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A comparison of two different flow types on performance ofa thermally coupled recuperative reactor containing naphthareforming process and hydrogenation of nitrobenzene

D. Iranshahi, E. Pourazadi, A.M. Bahmanpour, M.R. Rahimpour*

Department of Chemical Engineering, School of Chemical and Petroleum Engineering, Shiraz University, Shiraz 71345, Iran

a r t i c l e i n f o

Article history:

Received 20 August 2010

Received in revised form

20 November 2010

Accepted 8 December 2010

Available online 13 January 2011

Keywords:

Recuperative coupling

Catalytic naphtha reforming

Hydrogen production

Aromatic enhancement

Nitrobenzene hydrogenation

* Corresponding author. Tel.: þ98 711 230307E-mail address: rahimpor@shirazu.ac.ir (

0360-3199/$ e see front matter Copyright ªdoi:10.1016/j.ijhydene.2010.12.033

a b s t r a c t

Refineries have been looking for proper ways of improving reformer performance by

enhancing the octane number of the product via increasing the aromatics’ compounds. To

reach this goal, the endothermic catalytic naphtha reforming is coupled with the

exothermic hydrogenation of nitrobenzene to aniline in a multifunctional heat exchanger

reactor through the process intensification concept. Considering the higher thermal effi-

ciency as well as the smaller size of the coupled reactor, utilizing this reactor is given

priority. In this novel configuration, the first and the second reactor of the conventional

naphtha reforming process are exchanged with the coupled reactors contain the endo-

thermic naphtha reforming in the shell side and the hydrogenation reaction in the tube

side. Both co-current and counter-current modes of flow are examined during the opera-

tion considering various studies in literature which show the superiority of co-current flow

compared with the counter-current flow. The result of current study is compared with the

corresponding results for conventional tubular reactor (CTR). The results show higher

aromatic production as much as 18.73% and 16.48% in the co-current and counter-current

mode, respectively. Hydrogen molar flow rate increases about 5 kmol/h by using counter-

current flow regime, compared with the CTR.

Copyright ª 2010, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights

reserved.

1. Introduction catalytic reformer. The process licensors have responded to

1.1. Catalytic naphtha reforming process

Thecatalytic reformingunithasan important role in refineries.

This process includes the reconstruction of low-octane

hydrocarbons in the naphtha stream and production of more

valuable high-octane gasoline components without changing

the boiling point range. Over the past decade, refiners have

been forced to maximize their existing asset utilization. With

capital at apremium, refinersmustdelivermore fromtheunits

they already have. These double pressures from environ-

mental regulations and asset utilization have influenced the

1; fax: þ98 711 6287294.M.R. Rahimpour).2010, Hydrogen Energy P

these needs [1]. As a consequence, several research efforts

have been made and some modifications have been proposed

to enhance the production of the aromatic and hydrogen, as

well as the operating conditions of the process.

Studies which have been done on the catalytic naphtha

reforming process are categorized as bellow:

1.1.1. Studies on the kinetics of catalytic naphtha reformingprocessThe naphtha feed stream is composed of various hydrocarbons

and related isomers which are normally influenced by various

reactions. Thus a detailed kinetic model considering all the

ublications, LLC. Published by Elsevier Ltd. All rights reserved.

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 53484

components and reactions is too complex [2]. Accordingly,

various attempts have been made to find more acceptable

lumpedgroupsof reactions to represent the catalytic reforming

reactions. In this way, the first significant attempt to model

a reformingsystemwasmadebySmith [3]whoconsideredonly

threebasic components innaphthamainstructure:Thesewere

paraffins, naphthenes, and aromatics industrially known as

PNA.Weifeng et al. [4] developed anewkineticmodel including

20 lumped components and 31 reactions. The subdivision of

8-carbon aromatics and all the hydro cracking reactions of the

paraffin lumps were taken into consideration in their model.

Boyas and Froment [5] studied a very detailed model for the

catalytic reforming of naphtha based on its fundamentals.

The model provided a thorough insight into the process and

a wide predictive potential but contained a huge number of

parameters. The current model also imposed equilibrium for

thehydrogenationanddehydrogenationreactions. Stijepovic et

al. [6] postulated a semi-empirical kinetic model for catalytic

reforming. Their lumping strategy was based on paraffins,

olefins, naphthenes and aromatics (PONA) analysis.

1.1.2. Studies on the modeling and improvement of aromaticsand hydrogen yields of the CTRsLi et al. [7] and Taskar and Riggs [8] modeled and optimized

a semi-regenerative catalytic naphtha reformer by considering

most of its key constitutes. Juarez et al. [9] modeled and simu-

lated four catalytic reactors in series for naphtha reforming.The

process model used an extended version of the kinetic model

reported by Krane et al. Weifeng et al. [2] considered 18-lumped

kinetic models to simulate and optimize a whole industrial

catalytic naphtha reforming process by Aspen Plus platform.

