T, St for given dp decreases as the temperature increases; i.e. cy­

clonic separation becomes less effective. Using a more mechanisfic

analysis, derived from the worm of Muschelknautz /ll/ and Mothes and

Loffler /l2/, Abrahamson and Allen /13/ have shown that high-tempera-

ture cyclone performance is predictable from results at conditions

closer to ambient, so that there is no obvious difficulty in selecting

or designing cyclones for PfBC or gasification plant. This conclusion

is confirmed by the work of Wheel don et al. /14/, who reported on tests

at the Grimethorpe PfBC facility with cyclones similar but not identi­

cal to the standard high-efficiency Stairmand design /8/. Much of the

earlier work on cyclone performance at high temperatures and pressures

/see e.g. 9/ had been obscured by problems in measuring particle char­

acteristics and cyclone efficiency. Thus the work of Abrahamson and

Allen and of Wheeldon et al. provides a welcome, even if expected,

confirmation of familiar design approaches.

The general overall conclusion of cyclone testing on PFBC duty is that

two or three stages of cyclone can reduce particle loading to a level

acceptable to a gas turbine, but only provided that the operation

r it^. -i.

i r

f t

- 217 -

remains steady. Periodic upsets, caused by pidleg blockage for example,

can generate particle loadings sufficient to wreck a turbine in a few

minutes. Furthermore, in gasification plant where the higher turbine

inlet temperatures demand lower particles loadings, the suitability of ^JL' _i

cyclones remains to be verified. In either type of system, therefore,

some form of tertiary cleaning is essential. Possible approaches are

discussed below.

2.2 Other cyclone designs

Various attempts to improve the performance of gas cyclones have also

been reported. Syred and his colleagues at University College Cardiff

have tested a range of different geometries, one of which is shown in

Figure 1. The Cardiff cyclones feature "Vortex Collector Pockets" (VCP)

which remove part of the dust burden, and also use a "Vortex Weir" to

ensure smooth gas entry into the main cell (which is essentially a

Stairmand cyclone). It is claimed /e.g. 9) that these design features

III

f r i

FJ

Figure 1. Cardiff cyclone separator.

- 218 -

improve the collection efficiency for smaller particles. However, it

could turn out that the main advantages of this more complex design are

reduced height and pressure drop for any given duty.

The General Electric Co. has investigated two other variants /e.g. 16,

17/. The "gas-shielded" cyclone uses an annular vortex "shield" gas

flow between the dirty gas inlet and the exit tube, to reduce entrain­

ment of dirty gas and turbulent radial dispersion of particles. It is

far from clear that the increased complexity in design and operation

results in significant improvement in performance. Another variant is

the "electrocyclone", which has an electrode mounted along the axis of

the cyclone. When a potential is applied to this electrode a corona

discharge is set up, while an electric field is established between the

electrode and the cyclone body. The device therefore attempts to com­

bine by inertial separation with electrostatic precipitation: particles

are charged by the corona, and migrate towards the walls in the elec­

tric field. However, any resulting gain in efficiency is offset by the

reduction in efficiency of cyclonic removal caused by the presence of

the central electrode. The electrocyclone was not recommended by Rubow

/7/ for further investigation.

Further possible variants include cyclones in which part of the device

is driven to enchange rotation of suppress turbulence /18, 19/. There

is little indication that the increase in complexity, which is enormous

for operation at high temperature and pressure, is justified by im­

provement in performance.

Rotary flow cyclones use axial stationary vanes rather than tangential

inlet to impart "swirl" to the gas: see Figure 2. Fairly extensive

testing /e.g. 20, 21/ gives no indication that, for high temperature

high pressure duty, this type of cyclone has any advantage over conven­

tional reverse-flow cyclones.

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\

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EXHAUST (CLEAN GAS)

STATIONARY VANES 'GIVE AIR STREAA. HIGH CENTRIFUGAL ACTION

SECONDARY GAS INLET

SECONDARY AIRFLOW CREATES DOWNWARD SPIRAL OF OUST AND PROJECTS OUTER WALLS FROM ABRASION

-STATIONARY SPINNER

SEPARATED DUST ISDEPOSITED IN HOPPER

PRIMARY GAS DUST INLET *

CLEAN GAS SECONDARY GAS

---- DUST

]

I E l

t r i4- .'t

Figure 2. Aerodyne rotary flow cyclone ischematic).

2.3 Detection of upsets

Thus the need for tertiary cleaning remains, at least as "back-up"

during cyclone maloperation. An alternative approach might be to detect

"surges" in particle loading on-line, and then to take appropriate

action to bypass the turbines or otherwise to protect them. To do this,

it is necessary to measure particle size and concentration on-line and

rapidly, which in practice means directly without sampling. Tate et al.

/22/ have proposed a technique to achieve this, and reported on preli­

minary development of such a device. The principle is to use a "bluff

body" inserted into the gas flow (see Figure 3). The gas velocity on

the axial streamline, u, decreases to zero at the stagnation point.

However, the velocity of a particle, v, decays less rapidly, and the

difference between u and v gives a direct measure of the inertia of the

particle. Figure 4 shows the general configuration of an instrument to

220 -

r r :Figure 3. Gas and particle motion approaching stagnation point of a

bluff body.

exploit this principle. An optic fibre, carrying laser light, points

upstream at the stagnation point. Incoming particles "backscatter" the ▼ Tlight, which is collected and transmitted back by the fibre. The

OPTICAL FORE \ LENS

HCROCONPUTER

DOPPLER BURST COUNTER

FIBRE OPTIC DOPPLER ANEMOMETER

NUMBER DISTRIBUTION OF AERODYNAMIC DIAMETER

\

r-igu'-e 4. In-duct particle sizer (schematic).

- 221 -

Doppler shift of this returning light gives the particle velocity. The

frequencies of many "Doppler-bursts" are measured and the individual

measurements are used to build up the particle size distribution. Be­

cause the technique uses Doppler frequency and not light intensity, it

is free from the obscuration problems which bias diffraction or scat­

tering measurements in "dirty" process environments. Because the sens­

ing head is connected to the delicate photodiode and electronic compo­

nents only by a length of optic fibre, this technique is well suited to

on-line applications such as detecting cyclone upsets.

3 TERTIARY PARTICLE REMOVAL

Variants on most of the conventional gas cleaning techniques have been

tested for PFBC and gasification duty. Wet scrubbing can be eliminated

immediately, for lack of a suitable scrubbing liquid. Molten tin and

molten salts have been considered and tested /e.g. 23, 24/, but are

ruled out in practice and no alternative liquid has been proposed.

Sonic agglomeration reappears periodically as a way of pretreating

particle-laden gases to enhance the efficiency of more conventional

devices such as cyclones. While significant changes in particle size

distribution have been achieved in bench-scale tests at ambient condi­

tions /see 7/, there are major problems in both the science and the

technology of applying sonic agglomeration on a large scale at high

temperature and pressure.

I T ']

w

I I

3.1 Electrostatic precipitation

From the information then available, Rubow et al. /7/ identified elec­

trostatic precipitation as the most attractive tertiary cleaning tech­

nique. Subsequent experience has not so far lived up to this expecta-

tion, partly but not entirely as a result of operating problems in the ^

various test facilities. Tassicker /25/ has recently reviewed the test

programme under way in the U.S.A.

222 -

Electrostatic precipitation (ESP) involves three processes:

a) charging of suspended particles by corona discharge;

b) migration and deposition of the charged particles under the in­

fluence of an applied electric field;

c) detachment and recovery of deposited material from the collecting

electrode.

Stages a) and b) are usually combined, because the electrostatic field

also produces the corona discharge. For ESP to be viable, the field

strength necessary to establish a corona must be smaller than that at

which "sparkover" occurs between the electrodes. Therefore ESP cannot

be used at temperatures above 800 °C and ambient pressure. Increasing

pressure increases the sparkover voltage sharply while the corona

voltage increases more gradually /26, 27/. The pressure at which the

two become equal is called the "critical pressure". Negative corona

gives more stable operation. The critical pressure depends on electrode

geometry, but is typically a few bar for PFBC duty. Because of the

variation of corona and sparkover voltages with pressure and tempera­

ture, it is possible in principel to use higher voltages for high-pres­

sure high-temperature duty than at less extreme conditions /27/. Thus,

in principle, higher electrophoretic forces can be imposed, to offset

the higher drag resulting from the higher gas viscosity at elevated

temperatures.

The most extensive test programme, using the Curtiss-Wright PFBC faci­

lity in the USA, was beset by so many problems of maloperation that it

is far from clear what real conclusions can be drawn /28/. Smaller

scale tests at Denver Research Institue /29/ and the University of

Essen /30/ suggest that good collection efficiencies can be achieved.

However, the current densities needed appear to be high, even with the

best electrode geometries yet conceived. Furthermore, construction

materials to maintain the necessary electrical insulation between the

electrodes are needed; the Research-Cottrell preciptator used at

Curtiss-Wright, for example, used high-purity alumina which still had

to be cooled to about 800 K.

Thus it remains uncertain whether electrostatic precipitation can be

used for PFBC plant without unacceptable power consumption. Major

I

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- 223 -

difficulties can also be foreseen in applying ESP to gasification. It

is very uncertain whether it will be possible to maintain stable corona

discharge in the dirty reducing gases obtained from coal gasification.

Moreover, because the particles are so uncohesive, they are unlikely to

form a stable cake on the collecting electrode and very likely to re­

disperse. A build-up of conducting particles in the precipitator would

in turn destabilise the corona. Overall, electrostatic precipitation

remains potentially usable but unproven for PfBC plant, and very doubt­

ful for use in gasification cycles.

3.2 Barrier filtration

The term "barrier filter" is used here to mean a device using a rigid

or flexible filter medium, fixed in place within the filter housing and

functioning largely by interception of the dust particles. Such filters

must be cleaned in situ by some mechanical process such as vibration or

reverse pressure pulse or reverse gas flow. Particularly for high-tem-

perature high-pressure duty, where the cleaning action is difficult to

apply, cleanability is of as much concern as filter efficiency. The

durability of the filter medium is also of prime concern, and rules out

the use of metal filters.

Two general types of barrier filter have been investigated, both made

from suitable ceramic materials: they are woven or felted fabrics, and

rigid sintered or bonded elements. Fabrics are formed into cylindrical

bags, as used for conventional fabric filters. Table 1, after Chang et

al. /31/ summarises the differences between woven and felted media.

Ceramic fabrics have been widely tested, notably at Acurex Corp. /32/,

Westinghouse /33/ and the University of Essen /34/. Substantial diffi­

culties have been encountered, notably associated with embrittlemenet

of the ceramic fibres. Because these media are required to flex for

cleaning, embrittlement can lead to loss of mechanical integrity. This

is particularly serious with woven media, for which a relatively small

crack or tear leads to catastrophic loss of filter efficiency. In the

author's view, rigid ceramic filter media hold more promise.

224 -

Table 1. Comparison of woven and felted fabrics.