Khosravanipour and Rahimpour [10] presented a membrane

catalytic bed concept fornaphtha reforming inorder to enhance

hydrogen and aromatic production yield. Rahimpour et al.

[11e14] have presented novel reactor configurations and modi-

fications which lead to higher production rate of the desired

products. In the previous studies [14], the concept of using

thermally coupled reactors containing naphtha reforming

process and hydrogenation of nitrobenzene has been investi-

gated. In the present research, utilizing co-current and counter-

current flow regimes has been discussed and the related results

have been compared.

1.1.3. Studies on the catalyst and catalyst regenerationMore studies have been done previously on the new catalyst

preparation and catalyst deactivation due to the coke deposi-

tion and sintering. In this way Benitez et al. [15] and Boutzeloit

et al. [16] examined the reforming efficiency and product

Table 1 e Rate constants and heat of reactions fornaphtha reforming.

k ¼ AexpðB� E=1:8 TÞ A B E DH298 K

kf1 9.87 23.21 36,350 71,038.06

kf2 9.87 35.98 58,550 �36,953.33

kf3 1 42.97 63,800 �51,939.31

kf4 1 42.97 63,800 �56,597.54

Ke1 1.04 � 10�3 46.15 46,045 e

Ke2 9.87 �7.12 8000 e

enhancement through the catalyst with different metal basis.

Considering the coke deposition phenomenon on the catalyst

surface and regeneration of catalyst some researches have

been made, Mazzieri et al. [17] and Sugimoto et al. [18] inves-

tigated the coke formation and regeneration of catalyst during

the catalytic naphtha reforming process.

1.2. Catalytic hydrogenation of nitrobenzene process

1.2.1. Literatures for coupled reactorsOne of the most important advantages of using thermally

coupled reactors is to save energywhich consequently reduces

the capital and operational costs [19]. Considering the lower

capital costs due to the elimination of the inter stage heaters

and related piping system and operational costs due to lower

fuel consumption, as well as higher thermal efficiency, inves-

tigation of the thermally coupled reactors has been the first

priority in this research.Oneshouldconsider the fact that using

the presented configuration leads to combination of two

separate plantswhich indeed reduces the capital costs. Process

intensification (PI), known as the strategy of miniaturizing

process plants, is a novel design approach which can reduce

equipment size by several orders of magnitude leading to

substantial saving in capital cost, improvement of the safety

of the system and reducing environment impact [20]. In recent

years, many researches have been made in order to apply

this strategy in the coupling process. Bhat and Sadhukhan

[21] reviewed the PI for methane steam reforming by coupling

and using themembrane separation technologies to overcome

the mass transfer, heat transfer and thermodynamic limita-

tions. Altimari and Bildea [22] discussed about the design and

control of plant wide systems of coupling exothermic and

endothermic reactions in a steady state behavior. Patel and

Sunol [23] developed a distributed mathematical model for

thermally coupled membrane reactor to couple the methane

steam reforming through the exothermic methane combus-

tion. Glockler et al. [24] presented a theory to analyze and

optimize the reactor design to couple the exothermic and

endothermic reactions in an adiabatic fixed-bed reactor with

reverse flow mode. AboeGhander et al. [25] investigated

coecurrent and counterecurrent operation modes for the

above configuration.

Based on the researches made in literature, most of the

implemented studies on the coupled reactors have presented

higher conversion in the coecurrent mode compared with the

counterecurrent one. On the other hand, the results of some

specific studies showed the superiority of the counterecurrent

modeofoperation [25]. Thesedifferent resultsmadeuscompare

the performance of two reactor configurations, simultaneously.

Table 2 e Rate constants and heat of reaction forhydrogenation of nitrobenzene.

k ¼ A0expðB0 � E0=1:8 TÞ A0 B0 E0 DH298 K

k0 0.186 0 10 �44,3000

KH2 4.427 � 10�3 0 0 e

KNB 1.51 � 10�5 0 0 e

Table 3 e Specifications of reactors, feed, product andcatalyst of CTR.