Felted Woven

Very high collection efficiencies High collection efficiencies(> 99.9 %) (> 99 %)

High face veolocity Low face velocity(up to 0.1 m/s) (0.005 to 0.03 m/s)

Requires high cleaning energy. Low energy cleaning, by pulsing.usually by pulse shaking, or reverse flow

Rigid ceramic media have been tested fairly extensively in two forms;

"candles", which are hollow cylinders with closed ends, much like con­

ventional filter bags; and "crossflow" filters, which take the form of

blocks containing many small gas channels (see below). Candles have

been tested at Westinghouse /33/, University of Essen /34/, and by

British Coal and the University of Surrey /35/. Types of candle include

sintered granules, granule and fibre mixtures, bonded fibres, and gra­

nular substrates with a fine surface layer on the "dirty" face of the

filter. It appears possible to find formulations, for both PFBC and

gasification duties, which will maintain their strength and integrity

over extended periods. However, a more serious operating problem is

long-term rise in the resistance of the medium, due to "blinding" with

particles not removed by cleaning or to fouling with some other type of

deposit. Current developments therefore concentrate on finding media

and appropriate cleaning cycles to reduce "blinding". Because the be­

haviour of a dust cake depends on interparticle forces, which are in

turn very dependent on the chemical environment, it is likely that

media and cleaning strategies will differ between different applica­

tions. It appears that, for PFBC fly ash, pulse cleaning causes less

redispersion than does reverse flow cleaning /35/. However, redisper­

sion is likely to be more of a problem in gasification, because of the

low cohesiveness of gasifier fines.

The other geometry, the "crossflow" filter, is shown in Figure 5. This

type of filter has been tested most extensively by Westinghouse /36/,

i r

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Figure 5. Ceramic crossflow filter (schematic).

and was identified by Rubow et al. /7/ as promising. The gas enters on

one face of the filter and exits from an adjacent face, having passed

from one row of channels through the ceramic to the next orthogonal row

of channels. The advantage of this device is that it accommodates a

very high surface area within a small volume, and this is clearly par­

ticularly attractive for high pressure duty. As for ceramic candles,

filtration is essentially "absolute", and the operating pressure drop

appears to be comparable with other media /36/. Cleaning is achieved by

subjecting the exit face of the filter to a high pressure pulse, which

must be of sufficient duration to lift the cake from the ceramic mem­

brane and to transport it back out of the filter channels. This probab­

ly limits the length of channel which can be used in any single filter

I E

F

- 226 -

block. It is also possible that re-entrainment could occur for uncohe-

sive gasifier fines, although this remains to be investigated.

A rather different concept has been investigated at the Institute of

Gas Technology /37/, confusingly also called "crossflow filtration",

although the term "throughflow filtration" is a better description and

will therefore be used here. The concept is shown schematically in

Figure 6. The greater part of the gas flow, typically 5 * in the IGT

trials, passes through a ceramic medium and is filtered. The balance is

used to transport all the particulate matter. Thus the "throughflow

filter" serves to concentrate particles into a small proportion of the

gas flow, which can then be cooled and cleaned by more conventional

means. Clearly this concept is only applicable when the particles are

uncohesive, so that they do not form a cake and can therefore be trans­

ported through the filter. The IGT trials used ceramic cylinders, open

at both ends, with the dirty gas on the inside, but other geometries

are clearly possible.

In general, rigid ceramic barrier filters appear to represent the best

developed option for tertiary gas cleaning. They are being evaluated,

for example, at the Grimethorpe demonstration PFBC facility.

r' r-

I I

r

Dirty

Gas t t

1 fn

*/sss

Filtered Gas

Particle­

laden Gas

Figure 6. Concept of throughflow filter.

3.3 Granular bed filters

By contrast with barrier filters, granular beds use unbonded filter

elements. The filter may be cleaned in situ as for a barrier filter;

examples of this practice are the Lurgi "twist layer" filter /38/ and

the Ducon filter /39/. However, this type of design retains the diffi­

culties of intermittent cleaning which arise with barrier filters. The

potential advantages of granular bed filters include continuous opera­

tion, whith the medium passing continuously through the filter zone,

and cleaned elsewhere to be recycled or passed to disposal or pro­

cessing. This general type of device may have additional advantages for

simultaneous removal of particulates and vapours, such as sulphur

oxides or alkali salts. We therefore concentrate here on two types of

granular bed filter which embody this advantage: the Moving Granular

Bed Filter developed by Combustion Power Company (CPC) and given a

positive assessment by Rubow et al. /?/, and the Fluidised Bed Filter,

developed by the University of Surrey with British Coal /40/ and not

evaluated by Rubow.

Granular bed filters differ from membrane filters in a further im­

portant respect. The passages carrying gas through the filter are much

larger than the dust particles, so that "absolute" filtration is not

expected (unless the dust forms a "cake" on the upstream face of the

filter, and this cannot occur with moving or fluidised beds). In the

filtration process, dust must be brought into contact with the filter

granules. For the range of particle sizes of interest for turbine ex­

pansion. gases, the most important mechanism is inertial impaction /41/.

The dust must then adhere to the collector, without rebounding or being

re-entrained. This depends on a balance between the kinetic energy of

the impacting particle and the adhesion between the particle and the

collector /41/. For cohesive fly ash particles, retention can be good

provided that the filtration velocity is not too high /5, 41/. For

cohesive gasifier fines, there may be very little retention of the dust

in the filter, and efficiency is typically low.

The basic concept of the CPC Moving Granular Bed Filter is shown in

Figure 7. The dirty gas passes upward through a bed of filter granules,

I l

wi

- 228 -

MEOW RETURN LEG x. GAS INLET

'GAS OUTLET

INTERNAL MEDIA RESERVOIR

8 EQUALLY SPACED MEDIA FEED LEGS

INNER MEDIA FLOW ANNULUS

I DIRTY i MEDIA

Figure 7. CPC moving granular bed filter (schematic).

which are particles of gravel or alumina a few millimetres In diameter

/42/. The filter granules are drawn off at the base of the filter, and

raised In a high-velocity pneumatic conveying line to a fluidised bed,

from which they are returned to the filter. It appears that the depo­

sited particles are dislodged In the conveying line, and entrained off

from the fluidised bed. Thus the filter operates by transferring

particles from the process stream to the cooler air flow passing off

from the fluidised bed holding vessel, and probably by agglomerating

them in the process so that they are easier to collect in a cyclone.

Overall particle collection efficiencies above 90 % have been reported

for cohesive fly ash, with less than 1 kg of dust collected per 100 kg

of collector circulated /7, 43/.

i r

r ' t

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W

i_ i_

For higher collection efficiency still, a slightly different design has

been recommended /7/ in which the granular filter medium passes down

between two concentric cylindrical louvered panels as shown in Figure

8. The gas flows outwards through the annular filter. Corona electrodes

are mounted on the inlet side, to precharge the particles and thus

enhance collection by electrostatic attraction. Rubow et al /7/ recom­

mended testing this device on PfBC duty. In the meantime, it remains to

be seen whether the electrostatic enhancement is sufficiently strong to

justify the complexity and expense; given the difficulties already

experienced with conventional electrostatic precipitation, this appears

doubtful. Furthermore, for the reasons noted above, it is unlikely that

this type of filter will be of any use in gasification.

The fluidised bed filter is another device in which the filter medium

can be renewed continuously. The concept is shown schematically in Fig­

ure 9. The dusty gas is used to fluidise a bed of collector granules.IONIZWG

DIRTY GAS INLET —

CLEAN GAS OUTLET

GRAVEL BED FA.TER

Mgure 8. EFB electrostatically augmented granular bed filter.

- 230

FRESH FILTER MEDIUM

GAS ENTRY

SPENT FILTER MEDIUM

i r

Figure 9. Fluidised bed filter (schematic;.

Filtration occurs predominantly by inertial impaction in the jets which

form above the ofifices in the distributor plate through which the gas

enters the bed. The function of the bed itself is then to immerse these

jets so that filter granules are continuously entrained into them; thus

a bed only a few centimeters deep is needed. In some respects, the

fluidised bed filter is analogous to the "dry plate scrubber" developed

by Air Pollution Technology Inc. /46/. However, by using a fully

fluidised bed, it becomes possible to introduce means to enhance the

retention of dust by the filter /40/. For fly ash, for example, a che­

mical additive may be injected to the bed to reduce the sintering point

of the ash and so ensure retention. For gasifier fines, it is possible

to spray, a heavy fuel oil into the bed. The oil "cracks" to leave a

tarry residue which binds the particles to the bed granules. Thus the

device can be operated as an agglomerating filter, fed with seed

particles and producing as product agglomerates of dust onto "seed"

cores. Alternatively, such a device can be used as pre-agglomerator

before a cyclone separator /40, 47/.

The fluidised bed filter has yet to be tested on either PfBC or gasi­

fier duty. However, it has clearly reached a stage of development where

this is poortune, particularly in view of its advantages of compact

design and continuous operation.

7 7

l l

r-'i

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4 COLLECTION OF ALKALI SALTS

The extent to which it is necessary to remove alkali salts from the

gases derived from coal remains to be determined. It is clear that it

will be very difficult indeed to satisfy the tolerances on alkali level

set by most turbine manufacturers, but it is far from clear whether

these tolerances are actually required by the turbomachinery. Sodium

and potassium compounds are the main concern. In gross terms, sodium

and potassium chlorides are volatile at the temperatures of concern

here, whereas the sulphates are very much less volatile /3/. Thus if

the alkalis form as sulphates, they *i11 be removed to some extent with

the particulate matter. In PFBC gases, sulphates form by reactions like

2NaCl + H20 + S02 + 1/2 02 Na2S04 + 2HCL (2)

Thus sulphate condensation may be suppressed if the gases contain

significant levels of HC1.

This demonstrates a particular difficulty in assessing the release and

effects of alkali compounds: they are strongly dependent on the chlo­

rine level in the coal. It appears /4/ that high HC1 levels encourage

release of alkalis from minerals in the coal by reactions like that

with feldspar:

Na20*Al 203*2Si02 + 2HC1 2NaCl + Al 203*2Si02 + H20 (3)

and then inhibit sulphate condensation by reactions like (2). Most work

on alkali collection /e.g. 48/ has concentrated on absorbing or re­

acting the chlorides with mineral "getters". The getters could form a

bed through which the gases are passed, or might conceivably be added

directly to the combustor or gasifier along with the limestone used for

sulphur collection. An example of a "getter" is kaolin, in which case

the "gettering" reaction for sodium chloride is the reverse of reaction

B. Thus this reaction also serves to demonstrate why HC1 also reduces

the possible effectiveness of getters, by displacing thermodynamic

equilibrium to leave more alkali as chloride vapour.

i r !

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232

North American coals are typically low in chlorine. British, and some

other European, coals are high in chlorine. This difference explains

why the chlorine problem has been regarded rather differently on the

two sides of the Atlantic. Low-chlorine American coals are subject to

less alkali release, and "gettering" appears to be a feasible way to

reduce alkalis to acceptable levels /48/. The indications, at present

based largely on thermodynamic estimates /4/, are that known getters

will not achieve acceptable alkali levels for PfBC plant using high

chlorine coals unless the gases are cooled slightly.

For gasification the position is potentially more serious, because

sulphate condensation by reaction A cannot occur in a reducing environ­

ment. It appears that "gettering" can effect some reduction in alkali

levels, but to achieve acceptable levels it may be necessary either to

cool the gases or to remove HC1 /4/.

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1. Robson, B. Energy Research 1, 157 (1977).

2. Reed, G. P. "FilTech" Confce. London 1983, Filt. Soc. P. 384.

3. Scandrett, L. A. & Clift, R. J. Inst. Energy 57, 391 (1984).

4. Singh, A. et al. IChemE Symp. Ser. 99, 167 (1986).

5. Coury, J. R., Reed, G. P. 4 Clift. R. To be published.

6. Clift, R., Grace, J. R. 4 Weber, M. E. Pubbles, drops and

particles. New York 1978, Academic Press.

7. Rubow, L. N. et al. Technical and economic evaluation of ten high

temperature, high pressure particulate cleanup systems for pressu­

rized fluidized bed combustion. Rept. D0E/MC/19196-1654 (1984).

I L

rJL

8. Stairmand, C. J. Trans IChemE 29, 356 (1951).

9. Henry, R. F. & Podolski, W. F. Cyclone performance estimates for

pressurized fluidized bed combustion. Rept. ANL/CEN/FE-81-4

(1981).

10. Clift, R. et al. Textbook of particle technology. London,

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11. Muschelknautz, E. VDI Berichte 363, 49 (1980).

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(1986).

14. Wheeldon, J. M. et al. IChemE Symp. Ser. 99, 45 (1986).

15. Syred, N. et al. IChemE Symp. Ser. n:o 99, 17 (1986).

16. General Electric, US DOE Rept. FE-2357-70 (1980.

17. Giles, W. B. High performance cyclone development. 3rd Symp. on

Transfer and Utilisation of Particulate Control Technology,

Orlanda, FA, C0NF-810319-1 (1981).

18. McCabe, J. T. EPA/DOE Symp. on High-Temperature High-Pressure

Particulate Control. EPA-600/9-78-004, C0NF-770970, p. 55 (1977).

19. Paul, D. D., Razgaitis, R. & Jordan, H. IChemE Symp. Ser. 99, 1

(1986).

20. Ciliberti, D. F. & Lancaster, B. W. Chem. Eng. Sci 31, 499

(1976).

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w

21. General Electric, US DOE Rept. FE-2357-39A (1980).

234 -

22. Tate, A. H. J. et al. IChemE Symp. Ser. 99, 89 (1986).

23. McCreight, L. R. Proc. High-Temperature High-Pressure Particulate

and Alkali Control in Coal Combustion Process Streams. US DOE

Contractors' Meeting, C0NF-810249, p. 583 (1981).

24. Weber, E. et al. Environment International 6, 349 (1981).

25. Tassicker, 0. J. IChemE Symp. Ser. 99, 331 (1986).

26. Robinson, M. Electrostatic precipitation. In: Strauss, W. (ed.)

Air pollution control. Part 1. Wiley-Interscience 1971.

27. Bush, J. R., Feldman, P. L. & Robinson, M. JAPCA 29, 365 (1979).

28. Kumar, K. S. DOE Contractors' Meeting (1983).

29. Rinard, G. IChemE Symp. Ser. 99, 315 (1986).

30. Riepe, T. & Wiggers, H. IChemE Symp. Ser. 99, 297 (1986).

31. Chang, R. et al. 3rd Annual Contractors' Meeting on Contaminant

Control in Hot Coal-Derived Gas Streams. US/DOE/METC/84-6 (1983).

32. Chang, R. et al. IChenC Symp. Ser. 99, 177 (1986).

33. Lippert, T. E. et al. IChemE Symp. Ser. 99, 215 (1986).

34. Schmidt, D., Schulz, R. A Bender, J. IChemE Symp. Ser. 99, 245

(1986).

35. Cheung, W. et al. To be published.

36. Ciliberti, D. F. & Lippert, T. E. IChemE Symp. Ser. 99, 193

(1986).

IT

i r

f T•J1;. -ii.

235 -

37. Rush, W. F. et al. An exploratory study of a combined sonic

agglomeration and cross-flow filtration system for hot gas

clean-up. AIChE Conf., Chicago, Nov. 1980.

38. Use of gravel filter beds for dust removal from the exhaust air of

clinker coolers. Lurgi Apparate-Technik GmbH, 1977.

39. Kalen, B. & Zenz, F. A. AIChE Symp. Ser. 137, 397 (1974).

40. Ghadiri, M., Seville, J. P. K. A Clift, R. IChemE Symp. Ser. 99,

351 (1986).

41. Coury, J. R., Thambimuthu, K. V. & Clift, R. Powder Tech. 50, 253

(1987).

42. Guillory, J. L., Placer, F. M. A Grace, D. S. Environment

International 6, 387 (1981).

43. Moresco, L. L. A Cooper, J. L. US DOE Contractors' Meeting.

C0NF-810249, p. 507 (1981).

44. Doganogly, Y. et al. Trans. IChemE 56, 239 (1978).

45. Clift, R., Ghadiri, M. A Thambimuthu, K. V. in: Wakeman, R. J.

(ed.) Progress in filtration and separation, Vol. 2. London 1981,

Elsevier.

46. Calvert, S. et al. EPA Symp. on Transfer and Utilisation of

Particulate Control Technology, Vol. 3. EPA-600/7-79-044C, p. 405

(1979).

47. Barone, L. et al. Powder and Bulk Solids Tech. 1, 15 (1983).

48. Lee, S. H. D. A Myles, K. M. IChemE Symp. Ser. 99, 149 (1986).

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237 -

DIDER HOT GAS FILTER SYSTEMS

Weidner UlrichDidier filtertechnikD 6719 Eisenberg / Pfalz, FRG

1 INTRODUCTION

Didier, founded in 1834, is one of the biggest and most renown companies in refactory field, solving high-temperature problems up to 2 000 °C

Didier filtertechnik (DF) founded by Wilhelm Schuler in 1887 and acquired by Didier in 1957, is a leading manufacturer of filter material (porous ceramics), gas and liquid filters as well as of fine-bubble diffusers for sewage aeration.

For 20 years DF has been planning and constructing high-efficient hot gas filter system fot the direct, dry cleaning of hot gases at temperature ranges up to 1 000 °C. According to this process, in which the filter is located immediately after the firing systems, that is heat exchangers are installed behind the separators, corrosion problems are thus prevented - as no cases arise in which the temperature falls below the dewpoint Optimum heat recovery is also provided.

In the flue gas dust precipitation and product recovery the limiting temperature of application fin* ceramic candle filters - depending on the material of filter media - was for a long time approximately 400 °Gto 600 °G

In 1970 / 71 DF developed for the first time refactory-lined filters on a commersial scale for continuous temperatures 800 °C to 900 °C.

In 1981 / 82 a complete pilot plant was constructed by DF for Japanese licence, for

working temperatures of 1000 °C.

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- 238 -

The filters mentioned so far operate at atmospheric conditions. On the basis of developments in the Didier Group with regard to special high-tech tube-sheets, refactory- lined and high-temperature insulated pressure vessels, complete high-temperature high- pressure (HTHP) filters are designed since 1980. They are achieving excellent clean gas values (e.g. residual dust content in clean gas <1 mg/nm3) and thus satisfy the conditions required in pressurised fluidised bed combustion and gasification - based (PFBC/G) combined cycle applications before the gas turbines.

2 HOT GAS FILTER ELEMENTS

DF manufactures high efficient filter elements of alumina silicate and silicon carbide mateial, fired at temperatures around 1 400 °C. They are made in different shapes (1):

2J Cylindric hollow-todies;

Filter cylinders, open at both ends; filter candles, closed at one end with flange at the open end.

In hot gas filtration mainly flanged candles are used; the new hemispherical flange design(2) was developed in the 1970's, improving the backblowing behaviour of candles.

22"Cross-flow" (31 elements

DF makes "Kanalsteine" (4) among others. When using such type of filter elements the filtering surface can be increased in a given space.

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3 HOT GAS FILTERS FOR ATMOSPHERIC CONDITIONS

Highlights:From 1966 to 1969 DF developed hot gas pilot filters for temperatures up to 600 °C to 700 °C (5). In 1971 DF constructed for the first time scaled-up hot gas filters for a continuous operating temperature of 800 °C to 900 °C with refactory-lining and 91 - later on: 85 - pieces of flanged candles per unit (6).

In 1974 / 75 a complete DF pilot filter unit, consist of hot gas filter, blower, backblowing gas receiver, piping and measuring device, was tested with best results (7).

From 1980 larger hot gas filter units with elements of 1 500 mm length were commisioned

DF designed and constructed for the first time a complete pilot filter plant, operating at temperatures of 1 000 °C, with integral heat recovery system (8).

4 HTHP FILTERS

The use of combined cycle generation systems incorporating a gas turbine ahead of the syeam cycle is seen as a promising technology for raising the efficiency of electric power generation. For the protection of a gas turbine - operating with high entry temperatures - a paticle loading of 4 mg/nm3 will be tolerated. Conventionally the particle cleaning system would be wet scrubber. Hot gas filtration would raise the efficiency of PFBC as well as of gasification-based process (9).

DF development on HTHP filter was started in 1980. Already in 1981, some filters could be commissioned for designed pressure of 120 bars and a design temperature of 350 °C

(10).

In 1985/86 scaled-up filter units were designed, meeting the tubesheet and sealing problems (11).

Characteristics:Large filter units with a number of filter elements will be constructed in order to

minimize costs of filter vessel, valves and pipes.Special tubesheet of high-tech material are arranged in a big support plate in order

to meet tubesheet problems, especially the manufacturing costs and the dimension

(diameter) limitations.An optimal sealing of filter elements, gaskets, tubesheet and support plate will be

achiebed - and raw gas leakage will be prevented - by using ceramic materials only with

same thermal expansion.

- 240 -

References:

(1) Didier Filtertechnik, "Porous Ceramic Filter Media", Technical documentation no V 2.20 E, November 1983.

(2) Didier Filtertechnik, Sketch No. G/Wr, March 1985.

(3) Ciliberti D.F., and Lippert T.E.: "Ceramic Cross Flow Filters for Hot Gas Cleaning", September 1986.

(4) Hofbeck R., Kafla H., and Weidner U.: "Filtereinrichtung zur Russpartikel- filtration", October 1985.

(5) Hempelmann W., and Krause H.: " behandlung der festen radioaktiven AbfKlle im Kemforschungszentrum Karlsruhe", September 1970.

(6) Dirks F., and Hempelmann W.: "The Incineration Plant of the Karlsruhe Nuclear Research Center - a Regional Solution", 1984.

(7) David A., and Wittenmeier F.: "Filtration heissen Rauchgase iiber 600 °C", May 1976.

(8) Didier Filtertechnik, Sketch No. 4S -1101, May 1982

(9) Reed G.P.: "The Economics of Hot Gas Cleaning for Coal Gasification Combined Cycle Power Generation", September 1983

(10) Didier Filtertechnik, Sketch No. 4S - 1058, December 1980

(11) Weidner U.: "High-Temperature High-Pressure Ceramic Candle Filters for Pressurised Fluidised Bed Combustion and Gasification - Based on Combined Cycle applications", September 1986

Enclosed:Drawing and sketches of DF hot gas filter material, non-pressure filters and HTHP filters

(figures 1 to 18).

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Figure 1. Filter cylinders are open at both ends (outside diameter x inside diameter x length).

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Figure 2. Filter candles are closed at one end (outside diameter x inside diameter x length overall).

Figure 3. Flanged filter candles have a flange at the open end(outside diameter x inside diameter x length overall, Wwith outside diameter x height of the flange). ]_

242 -

Figure 4. Flanged candle with hemispherical shape. DF hot gas filter material: cylindric hollow bodies.