Parameter Numericalvalue

Unit

Naphtha feed stock 30.41 � 103 Kg/hr

Reformate 24.66 � 103 Kg/hr

H2/HC mole ratio 4.74 e

LHSV 1.25 hr�1

Mole percent of hydrogen in recycle 69.5 e

Diameter and length of 1st reactor 1.25, 6.29 m

Diameter and length of 2nd reactor 1.67, 7.13 m

Diameter and length of 3rd reactor 1.98, 7.89 m

Distillation fraction of naphtha feed and reformate

TBP Naphtha

feed (�C)Reformate

(�C)IBP 106 44

10% 113 73

Fig. 1 e A simplified process for conventional tubular reactor (CTR).

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 5 3485

2. Reaction scheme and kinetics

2.1. Endothermic side (catalytic naphtha reforming)

Considering the simplified model suggested by smith [3], four

dominant idealized reactions can be used to describe the cata-

lytic reforming process so that the catalytic reforming system

can be simplified significantly. The reactions are as follows:

Naphthenes (CnH2n) 4 Aromatics (CnH2n�6) þ 3H2 (1)

Naphthenes (CnH2n) þ H2 4 Paraffins (CnH2nþ2) (2)

Naphthenes (CnH2n) þ n/3H2 / Lighter ends (C1eC5) (3)

Paraffins (CnH2nþ2) þ (n�3)/3H2 / Lighter ends (C1eC5) (4)

The corresponding rate expressions are as follow:

r1 ¼�kf1

Ke1

��ke1pn � pap

3h

�(5)

r2 ¼�kf2

K

��ke2pnph � pp

�(6)

30% 119 105

50% 125 123

70% 133 136

90% 144 153

FBP 173 181

Typical properties of catalyst

dp 1.2 mm

Pt 0.3 wt%

e2

r3 ¼�kf3

pt

�pn (7)

r4 ¼�kf4

pt

�pp (8)

Re 0.3 wt%

sa 220 m2/g

rB 0.3 Kg/l

e 0.36 e

The reaction rate constants ðkfi Þ , equilibrium constants

ðKei Þ, activation energies (Ei) and standard heat of reactions

(DH298 K) are tabulated in Table 1 [26]. The value of activation

energies and standard heat of reactions are written. The

reaction rate and equilibrium constants are temperature

dependent. The temperature dependent equations for calcu-

lation of these constants and their related coefficients have

been presented.

Fig. 2 e A schematic concept for thermally coupled reactors (TCRs) (a) Co-current and (b) Counter-current.

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 53486

2.2. Exothermic side (hydrogenation of nitrobenzene toaniline)

Over the last 145 years aniline has became one of the hundred

most important building blocks in chemistry. Aniline is

known to be used in more than 300 different end products

such as, isocyanates, rubber processing chemicals, dyes and

pigments, agricultural chemicals and pharmaceuticals [27].

Nitrobenzene is used as rowmaterial for aniline production by

almost entire world producers, using fixed-bed or fluidized-

bed vapor phase reactors [27,28].

The hydrogenation of nitrobenzene to aniline in the tube

side can be expressed by:

C6H5NO2 þ 3H2/C6H5NH2 þ 2H2O (9)

The rateexpressionprovidedbyKlemmetal. [29] is as follow:

r ¼ k0KNBKH2PNBP:5

H2�1þ KNBPNB þ KH2

P:5H2

�2 (10)

The reaction rate constant (k0), equilibrium constants (Ki),

the activation energy (E ) and standard heat of reaction

Table 4 e Operating condition for hydrogenation reactionof nitrobenzene to aniline.

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 5 3487

(DH298 K) are all given in Table 2. As in Table 1, the related

values and temperature equations and coefficients have been

presented in Table 2.

Parameter Numerical value Unit

Inlet temperature 780 K

Inlet pressure 1.1 � 105 Pa

Catalyst density 1400 Kg/m3

Diameter of catalyst particle 4.7 � 10�3 m

Void fraction 0.46 e

1st reactor

Available cross-sectional

area of the tube side for flow

3.29 m2

Nitrobenzene molar flow rate 7.24 mol/s

Hydrogen molar flow rate 28.98 mol/s

Aniline molar flow rate 0.72 mol/s

Steam molar flow rate 72.46 mol/s

2nd reactor

Available cross-sectional

area of the tube side for flow

2.91 m2

Nitrobenzene molar flow rate 6.41 mol/s

Hydrogen molar flow rate 25.67 mol/s

Aniline molar flow rate 0.64 mol/s

Steam molar flow rate 64.19 mol/s

3. Process description

3.1. Conventional tubular reactor (CTR)

In conventional fixed-bed reforming the naphtha (middle

distillates of atmospheric distillation column or condensate

from gas reservoirs) feedstock is combined with a recycle gas

stream contain 60%e90% (by mole) hydrogen. Hydrogen can

adjust the H2/HC ratio through the reactors to reduce coke

formation and also it sweeps the products through the cata-

lyst pores. The total reactor charge is heated and passed

through the catalytic reformers, which are designed with

three adiabatically operated reactors and three kilns between

the reactors to maintain the reaction temperatures at design

levels. The catalyst weight distribution in the reactors in the

conventional process is 0.2/0.3/0.5 [30]. The effluent from the

last reactor is cooled, and then enters the product separators.