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- 243 -

E22X3 Reingaskonole Raw aas channels

I I Rohgaskanolp Clean gas channels

Figure 5. DIDIER cross-flow filter.

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Figure 6. DIDIER "Kanalstein" ("channel tile"). DF hot gas filter mate­rial: cross-flow filter (patented).

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Figure 12. Hot gas filter unit, fully-automatically controlled. Design temperature 600 °C. Scaled-up DF hot gas filters for atmo­spheric conditions: single-chamber filters.

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247 -

Back-washing gas Raw gas

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248 -

Filter elcinem

Figures 15 and 16. Hot gas filter, design temperature 600 C. Scaled-up DF hot gas filters for atmospheric conditions: multiple chamber filters.

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pressure 18 bar.

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- 251 -

PARTICLE MEASUREMENTS AT THE OTANIEMI PFBC/G TEST RIG

Seppo Hulkkonen, M.Sc.(Eng)Helsinki University of Technology ]

Laboratory of Energy Economics and Power Plant Engineering 02150 Espoo, Finland

Abstract

One of the most difficult problems in the commercialization of PFBC technology is hot gas cleanup. Hot and compressed gases have to be cleaned effectively enough before the gas turbine to ensure a reasonable gas turbine life time. At the PFBC/G test rig combustion gases are cleaned in a two-stage cyclone collector. The efficiency of the Erst cyclone is quite low and it is considered as a pre-collector. The secondary cyclone is similar but not identical to the standard high efficiency Stairmand design. The diameter of the secondary cyclone is 103 mm.

Particle measurements are based on isokinetic sampling. Isokinetic sample flow is extracted before and after the two cyclones. The particle concentration is measured using a total filter which is electrically heated to maintain the temperature at 200 *C. The system operates at plant pressure. The filter papers used are of 90 mm in diameter and made of glass or quartz fibres. The size distribution is determined from total filter samples using Coulter Counter analysis and also using a virtual impactor. The impactor is a four-stage cascade centripeter made by Bird&Tole (UK). The centripeter is used at atmospheric pressure and 140 *C temperature. The isokinetic sample flow is expanded through a critical nozzle and diluted using heated clean air before impactor measurement. The performance of the cyclones is determined from the exit dust sample and the samples taken from the captured dusts.

The first measurements were made at pressures in the 5 to 10 bar range at different excess air levels using crushed peat pellets as fuel. The elutriation from the reactor was 2000- 4000 mg/m3n and the cyclone system reduced the concentration of particles to 70-150 mg/m3n at the secondary cyclone outlet. The overall collection efficiency of the secondary cyclone was about 90 %. The mass mean diameter of the elutriated particles was about 30 |xm and at the outlet the mass mean diameter was 2-4 |im. Size distribution analysis by Coulter Counter agreed quite well with impactor measurements.

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1 INTRODUCTION

The purpose of this paper is to present particle measuring systems and the first results obtained at the pressurized fluidized bed combustion/gasification test facility in Otaniemi. The tests in autumn 1987 were carried out using crushed peat pellets as fuel at pressures in the S to 10 bar range. Particle concentration was measured using a total filter system and particle size distribution was determined using a virtual impactor and also by Coulter Counter analyses. The performance of the two cyclones was calculated based on the mass balance techniques.

2 EXPERIMENTAL EQUIPMENT AND PROCEDURES

The PFBC/G test rig used in this study consists of a combustor vessel lined to an inside bed diameter of 15 cm. The expanded bed height is 1.2 m. The diameter of the freeboard is 0.25 m and the height of the freeboard is 2.3 m. Vertically oriented watercooling coils are immersed in the bed. Crushed peat pellets as fuel are screw-fed into the bed through a lock hopper system. The feeding port is about 12 cm above the nozzles of the air distributor. The maximum solid feed rate is about 30 kg/h. The combustor is capable of operating at pressures up to 10 bar (abs), at temperatures up to 950 *C, and at superficial gas velocities up to 1.5 m/s with expanded beds up to 1.2 m. The bed height can be controlled by adding material from a pressurized hopper or by discharging material through the air distributor. Flue gas exiting from the combustor passes through two cyclone separators. Solids collected in cyclones are removed through lock hoppers. A schematic layout of the combustor and off-gas path is shown in Figure 1.

The cyclones are made of heat-resistant steel and they are located in a large pressure vessel. The first cyclone is a low efficiency pre-collector and the secondary cyclone is similar but not identical to the standard high efficiency Stairmand design. The barrel diameter of the secondary cyclone is 103 mm. The details of the secondary cyclone are

given in Table 1.

- 253 -

CONCENTRATION AND SEE DISTRIBUTION MEASUREMENT

CONCENTRATIONMEASUREMENT

ASH HOPPERS

TKK VTT IVQ

PFBC/G TEST RIG

Figure 1. PFBC/G test rig.

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Table 1. The dimensions of the secondary cyclone.

Dimensions relative to barrel diameter DPFBC/G Secondary Cyclone

Standard High EfficiencyStairmand

Height of inlet A/D 0.51 0.5Width of inlet B 1/D 0.205 0.2Gas outlet diam. Dl/D 0.54 0.5Outlet pipe insertion length S/D 0.57 0.5Height of cylindrical section Hl/D 1.52 1.5Total cyclone height H/D 4.04 4.0Dust outlet diam. B/D 0.40 0.375

The measurements were made during six tests of 5 to 15 hours duration giving the total time of about 60 h. The operating conditions of the combustor are given in Table 2.

- 254 -

Table 2. Combustor operating conditions.

TestNo.

Thermal input [kW]

Pressure[bar]

Bed temperaturera

Fluid, velocity [m/s]

Excess air [02-%]

6A 131 7.0 860 1.16 1.86 72 7.0 850 0.64 1.71 110 10.0 830 0.65 1.63 100 10.0 840 0.84 5.857A 67 5.0 870 0.83 1.59A 60 5.0 840 0.92 5.2

Natural silica sand is used as bed material. The mass mean diameter of the bed materialincluding also some fuel ash is 520 |im. The fuel properties are given in Table 3. The Y ITmass mean diameter of the crushed peat pellets is 4.85 mm.

Table 3. Fuel properties.

Effective heating value 15.6 MJ/kgMoisture content 20 J %Ash content (on dry base) 4.2%Particle size 4.85 mm (35 % >6 mm. 10 % < 1 mm)

2.1 Measuring procedures

Particle measurements are based on isokinetic sampling. Isokinetic sample flow is extracted before and after the two cyclones. The probes in use are null-type probes of 6 mm in diameter. The isokinetic conditions are maintained by keeping the static pressures inside and outside of the probe head equal. Additionally, the isokinetic flow is calculated from measured process values by means of a computer. The probes are air-cooled to reduce the temperature to 250 - 300 *C.

Particle concentration is measured using a total filter system. The filter papers are of 90 mm in diameter and the filter housing is electrically heated to keep the temperature above 200 *C. The filter materials in use are quarz and glass fibres (Munktell MK360 and Munktell MG160). Following the filter, the gas pressure is reduced across a flow control valve, the flow rate being controlled such that the sample is extracted isokinetically. Depending upon the dust loading in the duct, the sampling period can be of up to 90 minutes duration. The collected amount of dust is in the range from 100 mg to 2000 mg. After cooling and drying in an icebath condenser the gases pass through a dry gas meter. A schematic diagram of the sampling system is shown in Figure 2.

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MEASUREMENT OF PARTICLE CONCENTRATION

... ..Nitrogen

AIRCOOLINGAIRCOOLING

Cooling, drying

TOTAL FILTER

DRY GAS METER

Rotameter

Figure 2. Concentration measurement.

Particle size distribution at the secondary cyclone outlet is determined by a virtual

impactor and by making Coulter Counter analysis of the total filter samples.

The virtual impactor is a four stage cascade centripeter made by Bird&Tole (UK). The

cascade centripeter separates the particles to size ranges according to their aerodynamic

sizes. A sectional view of the cascade centripeter is shown in Figure 3.

The particles are collected on a filter paper (Whatman GF/A) backing each nozzle. About

3 % of the total sample flow is taken through the filter thus helping the separation of small

particles.

The impactor is used at atmospheric pressure and at 140 *C temperature but it is calibrated

at ambient temperature. The corrected cut sizes for the stages at higher temperatures can

be calculated from equation

d(T) = d(293) (1)

where d is particle size, T) is viscosity of the flue gas, and T is temperature (K). The cut

sizes at 140 *C are 14.5,4.4, and 1.6 |im from first to third stage. The last stage collects

all particles below 1.6 |im (aerodynamic diameter).

- 256 -

STAGE I

STAGE 2

STAGE 3

STAGE 4

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Figure 3. Cascade centripeter (Bird&Tole, UK) 131.

The isokinetic sample flow is expanded to atmospheric pressure through a critical nozzle

and diluted using clean air before impactor sampling. The flow can be varied by changing

the nozzle diameter. The nozzles in use are of 0.5 - 0.9 mm in diameter. Following the

impactor the sampled gases flow through a condenser unit to a leakless vacuum pump and

on through a dry gas meter. The sampling time is up to 2 h. The measuring facility is

presented in Figure 4.

The Coulter Counter analyses are performed of the dust samples captured on the total

filter. The dust cake on the filter is carefully scraped off and subjected to ultrasonic

dispersion for 5 minutes before Coulter Counting. This is done to brake the agglomerates

as the dust collected from pressurized fluidized bed combustor is known to agglomerate

once sampled. The electrolyte used is waterbased with 6 % NaCl. Particle size

- 257 -

distributions on a weight basis are obtained assuming that all observed particles were

spheres of equal density.

MEASUREMENT OF SIZE DISTRIBUTION

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Rotameter

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Figure 4. Impactor measurement.

Additionally the size distribution of the dust captured by the cyclones is determined using

Coulter Counter. The material has a wide size range from 0 up to 200 pm and so the

analysis is completed in two stages; Erst the entire sample is air-sieved (Alpine) to

determine the distribution of the particles over 11 pm and then the size distribution of the

fine particles is determined using Coulter Counter. The grade efficiency of the cyclones is

then calculated from the mass balances.

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3 RESULTS AND DISCUSSION

3.1 Elutriation and emissions f 7

The elutriation from the reactor varies in the range 2000 - 3000 mg/m3n (1700 - 2600

ppmw, by weight) and the cyclones reduce the concentration of particles down to 70 -150 mg/m3n (60 -130 ppmw) at the secondary cyclone outlet. Typical particle loading curves

are shown in Figure 5. The curves are based on Coulter Counter analyses.

258 -

CUMULATIVEPARTICLELOADING[mg/m3n]

10000D-10 bar

1000

1000DIAMETER [nm]

i rAT COMBUSTOR •O- at PRIMARY •X* AT SECONDARYEXIT CYCLONE EXIT CYCLONE EXIT

Figure 5. Particle loadings.

The elutriation rate depends on fluidizing velocity but also on excess air level. The

excess air has an effect on the fuel feed rate and thus also on the amount of ash fed in to

the reactor. The elutriated material is mostly fuel ash and the amount of unbumt fuel in

the dust is very small, typically 0.1 %.

The mass mean diameter of the elutriated particles is about 30 pm and the biggest partic­

les observed are 123 - ISO pm. The calculated critical diameter (below which all particles

would be elutriated) based on terminal velocity in the freeboard is 30 - 90 pm depending

on the operating conditions. In the bed, die critical diameter is 130 - 240 pm. As can be

expected the big diameter of the freeboard (compared with that of the bed) and the heigth

of the freeboard have quite a strong effects on elutriation rates. The elutriation is smaller

than reported by many coal burning test facilities. Also the ash content of the fuel is low

compared with that of coal. The density of the ash captured by the secondary cyclone was measured to be 2600 kg/m3.

The mass mean diameter at the secondary cyclone outlet is 2 - 4 pm. The size

distributions were obtained by Coulter Counter analyses and the virtual impactor. The

results of the two methods agreed quite well although! in some tests some variation was

experienced. The aerodynamic diameters of the impactor measurements are compared

1 l

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- 259 -

with the mass (volume) distribution determined by Coulter Counter. Data are reported in

terms of physical diameter. Following definition is used:

dp = dpa(C’pp)-M (2)

where dp is the physical diameter, dpa is the aerodynamic diameter, C is the slip correction factor, and pp is particle density.

Typical curves of size distribution at the secondary cyclone outlet determined by both techniques are shown in Figure 6.

SIZE DISTRIBUTION AT THE SECONDARY CYCLONE OUTLET

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Figure 6. Size distribution at the secondary cyclone outlet.

Figure 6 shows two independent Coulter and impactor measurements made during one

test of IS hours duration. The impactor measurements tend to show slightly higher

concentration of very fine particles. FI

3.2 Cyclone performance

About 95 % of the particles in flue gases leaving the reactor are removed in passage

through two cyclones. The material captured by the primary cyclone has a mass mean

- 260 -

diameter of 26 - 37 pm. The overall efficiency of this cyclone is only about 66 % and it

is considered as a pre-collector. The material captured by the secondary cyclone has a

mass mean diameter of 20 - 26 pm and an overall efficiency of 90 %. Particleconcentration in the flue gases leaving the secondary cyclone is generally 70 - 150 ]prmg/m3n. The operating parameters of the cyclones are presented in Table 4. The inlet

velocity of the cyclones is quite low, being normally 11-12 m/s.

Table 4. Operating conditions of the cyclones.

Inlet velocity [m/s] 10-20Pressure [bar] 5-10Temperature [*C] 600-700Pressure drop [kPa] 0.6-2.1Particle concentration 1 [mg/m3n] 2000-3000Particle concentration 2 [mg/m3n] 570-1100Gas flow [kg/h] 100-190Primary cyclone catch [g/h] 130-270Secondary cyclone catch [g/h] 60-110Exhaust dust flow fe/hl 8-14

The grade efficiency curves for the secondary cyclone are also prepared. Grade

efficiencies were calculated for each test period based on the exit dust sample and the

samples taken from the captured dusts. Figure 7 shows a typical grade efficiency curve

obtained and also the predicted grade efficiency determined by Leicht&Iicht model. The

overall collection efficiences are 1 - 3 units of percentage lower than predicted by the

model The measured cut sizes (50 % collection efficiency) of the secondary cyclone are

3.5 - 4.2 pm. The predicted cut sizes are in the range from 1.3 to 1.7 pm.

What is typical of all determined grade efficiency curves is the sudden improvement in

collection efficiency when the particle diameter is below 1,5-2 pm. Similar phenomena

have also been detected by other workers /l, 2/. This is probably caused by particle

aggregation promoted within the cyclone whereby fine particles, which would otherwise

pass through the cyclone uncollected, adhere to bigger particles and are then consequently

collected. However, the agglomerates are broken down and re-dispersed during the

particle size analysis. Also the fine particles can agglomerate with other small particles

thus forming big particles which are then captured by the cyclones. It seems that the

behaviour of small particles at high pressures cannot be predicted by the Leicht&Licht

model.

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- 261 -

The comparison of the results obtained by the impactor and Coulter Counter, however,

suggests that after the cyclones there are no such agglomerates. The agglomerates might

have been broken during the pressure reduction process of the impactor measurement but

to this extent it, however, seems unlikely. The whole question of agglomeration needs w rr

further investigations. The research is to be extended to include electron microscopy

studies and chemical analyses of the collected samples.

COLLECTION EFFICIENCY OF THE SECONDARY CYCLONE

GRADE EFFICIENCY [%1

p- 10 barv-11.5 m/s'

-o- TEST 1 LEICHT&LICHT MODEL

Figure 7. A typical grade efficiency curve for the secondary

cyclone.

4 CONCLUSIONS

The very first measurements and results of particle emissions and cyclone performance

when burning peat at high pressures were obtained at the PFBC/G test rig during the test

period in autumn 1987. The particle concentration was measured using a total filter

system and size distribution was determined by virtual impactor and Coulter Counter

analyses.

The elutriation from the reactor varied from 2000 to 4000 mg/m^n and the two cyclones

reduced the concentration to 70 - 150 mg/m3n at the secondary cyclone outlet. The

overall efficiency of the Stairmand type secondary cyclone was 90 % being 2 units of

I _E 1

T T 1_ -it- -ii. ~ -i

I

w

- 262 -

percentage lower than predicted by the Leicht&Licht model. The results, however,

showed that the Leicht&Licht model appears to be insufficient for predicting the

behaviour of very fine particles. Further investigation of agglomeration of fine particles isneeded. Electron microscopy studies of the particles collected will be undertaken. f" Jr

References:

1. Pillai KX., Wood MA., Emissions from a pressurized fluidized bed coal

combustor. Journal of the Institute of Energy. December 1980. p 159-175.

2. Wheeldon J.M.,etxil., The performance of cyclones in the off-gas path of a

pressurized fluidized bed combustor. IChemE Symposium Series No.99. 1986.

3. Air pollution measurement, Bird&Tole brochure.

T T

S l

- 263 -

THERMODYNAMIC STUDY OF ALKALI METALS RELEASE IN PRESSURISED

FLUIDISED-BED COMBUSTION AND GASIFICATION OF PEAT

Wahab Mojtahedi* & Rainer Backman **

* Laboratory of Fuel Processing and Lubrication Technology

Technical Research Centre of Finland

Espoo, Finland.

** Department of Chemical Engineering

Abo Akademi

Turku, Finland.

ABSTRACT

A combined-cycle power generation system incorporating pressurised

fluidised-bed combustion (PFBC) or gasification is considered a promi­

sing approach for electricity generation using solid fuels such as

peat. In these systems, the high-pressure hot flue gas is expanded in a

gas turbine. Peat contains sodium and potassium which are released in

combustion and gasification. These are corrosive elements that can

cause severe damage to the turbine blades if not suppressed.

Multicomponent, multiphase equilibrium calculations were carried out

for atmospheric and pressurised fluidised-bed operating conditions to

determine the relative distribution of the two metals (Na and K) in the

gas and condensed phases. Dependence of the alkali volatilisation on

the operating temperature, pressure, the chlorine-content and the total

alkali-content of the feedstock was studied. The results show that the

alkali release in the vapour-phase could be much higher than acceptable

to a gas turbine, particularly under gasification conditions. Hence the

necessity to remove the volatilised alkali-metal compounds is more

acute in gasification than in combustion.

Both sodium and potassium are present as chlorides and to a lesser

extent as hydroxides in the gas phase in both modes of operation (i.e.

combustion and gasification). However, whereas under combustion

conditions both metals seem to condense as sulphates (Na2S04 and

]

I r 1

wi

264 -

K2S04), in gasification, chlorides and carbonates dominate in the

condensed phase. The alkali-metals volatilisation shows strong de­

pendence on the operating pressure of the system as well as on the

chlorine-content of the feedstock. It decreases markedly with the

former but increases sharply with the latter.

1 INTRODUCTION

Fluidised-bed gasification of peat to produce a low calorific value

fuel gas for use in a combined-cycle power generation scheme, provides

a great potential for improved thermal conversion efficiency. Similar­

ly, in a pressurised fluidised-bed combustion/gas turbine combined-

cycle system, the high temperature, high-pressure flue gas generated

from the PFB combustor is expanded through a gas turbine to recover

energy for improved overall thermal cycle efficiency.

The major obstacle remaining in the utilisation of solid fuels such as

coal and peat in such systems, is the high-temperature high-pressure

gas cleanup necessary before the flue gas can be admitted to the gas

turbine. The principal contaminants include

(I) fine particulates (small ash, sorbent and unburnt carbon

particles),

(II) alkali-metal compounds (principally sodium and potassium

chlorides and sodium and potassium sulphates), and

(III) sulphur dioxide and sulphur trioxide.

The contaminants are not effectively removed from the hot gases by

conventional hot gas cleaning methods (mainly cyclones) either because

of their very small size or because they exist in the vapour phase. The

fine particulates in the hot gas streams cause most of the erosion and

fouling that occurs in gas turbines. The major causes of hot corrosion

are the alkali-metal compounds. These compounds generally corrode the

turbine blades by attacking and destroying the protective oxide layer

normally present on the metal at high temperatures. The different tem­

perature ranges in which accelerated corrosion takes place can be rela­

ted to the different melting temperatures of various alkali compounds.

Manufacturers' specifications for acceptable levels of alkali in tur­

bine expansion gases vary reflecting a real lack of experience in the

effect of alkalis on turbomachinary /!/. The current industrial gas-

turbine specification limit for alkali-metal compounds in the com­

bustion gas entering a turbine is equivalent to 0.024 ppm (v) /2/.

Singh et al. /3/ reporting on alkali removal from hot coal-derived

gases, state that "conventional standards for turbomachinery require

alkali levels equivalent to typically 0.1 ppm (w) in a gaseous fuel".

As a result of the catastrophic erosion and corrosion of gas turbine

hardware, it is imperative to find a method to clean up gasifier pro­

duct gas and combustor flue gas before combined-cycle power generation

can be operated successfully. The removal of particulates and alkali-

metal compounds at high temperatures is investigated in the Laboratory

of Fuel Processing and Lubrication Technology of the Technical Research

Centre of Finland. This article deals with some theoretical considera­

tions concerning alkali release from fluidised- bed combustion and

gasification of peat.

2 THERMODYNAMIC EQUILIBRIUM CALCULATIONS

A theoretical distribution of sodium and potassium in combustion and

gasification products was studied assuming equilibrium conditions.

The equilibrium composition was established by minimising the total

Gibbs free energy of the system. This was done by using a computer

model developed by G. Eriksson /4/. Altogether 65 chemical species (45

gaseous and 20 condensed) were considered. These are listed in Table

(1). Thermodynamic data for the species were taken from references /5 -

8/. The notations used in Table 1 (and in this paper) are as follows:

- name-g gaseous

- name-1 1iquid

- name-s solid

condensed (i.e. solid or liquid)name-c

Table 1. Chemical species considered in the equilibrium calculations

1 6 10 bar

A: 0.3 & 1.2 bar

500-1000‘C

200-600 ppm

190-5700 ppmK+Na=

Gaseous phase: co2, h2o, o2, n2, oh. h2, CH^ HCO, H, 0. N.