The off gases and reformates are separated from the above

and bottom of the separator, respectively. Fig. 1 shows the

schematic of the process and Table 3 shows the catalyst

specification and operational conditions for the CTR.

3.2. Thermally coupled reactor configuration (TCR)

Fig. 2(a) and (b) provide schematic diagram of co and coun-

terecurrent modes of operation for coupled reactors,

respectively. The first and second reactors in the coupled

configuration composed of two compartments separated by

a solid wall. Catalytic reforming is supposed to takes place in

the shell side and the exothermic hydrogenation of nitro-

benzene to aniline is considered to provide the heat for

reforming. In the counterecurrent configuration the blend of

nitrobenzene and hydrogen feed the reactors from the

bottom. The external walls are assumed to be adiabatic.

The catalytic naphtha reforming reactions are carried out

over commercial Pt/Re/Al2O3 catalyst (0.3% Platinum, 0.3%

Rhenium). The outer compartment is filled with the spherical

commercial PdeAl2O3 (1.1 wt.%) catalyst [31]. Table 4 shows

the specific properties and operational conditions of the

exothermic side (hydrogenation of nitrobenzene). The molar

flow rates for nitrobenzene and steam are considered as those

reported by AboeGhander et al. [25].

Fig. 3 e Axial elements for the co-current and counter-

current configurations.

4. Reactor model

In this study a homogeneous oneedimensional model has

been considered. The basic structure of this model is

composed of heat and mass balance conservation equations

coupled through thermodynamic and kinetic relations, aswell

as auxiliary correlations for predicting physical properties.

Differential elements (Fig. 3) are considered along the axial

direction inside the reactor configurations (co and counter-

ecurrent). The mass and energy balances are written for fluid

phase in each side. The results are summarized in Table 5. The

positive sign in the left sides of equations (11) and (14) is used

for the coecurrent setup. Simultaneously, the negative one

demonstrates the counterecurrent regime. The negative sign

in the right side of equation (14) is used for the exothermic

reaction in the tube side.

5. Simulation and results

5.1. Steady state model validation for CTR

A comparison between the proposed model and conventional

tubular packed bed reactors at steady state condition has been

demonstrated in Table 6. The comparison of experimental

data and the results obtained from the model for 771 days of

operation has been presented in Table 7. As it can be seen,

there is an acceptable agreement between the plant data and

predicted mole fractions of components at the output of the

system. Analysis of the components (paraffin, naphthene and

aromatic) is performed by PONA Test in Stan Hop Seta appa-

ratus. The aromatic is tested especially by ASTM 2159 equiv-

alent to UOP 273 method [30].

Table 5 e Mass & energy balances and auxiliary correlations.

Definitions Equations

Fluid phase

� 1Acj

dFi;j

dz¼ ri;jrb (11)

dFtj ¼Xi

dFi;j (12)

dMave;j ¼dFtj

F2tj

_mj (13)

� 1Acj

Cgpj

d�FtjTj

�dz

¼ �pDj

Acj

UðTTube � TShellÞ þ rb

Xi

ri;j��DHf ;i

�(14)

dPdz

¼ 150m

f2sd

2p

ð1� eÞ2e3

QAc

þ 1:75rfsdp

ð1� eÞe3

Q2

Ac(15)

Boundary conditions

z ¼ 0; Fi;j ¼ Fi0;j;T ¼ T0

z ¼ L; Fi;j¼2 ¼ Fi0;j¼2;T2 ¼ T0;2(16)

Parameter Equation Reference

Component heat capacity Cp ¼ aþ bTþ cT2 þ dT�2 [32]

Viscositym ¼ C1TC2

1þ C3T þ C4

T2

[33]

Thermal conductivity Based on local compositions [34]

Re ¼ 2Rpugr

m

hi ¼ 4ðKwdpÞðrugdp

.mÞ0:3 [35]

Overall heat transfer coefficient 1U ¼ 1

hiþ Ai lnðDo=DiÞ

2pLKwþ Ai

Ao

1ho

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 53488

5.2. Endothermic side

5.2.1. ParaffinFig. 4(a) and (b) compares theconsumption rates of the reactants

for three configurations (CTR, co-current and counter-current

coupled reactors). Paraffin consumption is not considerable in

the first reactor as it is shown in Fig. 4(a). The dehydrogena-

tion reaction mainly occurs in this reactor. Providing a heating

medium for the reaction side (for endothermic section) can

Table 6 e Comparison between model prediction and plant da

Reactor no. Inlet temperature (K) Inlet pressure (Kpa)