HCN, C, NH3, HC1, Cl, C12, S03, S02, S, S2,H2S, COS, HS, CS, CS2, S20, H2S0l4, Ma, Na2,

NaCl, Na2Cl2, NaOH, NaH, NaO, Na2S0„, K, K2,

kci, k2ci2, koh, kh, ko

Liquid phase: Na2C03, Na2S04, Na2S, NaCl, NaOH, K2C03,

K2S0^, k2s, KCI, KOH

Solid phases: Na2 C03, Na2S04, Na2S, NaCl, NaOH, K2C03,

K2s°4, K2S, KCI, KOH

1 i r...i 1 hi

r i h-« M

267

In the results presented here, equilibrium compositions are calculated

assuming that gas mixtures are ideal and that all solid condensed

phases are immiscible. The first assumption is not a significant

limitation for the moderate pressures of interest here. Mixtures of

condensed liquid phases, however, have been considered except in the

case of sodium and potassium sulphates and carbonates.

The calculations were carried out to predict the form in which alkali

metals are released from xombustion or gasification of peat. It is of

interest to know whether the compounds of sodium and potassium will be

present in the aerosol or the vapour phase in the hot flue gas.

According to the vapour pressure data for sodium sulphate, if the

alkalis are present as sulphates in the hot gas streams, they are most

likely to be in the aerosol phase. If they are present primarily as

chlorides, they will be in the vapour phase. The actual form will have

a serious effect on the ultimate design of an industrial-sized hot gas

alkali removal system.

Table 2 lists the composition of the fuel used for the equilibrium

calculations. These were taken from reference /9/.

i r

r r.A

Table 2. Feedstock composition.

Element mol/1000 mol of C t/t of dry solid

C 1000.0 0.55H 1191.6 0.0550 444.98 0.326N 26.512 0.017S 1.363 0.002h2o* 808.54 0.667

El ement Minimum Maximummg/kg of peat mg/kg of peat

Cl 200 600K 140 2210Na 50 790

* The value for H20 corresponds to 40 % water-content in peat.

- 268 -

3 COMBUSTION

Equilibrium calculations were carried out for atmospheric as well as

pressurised fluidised-bed combustion (AFBC and PFBC) conditions. An air

ratio of 1.2.(i.e. 20 % in excess of the stoichiometric requirement)

with both operating pressures (of 1 and 10 bar) were employed. Some

calculations were also carried out to simulate isothermal fluidised-bed

combustion of peat. In this case, a temperature of 850 °c was employed

which represents a typical fluidised-bed combustion operation.

The calculations were repeated for three different fuel compositions

with regard to the chlorine content of the fuel (200, 400 and 600 mgAg

of peat) and again, for three alkali-metal contents in the feedstock.

Figures 1-5 show the results of some of these calculations. Figure 1

presents the distribution of the two metals in the three different

phases at 500 - 1000 °C temperature range. A chlorine content of 400

mg/kg and alkali content of 500 mg/kg (K = 400 and Na s 100 mg/kg),

used in these calculations, correspond to a "mean" peat composition

with regard to these elements.

The maximum of 100 % on the Y-axes corresponds to the total amount of

the element in question, in the system. Table 2, (e.g. 100 mg/Kg of dry

peat for Na). The composition of the gas can be read from Figure 2,

which shows the variation of the partial pressure of the gaseous alkali

and chlorine compounds with temperature. The bulk of both metals are

present in the form of chloride and to a lesser extent hydroxide in the

volatilised phase.

The liquid phase (Figure 1) consists almost entirely of the sulphates

of sodium and potassium (Na2S04-1 and K2S04-1). Solid sodium sulphate

seems to form below 740 °C but K2S04-s could be present at even 900 °C.

The concentration of both metals in the flue gas can be read from Table

3, which shows that the total alkali content in the vapour phase never

exceeds 0.05 ppm (v), (Na = 0.1, K = 0.04) at 800 °C under PFBC condi­

tions with a chlorine-content of 400 mg/kg and alkali-content of 500

mg/kg in the feedstock. This value represents about 0.06 * of the

JL

r r

Distribution of K

LIQUID

500 600 700 000 900 1000 °C

Temperature

Distribution of Na

Fuel

Pressure

X

Temp.

K + Na

Cl

PEAT

10 bar

1.2

variable

400 + 100 ppm (wt.)

400 ppm (wt.)

Distribution of Cl

LIQUID

500 600 700 800 900 1000 °C 500 600 700 GOO 900 1000 °C

Temperature Temperature

Hgure 1. Distribution of Na, K and Cl in three different, phases under PFB combustion conditions (fuel: peat).

270 -

sodium and 0.12 % of the potassium in the system. The corresponding figure for 900 °C shows a value of 0.67 ppm (v) for the total volati­

lised alkali (Na = 0.13, K = 0.5), an increase of almost 10-fold.

This figure represents about 1.0 %> of the total sodium and 1.9 % of

the total potassium entering the system.

GASEOUS K-, Na- and Cl-species

-0.001

900 1000 °C700 BOO Temperature

i r

f TJL A

IE

Figure 2. Variation of the partial pressure of the gaseous alkalispecies with temperature under PFB combustion conditions(fuel: peat).

- 271 -

The volatilised alkali metals are most probably initially present as

hydroxide or chloride vapours /10/. They then react with hydrogen

chloride or sulphur oxide and o*ygene in stages, to form species which

eventually result in deposits of sodium and potassium carbonates,

sulphites, but mainly chlorides and sulphates. It has been a common

practice to show the formation of these compounds by net chemical

reactions where a number of elementary reaction steps are involved. A

few of such net reactions could be /10/:

Cl = 600 mg/kg

Cl = 400 mg/kg

Cl = 200 mg/kg

Temperature C

Figure 3. Variation of the vapour-phase Na concentration in the fluegas with temperature for three different Cl contents. PFBcombustion conditions (fuel: peat).

r r

v rJL -J.

\

ppm

(v) K i

n the v

apou

r phas

e

272 -

KOH (g) + HC1 (g) - KCI (g) + H20 (g) (1)

2 NaCl (g) + H20 (g) + S02 (g) + 1/2 02 (g) -

Na2S04 (c,g) + 2 HC1 (g) (2)

2 NaOH (g) + S02 (g) + 1/2 02 (g)

Na2 SO^ (c,g) + H20 (g) (3)

C! = 600 mg/kg

Cl = 400 mg/kg

Cl = 200 mg/kg

800 850 900 950 1000

Temperature °C

Figure 4. Variation of the vapour-phase K concentration in the fluegas with temperature for three different Cl contents. PFBcombustion conditions (fuel: peat).

W- _

I E .

- 273 -

GASEOUS K-, Na- and Cl-species

100

10

1

>

EQ.CL

0.1

0.01

0.001

o 1000 2000 3000

(K+Na)-content/ mg/Kg of peat

Figure 5. Variation of the partial pressure of gaseous alkali species with the feedstock alkali-content. PFB combustion conditions, (P = 10 bar, X = 1.2, T = 850 °C, Cl-content = 200 ppm (wt.) fuel:peat)).

r r i- _*4 J

I T 1i

T fp

F

274 -

Table 3. Sodium and potassium concentrations 1n the vapour phase In the flue gas under combustion conditions (x = 1.2) at 1 and 10 bar pressures. (Fuel : peat. Cl-content = 400 mg/kg).

Temperature(°C)

P = 1 bar P = 10 bar

Na(ppm v)

K(ppm v)

Na + K (ppm v)

Na(ppm v)

K(ppm v)

Na + K (ppm v)

800 0.12 0.56 0.68 0.01 0.04 0.05810 0.16 0.75 0.91 0.01 0.05 0.06820 0.22 1.00 1.22 0.01 0.06 0.07830 0.29 1.32 1.61 0.02 0.09 0.11840 0.39 1.72 2.11 0.03 0.11 0.14850 0.51 2.23 2.74 0.03 0.15 0.18860 0.66 2.85 3.51 0.04 0.19 0.23870 0.85 3.61 4.46 0.06 0.25 0.31880 1.08 4.51 5.59 0.08 0.33 0.41890 1.36 5.56 6.92 0.10 0.42 0.52900 1.68 6.74 8.42 0.13 0.54 0.67910 2.05 8.06 10.11 0.17 r\ c -\.: - . 0.86920 2.46 9.48 11.94 0.21 1.09930 2.92 10.97 13.89 0.2? -ii 1.38940 3.41 12.51 15.92 0.34 1.39 1.73950 3.92 14.06 17.98 0.43 1.73 2.16960 4.47 15.61 20.08 0.53 2.14 2.67970 5.04 17.13 22.17 0.66 2.62 3.28980 5.63 18.64 24.27 0.82 3.20 4.02990 6.26 20.15 26.41 1.00 3.87 4.87

1000 6.95 21.64 28.64 1.21 4.64 5.85

However, in thermodynamic equilibrium analyses such as those reported

in this paper, it is more appropriate to consider the chemical system

as a whole rather than single reactions. Hence, the compounds which can

exist at equilibrium under given conditions of temperature, pressure

and composition are considered in our analyses and not these reactions.

This study shows that both sulphates of sodium and potassium are likely

to be found in the condensed phase under the conditions prevalent in

PFBC systems. (Na2S04-g concentration never exceeds 0.01 ppmv even at

950 °C). Scandrett /10/ studying the possible forms of the alkali-metal

compounds in pressurised fluidised-bed combustion of coal, also

reported that both sodium sulphate and potassium sulphate are most

likely to be present in the condensed phase.

S l

- 275 -

The saturation concentrations of the chlorides of sodium and potassium

at 900 °C and a pressure of 10 bar are reported to be of the order of

10"3 bar in a flue gas with similar composition to that considered here

/10/. Our calculations show the concentrations of NaCl-g and KCL-g to

be well below the saturation levels (of the order of 10"5 for both

chlorides). Therefore, if the alkalis are present as chloride or

hydroxide, they will be present in the vapour phase. However, the

estimated vapour saturation concentrations for the sulphates are below

their concentrations in the flue gas. This suggests that, if the

alkalis are present as sulphates, they will be present predominantly as

condensed phase.

It is interesting to note the effect of the chlorine-content of the

fuel on the relative distribution of the two metals in the gas and

condensed phases. Most of the fuel-bound chlorine is present as HCl-g

in the flue gas. Figure 2. Because the HCl-g concentration is much

higher than that of the alkali chlorides, its formation has negligible

effect upon the chloride concentration in the gas. However, Figures 3

and 4 show the volatilisation of both sodium and potassium to be

strongly dependent on the chlorine-content of the in-going fuel. The

higher the Cl-content the higher the vapour-phase alkali chlorides in

the flue gas. This indicates the relative importance of the peat compo­

sition used as fuel in pressurised fluidised-bed combustion process.

The flue gas composition with respect to the volatilised alkali com­

pounds (principally chlorides), the necessity to remove these com­

pounds, and the removal mechanism selected, (should it be required),

would depend to some extent on the type of peat combusted, (i.e. the

initial feedstock composition).

Calculations similar to those reported above, were carried out for

operating conditions corresponding to fluidised-bed combustion of peat

at normal pressure; the operating conditions were identical except for

pressure which is one bar. Again, the higher chlorine content of the

fuel enhances alkali metals volatilisation (in the form of vapour-phase

chlorides). A comparison of the results obtained at P = 1 bar with

r t

t r

- 276 -

P = 10 bar (Table 3), shows the effect of increasing the operating

pressure of the system on the concentration of alkali metals released

in the gas phase at equilibrium. The latter decreases markedly with

increasing pressure. At a temperature of 850 °C, the equilibrium

concentration of Na in the gas-phase increases from almost 0.03 ppm (v)

to about 0.5 ppm (v) when the pressure is reduced to 1 bar, (chlorine-

content = 400 mg/kg of peat, (Na + K)-content = 500 mg/kg). Table 3.

Similar calculations with potassium show an increase of more than

10-fold (0.15 ppm (v) to 2.2 ppm (v)) in the concentration of

volatilised potassium released into the flue gas. Table 3.

Figure 5 shows the dependence of the alkali-metal compounds volatilisa­

tion on the alkali-content of the feedstock between 500 and 3 000 mg/kg.

These calculations were carried out for two different chlorine-contents

in the peat. Figure 5 corresponding to 200 mg/kg. Within this range,

all the major alkali (Na and K) compounds (i.e. chlorides and

hydroxides) show a steady monotonous increase with increasing alkali-

content. The weight ratio of the two metals (K/Na) was kept constant at

2.8 as well as the temperature of 850 C and the sulphur content (0.2 % wt.) of the feedstock. The curve corresponding to NajSO^ shows no

change of gradient which would seem to indicate a sufficient sulphur

supply for condensed-phase sulphate formation at equilibrium.

4 GASIFICATION

Thermodynamic equilibrium calculations were carried out for the opera­

ting conditions corresponding to atmospheric as well as pressurised

fluidised-bed gasification of peat (an air ratio of 0.3, a temperature

of 850 °C and two operating pressures of 1 and 10 bars, the fuel compo­

sition as in Table 2). Figures 6-9 show the results of some of these

calculations. In fuel rich conditions, sulphur oxides are present at

extremely low concentrations. Most of the sulphur in the feedstock

forms hydrogen sulphide H2S-g and the alkali sulphates will not form.

II

T ' T

- 277 -

Compared to fluidised-bed combustion, a much greaterproportion of both

metals (Na, K) are present in the vapour phase at 700 - 900 °C tempe­

rature range (Figures 1 and 6). The volatilised phase consists prima­

rily of chlorides and hydroxides of sodium and potassium. Whereas in

F-B combustion the condensed phase consists almost entirely of the

sulphates of the two metals, in F-B gasification both metals condense

as chlorides and carbonates. KC1-1 is the dominant compound of potas­sium in the condensed phase above 600 °C. Below this temperature, the

solid phase forms and KCl-s is the dominant compound of potassium. A

relative distribution of potassium between KC1-1 and K2C03-1 was cal­

culated and the results show the concentration of the former to be al­

most an order of magnitude greater than the latter. This pattern pre­vails up to almost 920 °C, above which no condensed phase is revealed.

A similar pattern is revealed with sodium in that, at equilibrium, this metal will condense as NaCl-1 and NagCO^-l at 600 - 900 °C tempe­

rature range. However, the relative distribution of sodium in these

two compounds varies as the temperature increases. Most of the metal is bound as NaCl-1 at about 700 °C but the carbonate percentage becomes

significant above 850 °C.

Figure 7 shows that the chlorides (KCl-g and NaCl-g) and to a lesser

extent hydroxides (KOH-g and NaOH-g) are the dominant volatilised al­

kali compounds. Diehlorides of both metals (I^Clg-g and Na2Cl2-g) are

also formed but by almost an order of magnitude less than monochlorides.

The concentrations of the volatilised alkali metals in the flue gas at various temperatures are tabulated in Table 4. Even at 700 °C a rela­

tively significant concentration (at ppm level) of both metals is re­

leased in the vapour phase in the flue gas (7.5 ppm (v), with Na =

0.6 ppm (v) and K = 6.9 ppm(v)), Table 4.

Our analyses show that, depending on the alkali and chlorine content

in the feedstock, as much as 20 % of total sodium and 70 % of total potassium in the fuel could be present in the vapour phase at 800 °C,

under gasification conditions. Figure 8. Again, the higher chlorine

content in the fuel seems to enhance the alkali-metal volatilisation.

TF-L. - ]

T J

El

wi

Distribution of K

KCI-s LIQUID

900 1000Temperature

Distribution of Na

Fuel

Pressure

X

Temp.

K + Na

Cl

PEAT

10 bar

0.3

variable

400 + 100 ppm (wt.)

400 ppm (wt.)

Distribution of Cl

1000 °C: 500 600 700 800 900Temperature

Figure 6. Distribution of Na, K and Cl in three different phases under PFB gasification conditions (fuelrpeat).

I

I

\ 4 lx*/ Mi N hi•

r I H 1

\ t

I

I

P

278

- 279 -

GASEOUS K-, Na- and Cl-species

100

10

1

0.1

o.oi

o.ooi

>

CLCL

500 500 700 800 900 1000 °CTemperature

Figure 7. Variation of the partial pressure of the gaseous alkalispecies with temperature under PFB gasification conditions (fuel: peat).

r w i___j. . j

i _£

T T

The variation of the partial pressure of the volatilised alkali and

chlorine compounds with the alkali content of the feedstock is pre-w

sented in Figure 9. Within the range investigated (500 - 3 000 mg/kg ]_

of dry peat), the chlorides of both sodium and potassium (KCl-g,

K2Cl2-g, NaCl-g, Na2Cl2-g) show a different pattern of dependence on

the total alkali content of the fuel than those of hydroxides. KOH-g

and NaOH-g concentrations show a very slight increase with increasing

Potassium

in v

apour

Cl-content: 200 g/t ds.400BOO

1 Pressure - 10 bar X - 0.3

Potasslum

> 60

1000 °C 1000 °CTemperature Temperature

Figure 8. Percent of total Na and K in the vapour phase for three different Cl contents in the feedstock. PFB gasification conditions (P - 10 bar, x - 0.3, fuel: peat).

WA

I £

wJL

r

(Na + K) content after a threshold value of about 500 mg/kg, or remain

constant. The chloride concentrations on the other hand, show the oppo­

site trend of steady decline with increasing (Na + K) content after

the threshold value of about 500 mg/kg. In these calculations the K/Na

weight ratio of 2.8, the sulphur content of peat (0.2 % wt.), the

chlorine content of the feedstock (of 200 mg/kg) and the operating

temperature of 850 °C are kept constant. Similar calculations with a

higher chlorine content (of 600 mg/kg in the feedstock) show very

little change in the overall pattern seen in Figure 9. This seems to

indicate that above a certain minimum requirement of total alkali

presence in the system (necessary for the chloride and hydroxide

formation), the concentrations of the volatilised alkali compounds

are not significantly affected by an increase in the total (Na + K)

content of peat, within the range investigated.

- 281 -

GASEOUS K-, Na- AND Cl-species

Figure 9. Variation of the partial pressure of gaseous alkali species with the feedstock alkali-content.PFB gasification conditions (P = 10 bar, X = 0.3,T = 850 °C, Cl-content = 200 ppm (wt.) fuel:peat ).

4

- 282 -

Table 4. Na and K concentration In the vapour-phase In the product gas under gasification conditions (\ = 0.3). (Fuel : peat, Cl-content s 400 mg/kg).

Temperature(°C)

P = 1 bar P = 10 bar

Na(ppm v)

K(ppm v)

Na K (ppm v)

Na(ppm v)

K(ppm v)

Na + K (ppm v)

700 7.31 49.98 57.29 0.62 6.92 7.54710 7.80 51.12 58.92 0.86 9.02 9.88720 8.36 52.21 60.57 1.12 11.51 12.63730 9.03 53.27 62.30 1.41 14.08 15.49740 9.84 54.28 64.12 1.76 17.05 18.81750 10.82 55.23 66.05 2.18 20.41 22.59760 12.02 56.11 58.13 2.70 24.09 26.79770 13.48 56.91 70.39 3.30 27.98 31.28780 15.27 57.62 72.89 4.00 31.92 35.92790 17.43 58.24 75.67 4.78 35.73 40.51800 20.04 58.76 78.80 5.61 39.24 44.85810 23.17 59.19 82.36 6.49 42.38 48.87820 25.38 59.48 84.86 7.41 45.11 52.52830 25.43 59.62 85.05 8.40 47.50 55.90840 25.48 59.75 85.23 9.46 49.57 59.03850 25.52 59.85 85.37 10.62 51.38 62.00860 25.55 59.94 85.49 11.91 52.96 64.87870 25.57 60.01 85.58 13.35 54.33 67.68880 25.59 60.07 85.66 14.97 55.52 70.49890 25.61 60.12 85.73 16.78 56.53 73.31900 25.62 60.15 85.77 18.80 57.39 76.19

CONCLUSIONS

The global thermodynamic equilibrium analysis reported in this paper

was undertaken to identify the major alkali (Na and K) compounds in

fluidised-bed combustor and gasifier effluent streams. It was also

intended to study the effect of temperature, pressure, the chlorine

content and the alkali content of the feedstock on the total volati­

lised alkali content of the flue gas. The interactions of the two

metals with the fly ash constituents (S102, CaO, FeO) were not con­

sidered in our "worst-case" projections. There is experimental evi­

dence that some of the volatilised alkali-metal compounds will con­

dense onto, be adsorbed by or interact with the fly ash particles

and hence the measured alkali concentration in the hot flue gas could

be much lower than these calculations suggest.

- 283 -

The removal of the alkali-metal compounds from the hot flue gas will

depend to some extent on the form in which these species will exit

from the reactor in the gas stream. In other words, the necessity to

remove these compounds and the choice of the removal device will de­

pend on the form in which the alkali metals exist in the hot flue gas.

It was shown that the alkali metals will be present predominantly in

the vapour phase, should they form chlorides or hydroxides, but as

aerosols, if present as sulphates.

Our analyses of the alkali-metal release under conditions prevalent

in fluidised-bed combustion and gasification of peat show that:

The concentration of the vapour-phase alkali-metal compounds in the

flue gas is strongly dependent on the operating temperature and

pressure. It increases sharply with increasing temperature but

decreases markedly with an increase in the total pressure of the

system. In the temperature range of interest (700 - 1 000 °C), the

dominant species of both sodium and potassium are chlorides

(NaCl-g), KCl-g) and hydroxides, (NaOH-g and KOH-g). Whereas in

combustion conditions, the condensed phase consists almost entirely

of sulphates (Na2S04-c and K2S04-c), under gasification conditions

chlorides (NaCl-c and KCl-c) and carbonates (Na2C03-c and K2C03-c)

prevail.

Relatively, a greater percentage of both metals are released in the

vapour phase in gasification than in combustion. Hence, gettering

(to remove the volatilised alkalis) is more relevant to

gasification than combustion.

A strong dependence on the chlorine-content of the peat is also

revealed. Relatively higher concentrations of volatilised alkali-

metal compounds are revealed with higher chlorine-content in the

feedstock. This holds for both modes of operation (i.e. combustion

and gasification).

The total alkali-content of peat does not seem to affect the alkali

metals volatilisation greatly. Above a certain minimum requirement

- 284 -

(necessary for the alkali' compounds formation) the presence of

additional Na and K simply results in a slight increase in the

volatilised alkali-metal compounds under PFBC conditions, the

chlorine-content and the sulphur-content of the fuel being suffi- _

cient to form the chlorides (of Na and K) as well as the condensed —a

phase sulphates. However, under PFB gasification conditions, the

reverse obtains and the alkali chloride concentrations decline with

increasing (Na + K) content above this threshold value.

REFERENCES

1. Scandrett, L. A. & Clift, R. Thermodynamics of alkali removal from

coal-derived gases. Journal of the Inst, of Energy (1984) Dec., p.

391 - 397.