1 777 3703

2 777 3537

3 775 3401

Outlet temperature (K)

Reactor no. Plant CTR TCR

Co-current Counter-cur

1 722 727 760.8 767.2

2 753 751 801.2 802

3 770 771 799.6 798.9

changethereaction ratesandequilibriumconditions.Thiscould

be achieved by utilizing the coupled reactor configuration. The

generated heat in the exothermic side is straightly consumed

by the endothermic section in the co-current regime and

this increases the paraffin consumption rate in the co-current

reactor in comparison with the counter-current one. A sharp

slope in this figure for CTR and TCRs can be seen for the second

and third reactors. This means that the forth reaction will be

predominant in these two reactors.

ta for fresh catalyst.

Catalyst distribution (wt%) Input feedstock (Mole%)

20 Paraffin 49.3

30 Naphthene 36

50 Aromatic 14.7

Aromatic in reformate (Mole %)

Plant CTR TCR

rent Co-current Counter-current

e e 44.02 39.96

e e 62.10 53.35

57.7 56.2 85.77 81.75

Table 7 e Comparison between predicted production rateand plant data.

Time(day)

Naphtha feed(ton/h)

Plant(kmol/hr)

Model(kmol/hr)

Devi%

0 30.41 225.90 221.5802 �1.91

34 30.41 224.25 222.5122 �0.77

62 31.00 229.65 227.9313 �0.75

97 30.78 229.65 226.7749 �1.25

125 31.22 229.65 230.7985 þ0.50

160 31.22 229.65 231.2730 þ0.71

188 28.55 211.60 209.8377 �0.83

223 30.33 222.75 224.7291 þ0.88

243 31.22 233.05 232.1821 �0.37

298 30.67 228.65 228.1752 �0.21

321 30.76 227.64 229.0932 þ0.64

398 42.35 317.30 324.5555 þ2.28

425 42.32 317.94 324.4826 þ2.05

461 42.32 317.94 324.6987 þ2.12

490 42.32 317.94 324.8622 þ2.17

524 42.32 313.09 325.0433 þ3.81

567 42.54 317.94 327.0586 þ2.86

610 42.54 313.90 327.2581 þ4.25

717 37.86 286.15 289.3742 þ1.12

771 38.51 282.10 294.9026 þ4.53

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

0

10

20

30

40

50

60

70

80

90

100

Mass of catalyst (Dimensionless)

Nap

hthe

n m

olar

flo

w r

ate

(km

ole/

hr)

0.8 0.9 1

1

2

3

4

CTR

Co Current

Counter Current

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

20

40

60

80

100

120

140a

b

Mass of catalyst (Dimensionless)

Par

affi

n m

olar

flo

w r

ate

(km

ole/

hr)

Counter Current

Co Current

CTR

Fig. 4 e Molar flow rates for (a) paraffine and (b) naphthene

vs. the dimensionless mass of catalyst in the endothermic

side for CTR and TCRs.

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 5 3489

The naphthenes consumption profile is illustrated in Fig. 4

(b). The achieved results from the coupled reactor configura-

tion show that the configuration with just two coupled reac-

tors has the same results as the CTR (with three reactor) and

even more acceptable.

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

20

40

60

80

100

120

140

160a

b

c

Mass of catalyst (Dimensionless)

Aro

mat

ic m

olar

flo

w r

ate

(km

ole/

hr)

CTR

Co Current

Counter Current

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

1250

1300

1350

1400

1450

1500

1550

Mass of catalyst (Dimensionless)

Hyd

roge

n m

olar

flo

w r

ate

(km

ole/

hr)

CTR

Co Current

Counter Current

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

550

600

650

700

750

800

Mass of catalyst (Dimensionless)

Lig

ht e

nds

mol

ar f

low

rat

e (k

mol

e/hr

)

CTR

Co Current

Counter Current

Fig. 5 e Products’ molar flow rates of (a) aromatic, (b)

hydrogen and (c) the light ends along the CTR and TCRs.

Table 8 e Third reactor outlet and Reformate compositions.