2. Lee, S. H. D. & Miles, K. Measurement of alkali vapour in PFBC

flue gas and its control by a fixed granular bed of activated

bauxite. Proc. Symp. on Gas Cleaning at High Temperatures, 16 - 18

Sept, 1986, p. 149 - 166.

3. Singh, A. et al. Thermodynamic calculations of the effects of

chlorine on alkali removal from hot coal-derived gases. Proc. Symp.

on Gas Cleaning at High Temperatures, 16 - 18 Sept. 1986, p. 167 -

176.

4. Eriksson, G. Thermodynamic studies of high temperature equilibria.

Chemica Scripta (1975)8, p. 100 - 103.

5. Barin, I. & Knacke, 0. Thermochemical properties of inorganic

substances. Berlin 1973, Springer-Verlag. 921 p.

6. Barin, I., Knacke, 0. & Kubaschewski, 0. Thermochemical properties

of Inorganic substances. Supplement. Berlin 1977, Springer-Verlag.

861 p.

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w{

- 285 -

7. JANAF thermochemical tables with supplements. Washington, D.C.,

1971 - 1978, National Bureau of Standards of the U.S. Department of

Commerce. 1 141 p.

8. Pejryd, L. & Hupa, M. Bed and furnace gas composition in recovery

boilers. Paper presented at the 1984 Tappi Pulping Conf., November

12 - 14, San Francisco, U.S.A., 12 s.

9. Lappi, M. Branntorvmyrarnas fbrbranningstekniska och kemiska egen-

skaper samt asksammansattning (Combustion-technical and chemical

• characteristics of mires used for fuel peat production, and the

composition of ash). Esbo 1985, Technical Research Centre of Fin­

land, NE 1460 76 SVEFIN TORV, Properties of industrialpeat and

their test methods. Sub-report 2. 63 p.

10. Scandrett, L. A. The removal of alkali compounds from gases at

high temperatures. Ph. 0. thesis. Cambridge 1983, University of

Cambridge.

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APPENDIX 1/1

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Participants of PFBC/G-SeminarBoij Johan Helsinki University of Technology

/ EVO-laboraioryOtakaari 4 02150 Espoo

Clift Roland University of Surrey England

Ernsten Svante Suomen Brown Boveri Oy PI 10 00721 Helsinki

Fagerholm Nils-Erik Helsinki University of Technology / Power plant machinery

SihkOmiehemie 4 02150 Espoo

Frilund Bjarne ASEA-PFBC Ab S-612 20 'Finspong Sweden

Haasiosalo Taisto Neste Oy Keilaniemi 02150 Espoo

Hanelius Antti Suomen Voimalaitosyhdistys r.y. LOnnrotinkatu 4 B 00120 Helsinki

Hannus Seppo Ministry of Trade and Industry / Energy Dept.

PI 37 00131 Helsinki

Heinonen Olli A. AhlstrOm Oy / Research Dept PI 18 48601 Karhula

Hokajarvi Risto Neste Oy / Energy technology Lundinkatu 6-8 06100 Porvoo

Holm Suvi Tampere University of Technology / SMtOtekniikka

PI 527 33101 Tampere

Horn Lars Helsinki Energy Board P1469 Helsinki 00101

Horvath Andras Helsinki University of Technology / EVO-laboratory

Otakaari 4 02150 Espoo

Hoy Raymond Hoy Associates England

HuHdconen Seppo Helsinki University of Technology / EVO-laboratory

Otakaari 4 02150 Espoo

HwMtari Juba Technical Research Centre ef Finland / KPA

PI 221 40101 JyviskyU

Hupa Mikko Abo Akedrmi Piispankuja 8 20500 Turku

HyntyJJaav® ^ Tampdia / Boiler Division PO:BOX 626 33101 Tampere

Hamali Ap^ero Taaapeiia Oy PI 256 33101 Tampere

Hamala Sirpa Tampella Oy PI 256 33101 Tampere

Isaksson Juhani A. AUstrtm Oy / Research Dept PI 18 48601 Karhula

Isoniemi Markku Helsinki University of Technology / EVO-Uboralory

Otakaari 4 02150 Espoo

Jahkola Antero Helsinki University of Technology / EVO-laboratory

Otakaari 4 02150 Espoo

Janka Pentti Tampella / Boiler Division P:0:B0X 626 33101 Tampere

Kaukanen Eila Enconsults Oy / SF-Cleanair PI 50 20781 Kaarina

Keranen Hannu Neste Oy Keilaniemi 02150 Espoo

Kinni Jouni Tampella / Boiler Division P:0:B0X 626 33101 Tampere

C

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APPENDIX 1/2

Participants of PFBC/G-SeminarKiukaanniemi Eino Oulu University of Technology Lirmamnaa 90570 Oulu

Korkiakoski Martti Ministry of Trade and Industry / Pohjoinen 00130 HelsinkiEnergy Dept. Makasiinikatu 6

Kuorelahti Kalevi Kemirm Oy

Kurkela Esa Technical Research Centre of Finland / POV

Biologinkuja S 02150 Espoo

Kurki-Suonio Ilmari Helsinki University of Technology Otakaari 4 02150 Espoo

Kurppa Erkki Limpdvoima Oy Mertaniemi 53920 Lappeenranta

Lehtonen Olof Tampella / Boiler Division P:0:B0X 626 33101 Tampere

LVK / Jarmo Tervo Helsinki University of Technology / EVO-laboratory

Otakaari 4 02150 Espoo

Maskuniitty Hannu Ekono Oy PL 27 00131 Helsinki

McKeough Paterson Technical Research Centre of Finland / POV

Biologinkuja S 02150 Espoo

Mikkanen Pirita Helsinki University of Technology / EVO-laboratory

Otakaari 4 02150 Espoo

Mojtahedi Wahab Technical Research Centre of Finland/POV

Muller Rainer KWU West Germany

Nieminen Jorma A. AhlstrOm Oy / Boiler division PI 184 78201 Varicaus

Nieminen Matti Technical Research Centre of Finland/POV

Biologirikuja S 02150 Espoo

Nitschke Eberhard UHDEGmbH West Germany

Nurmi Jari Eknnalik Oy Vihiheikkilintie 37 20810 Turku. . 3 ^

Olsen Aksel Risoe National Laboratory 4000 Roskilde Demhink

Orjaia Markku Technical Research Centre of Finland/KPA

PI 221 40101 JyviskyU

Pesonen Pekka Technical Research Centre ofFinland / KPA

PI 221 40l01 JyviskyU

Pikku-Pyhilto Raija Ministry of Trade and Industry / Pohjoinen 00130 HelsinkiEnergy Dept Makasiinikatu 6

Pillai Krisna ASEA-PFBC Ab Sweden

Raiko Markku Imatran Voima Oy PI 138 00101 Helsinki

Raiko Risto Tampere University of Technology PI 527 33101 Tampere

Riala Pekka Imatran Voima Oy PI 138 00101 Helsinki

Rinne Risto University of Oulu / Energy laboratory

90570 Oulu

Salmi Jorma Jaakko FOyry Oy Kanpintie 3, PI 16 00441 Helsinki

X

APPENDIX 1/3

Participants of PFBC/G-SeminarSalminen Pekka Imatran Voima Oy PI 138 00101 Hebinki

Salo Kari YTT - Yhtymft Oy Hlmeenlinna workshop

PI 537 13111 Hlmeenlinna

Salo-Asikainen Sirpa Ministry of Environment PI 399 00121 Hebinki

Saviharju Kari Technical Research Centre ofFinland / POV

Seppa Matti Helsinki Energy Board PI 469 00101 Hebinki

Sihvola Ali Imatran Voima Oy PI 138 00101 Hebinki

Sipila Kai Technical Research Centre ofFinland / POV

Biologinkuja 3-5 02200 Espoo

Sipilp Kari Technical Research Centre ofFinland / SAH

Otakaari 5 i 02150 Espoo

Solantausta Yrjo Technical Research Centre ofFinland / POV

Suokas Mervi Tampella Oy PI 256 33101 Tampere

Thun Rabbe Technical Research Centre ofFinland / POV

Biologinkuja 5 02150 Espoo

Vainio Pirjo-Liisa Ministry of Trade and Industry / Pohjoinen 00130 HebinkiEnergy Dept. Makasiinikatu 6

Vessonen Kari Imatran Voima Oy PI 138 00101 Hebinki

Vilppunen Pekka University of Oulu / Energy laboratory

90570 Oulu

Virta Kari A. AhlstrOm Oy / Termoflow P.O. BOX 15 20781 Kaarina

Wallman Henrik Chalmers University of Technology Institutionen for S-41296 GOteborg,energiteknik Sweden

Weidner Ulrich Didier Filtertechnik Konrad-Adenauer-Str. 4 D-6719 Ebenberg, FRG

Wilin Carl Technical Research Centre ofFinland/POV

Biologinkuja 5 02150 Espoo

Ylhaisi Mikko Ministry of Trade and Industry / PI 37 00131 Hebinki

Astrdm LeifEnergy Dept.Outokumpu Oy / Engineering PtOtBOX 86 02201 Espoo

Xijalfi Martti Imatran Voima Oy PI 138 00101 Hebinki

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Published by Technical Research Centre of Finland Vuorimiehentie 5 SF-02150 Espoo, Finland

IT TV phone internal. + 358 0 4561 \l */ telex 122972 vtthasf

Name, number and report code of publicationVTT Symposium 83 FI+VTTSYMP-88/83

DateMarch 1988

Project numberPOVT00838

Authors Name of projectPFBC-seminaari

Ed. byMaija Korhonen

Commissioned by Helsinki University of Technology,Technical Research Centre of Finland (VTT), Laboratory of Fuel Processing and Lubricant Technology, Imatran Voima Oy

Title

PRESSURIZED FLUIDIZED BED COMBUSTION ANDGASIFICATION POWER SYSTEMS

Abstract

The Seminar on Pressurized Fluidized Bed Combustion and Gasification Power Systems was held in Espoo, Finland, on 23 - 24 November, 1987. The topics of the Seminar were related to a joint research project between die Technical Research Centre of Finland (VTT), the Helsinki University of Technology, and Imatran Voima Oy. The objective was to take a general review of the state-of-the-art of combined cycle power plant processes based chi pressurized combustion and gasification. Foreign specialists had been invited to lecture at the Seminar. Research done in Finland was also introduced.

The lectures given on the first day focused on potentials offered by pressurized combustion technology in Finland, and on research and development in the field of pressurized fluidized bed combustion. A visit was also organized to see die test rig constructed for research into pressurized fluidized bed combustion and gasification at the Laboratory of Fuel Processing of VTT. The lectures on the second day concerned gasification technology and gas cleanup from particulates and alkali metals.

Activity unitLaboratory of Fuel Processing and Lubricant Technology, Biologinkuja 3-5, SF-02150 Espoo, Finland

ISSN and key name0357-9387 VTT Symposim - Valtion teknillinen tutkimuskeskus

ISBN951-38-3105-1

LanguageEnglish

Class (UDC)

662.74:662.76:66.096.5

Key words

combustion, fluidized bed processors, pressurizing, gasification, coal, peat, electric power plants

Sold by Government Printing Centre P.O. Box 516 SF-00101 HELSINKI phone internal. + 358 0 56601

Pages

285 p. + app. 3 p.

PriceFIM95

Note

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VTT:N JULKAISUSARJATValtion teknillisen tutkimuskes-kuksen julkaisusarjat ovat

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