Paraffin Naphthene Aromatic Hydrogen Light end

Third reactor outlet composition (% w/w)

CTR 19.7849 0.7744 26.8334 5.1442 47.4630

Co-current 4.8546 0.2776 31.4891 5.1226 58.2560

Counter-current 6.4985 0.3300 31.1464 5.1412 56.8839

Reformate composition (% w/w)

CTR 41.7468 1.6339 56.6193

Co-current 13.2563 0.7582 85.9856

Counter-current 17.1126 0.8690 82.0183

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 53490

5.2.2. Aromatics, hydrogen and lighter endsFig. 5(a) and (c) reveals the products molar flow rates in the

endothermic section of the coupled reactor configuration. An

interesting result is achieved by the modeling of TCRs (co-

current and counter-current regimes). Higher aromatic and

lower hydrogen production rates are achieved in co-current

flow compared with the counter-current. This can be seen in

Fig. 5(a) and (b).Generally, theuseofproperflowregimeofTCRs

depends on the domestic use of plant products and down-

streamunits. Themain goal of the catalytic naphtha reforming

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

23

23.5

24

24.5

25

25.5

26

26.5

27

27.5

28

Mass of catalyst (Dimensionless)

Mol

ecul

ar w

eigh

t (k

g/km

ole) CTR

Co Current

Counter Current

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

2050

2100

2150

2200

2250

2300

2350

2400

2450

2500a

b

Mass of catalyst (Dimensionless)

Tot

al m

olar

flo

w r

ate

(km

ole/

hr)

CTR

Co Current

Counter Current

Fig. 6 e Variation of (a) total molar flow rate and (b) the

molecular weight in the CTR and TCRs.

is to produce high-octane gasoline. Therefore, the co-current

regime is more favorable. Also the co-current reactors operate

more easily than the counter-current ones. Fig. 5(a) and (b) also

shows higher aromatic and hydrogen production in the outlet

of the second reactor of TCRs configuration in comparisonwith

the outlet of the third reactor in the CTR. As a result, the third

reactor, with the highest catalyst load (as much as the sum of

the first and second reactors) can be neglected. This achieve-

ment leads to ahugecapital andoperational investment saving

(the catalyst contains Pt metal). Fig. 5 (c) reveals the light end

production. These gases are used as a feed stock for LPG

production. Table 8 shows the composition of reformate and

product stream of the third reactor for each case.

5.2.3. Change of molar flow rate and molecular weightAs it is demonstrated in Fig. 6(a) the variation of the totalmolar

flow rate from the entrance of the first reactor to the outlet of

the last reactor is about 350 kmol/h (about 17% changes in

molar flow rate). This is considered during the modeling and

amore comprehensivemodel comparedwith previous studies

is present in this research. The TCRs produce lighter hydro-

carbons such as lighter ends and aromatics comparedwith the

CTR. Accordingly, the average molecular weight is lower for

the TCRs (see Fig. 6(b)). Also the Heat capacity, viscosity,

density and other physical properties are considered to be

variable in both shell and tube sides in this study.

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

3200

3300

3400

3500

3600

3700

3800

Mass of catalyst (Dimensionless)

Pre

ssur

e (k

Pa)

CTR

Co Current

Counter Current

Fig. 7 e Pressure drop along the endothermic side.

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

4.5

5

5.5

6

6.5

7

7.5

8

8.5

Mass of catalyst (Dimensionless)

H2/H

C m

olar

rat

io

CTR

Co Current

Counter Current

Fig. 8 e H2/HC ratio through the reactor for CTR and TCRs.

0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5

0

5

10

15

20

24.84

26.06

30a

b

Mass of catalyst (Dimensionless)

Ani

lin m

olar

flo

w r

ate

(km

ole/

hr)

Co Current

Counter Current

0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5

265

270

275

280

285

290

295

300

305

312.5

310

314.8

320

Mass of catalyst (Dimensionless)

H2O

mol

ar f

low

rat

e (k

mol

e/hr

)

Co Current

Counter Current

Fig. 9 e Changes of (a) aniline and (b) water in the

exothermic side for TCRs (co-current and counter-current).

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

720

730

740

750

760

770

780

790

800

810a

b

c

Mass of catalyst (Dimensionless)

Tem

pera

ture

of

shel

l sid

e (K

)

Co Current

Counter Current

CTR

0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5

0

0.5

1

1.5

2

2.5

x 10

4

Mass of catalyst (Dimensionless)

Hea

t (k

J/kg

cat.hr

)

0.19 0.195 0.2

1000

1200

1400

Endothermic Heat of Reaction

Transfered Heat

0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5

0

0.5

1

1.5

2

2.5

x 10

4

Mass of catalyst (Dimensionless)

Hea

t (k

J/kg

cat.hr

)

0.02 0.04

0

500

1000

1500Enothermic Heat of Reaction

Transfered Heat

Fig. 10 e (a) Thermal profiles of CTR and TCRs, and the

consumed heat of endothermic side and transfer heat

through the solid walls of (b) co-current and (c) counter-

current.

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 5 3491

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 53492

5.2.4. Pressure dropSince the total molar flow rate increases, the velocity and

related viscose losses as well as the pressure drop increase.

Fig. 7 depicts the pressure profile along the CTR and TCRs. The

related equation (Ergun equation) was shown in Table 5.

5.2.5. H2/HC ratioClearly theH2/HC ratio increases along the reactor according to

thehydrogenproduction. Increasing this ratio is a guarantee to

enhance the catalyst life time and also reduces load on the

catalyst. On the other hand, it is a constraint and can shift the

dehydrogenation reaction in the reverse direction to reduce

thehydrogenandaromaticproductions.Higher temperature in

the TCRs can aid to reduce the undesirable effect of increasing

in the H2/HC. Fig. 8 shows this ratio in the CTR and TCRs.

5.3. Exothermic side

5.3.1. Molar behaviorFig. 9(a) and (b) presents themolar flow rate of the exothermic

side of the reactor for both coecurrent and counterecurrent

mode of operation. The same value for the space velocity

0 0.025 0.05 0.075 0.1 0.125 0.15 0.175 0.2

750

800

850

900

950

1000

1050

1100a

b

Mass of catalyst (Dimensionless)

Tem

pera

ture

of

exot

herm

ic s

ide

(K)

N=1

N=10

N=100

N=750

Counter CurrentCo Current

0.2 0.25 0.3 0.35 0.4 0.45 0.5

650

700

750

800

850

900

950

1000

Mass of catalyst (Dimensionless)

Tem

pera

ture

of

exot

herm

ic s

ide

(K)

N=1

N=10

N=100

N=750

Co CurrentCounter Current

Fig. 11 e Effect of tube’s number on the developed hot spot

in the exothermic sides of co and counter-current reactors

for (a) first reactor and (b) second reactor.

which was considered by AboeGhander et al. [25] has been

applied in our study. The molar flow rates of the exothermic

components are basedon stoichiometric values. Adding steam

to the feed lineprevents the coke formationon the catalyst and

reduces the temperature along the hydrogenation side and

maintains it low by diluting the reacting mixture on the

hydrogenation side [25,36]. The hydrogenation of nitroben-

zene is an irreversible reaction and it is not limited due to the

thermodynamic equilibrium and it continues to the end. Fuel

depletion will impress the reaction rate and will stop it. The

concentrated feed generates high heat and consequently

a considerable peak in the temperature. Sowe divided the feed

stream into two parallel lines to avoid this phenomenon. The

conversions for nitrobenzene in the first and the second reac-

tors are 86.40e86.38% and 90.60e90.59% respectively for co-

current and counter-current regimes.

5.4. Thermall behavior

5.4.1. Endothermic sideFig. 10(a) presents the temperature profile in the endothermic

side of the CTR and TCRs simultaneously. The temperature in

0 0.02 0.04 0.06 0.08 0.1 0.12 0.14 0.16 0.18 0.2

0

200

400

600

800

1000

1200

1400

1600

a

b

Mass of catalyst (Dimensionless)

Hea

t (k

j/kg

cat.hr

)

Exothermic Heat of Reaction

Transfered Heat

Counter Current

Co Current

N=750

0.2 0.25 0.3 0.35 0.4 0.45 0.5

-200

0

200

400

600

800

1000

1200

1400

1600

1792

Mass of catalyst (Dimensionless)

Hea

t (k

j/kg

cat.hr

)

Exothermic Heat of Reaction

Transfered Heat

Counter Current

Co Current

N=750

Fig. 12 e Exothermic heat of reaction and the transferred

heat in the TCRs for (a) first reactor and (b) second reactor.

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 5 3493

the CTR falls in each reactor due to the predominant endo-

thermic reaction. The temperature reduction of the first reactor

in TCR configuration for co-current flow pattern is lower

compared with the CTR because of the heat source (exothermic

reaction inthe tubeside).Considering thecaseof co-currentflow

in the thermally coupled reactor configuration (Fig. 10b), the

consumption heat in the endothermic side is higher compared

with the transferred heat through the tube wall in the first

reactor.Thiscan justify the temperature reductionalong thefirst

reactor. In the second reactor, the amount of transferred heat is

higher comparedwith theheat consumption in theendothermic

side which raises the temperature of the system. The heat

transfer for the counter-current flow regime has been shown in

Fig. 10(c).As it canbeseen, at theentranceof thereactor, theheat

consumption rate is higher compared with the transferred heat

through the tubewall. Therefore, the temperature of the system

reduces. In a certain axial position in Fig. 10(c), (dimensionless

mass of catalyst ¼ 0.1) the amount of transferred heat exceeds

the amount of consumed heat in the endothermic side which

justifies the temperature increase in Fig. 10(a). The same trend

can be seen in the second reactor and a minimum in tempera-

ture profile of the endothermic side occurs for the counter-

current flow regime of the TCR in this reactor.

5.4.2. Exothermic sideFig. 11(a) and (b) shows the exothermic temperature profiles in

the first and second reactors for both co and counter-current

flows.As thenumber of tubes increases, thehot spot decreases

and the thermal behavior improves.

Fig. 12 demonstrates the heat generation in the exothermic

sides for co-current and counter-current flows. The generated

heat in the exothermic side is higher for the co-current config-

uration in the first reactor (also the transferred heat through

the solid wall). Accordingly, the temperature of the exothermic

side remains higher (higher maximum temperature for co-

current). Fig. 12(b) demonstrates a reverse phenomenon for the

exothermic side of the second reactor (higher maximum

temperature for counter-current).

6. Conclusion

Inorder tosaveenergyandminimize thecapitalandoperational

costs, a novel reactor configuration has been proposed for the

catalytic naphtha reforming process in previous studies. In this

research, two different modes of operation have

been investigated in order to determine their advantages and

disadvantages. Considering aromatics and hydrogen as the

main products of the naphtha reforming process, the variation

of the production rate of these two componentsmay clarify the

situation. According to the calculations made by this research

group, higher aromatics production rate can be achieved by

utilizing the co-current mode of operation. On the other hand,

by utilizing the counter-current flow regime, the hydrogen

production rate increasedbyabout 5 kmol/h comparedwith the

co-current flow. The pressure and temperature profile also has

been presented in this study. Generally, according to all the

calculations made in this research, utilizing the co-current

mode of operation is given priority. These results suggest that

the counter-current configuration can be a compelling way

to boost hydrogen production and the co-current mode of

operation enhances the aromatic production rate and ensure

the other desired results, as well. However, an investigation in

relation to the environmental aspects, commercial viability and

economic feasibility of the proposed configuration is necessary

in order to consider commercialization of the process.

Nomenclature

a catalyst activity

A moles of aromatic formed, kmol h�1

Ac cross-section area of reactor, m2

cp specific heat capacity, kJ kmol�1 K�1

dp particle diameter, m

Ed activation energy of catalyst, J mol�1

Ei activation energy for ith reaction, kJ kmol�1

Fi molar flow rate of component i

hf heat transfer coefficient, W m�2 K�1

H2 hydrogen, kmol h�1

keff effective thermal conductivity, W m�1 s�1

kf1 forward rate constant for reaction (1),

kmol h�1 kg cat�1 MPa�1

kf2 forward rate constant for reaction (2),

kmol h�1 kg cat�1 MPa�2

kf3 forward rate constant for reactions (3),

kmol h�1 kg cat�1 MPa�2

kf4 forward rate constant for reactions (4),

kmol h�1 kg cat�1 MPa�2

Ke1 equilibrium constant, MPa3

Ke2 equilibrium constant, M Pa�1

L length of reactor, m

Mi molecular weight of component i, kg kmol�1

Mw average molecular weight of the feedstock,

kg kmol�1

n average carbon number for naphtha

NA molar flow rate of aromatic, kmol h�1

Ni molar flow rate of component i, kmol h�1

p moles of paraffin formed, kmol h�1

Pi partial pressure of ith component, kPa

P total pressure, kPa

Q volumetric flow rate, m3 s�1

ri rate of reaction for ith reaction, kmol kg cat�1 h�1

R gas constant, kJ kmol�1 K�1

sa specific surface area of catalyst pellet, m2 kg�1

T temperature of gas phase, K

yi mole fraction for ith component in gas phase

e void fraction of catalyst bed

m viscosity of gas phase, kg m�1 s�1

r density of gas phase, kg m�3

rb reactor bulk density, kg m�3

fs sphericity

DH heat of reaction, kJ kmol�1

a aromatic

cal calculated

h hydrogen

i number of components

j number of reaction sides (1: Endothermic, 2:

Exothermic)

n naphthene

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 3 4 8 3e3 4 9 53494

out outlet

p paraffin

CTR conventional tubular reactor

FBP final boiling pint, �CHC Hydrocarbon, kmol h�1

IBP initial boiling pint, �CLHSV liquid hourly space velocity, hr�1

Pt platinum

Re rhenium

RON research octane number

TBP true boiling point, �KTCR thermally coupled reactor

WHSV weight hourly space velocity, hr�1

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