Reconfiguration of the E&C Refinery Co. Crude Distillation ...

Upgrade Project: Reconfiguration of the E&C Refinery Co. Crude Distillation Column in Irvine

Cara Chan

Justin Chung

Caitlin Lim

Jacqueline Medina

David Nguyen

CBEMS 149B: Group 6

Professor Cesar Ruiz, P.E. University of California, Irvine

Irvine, California

June 9th, 2017

Letter of Transmittal

Date: June 9th, 2017

To: Cesar Ruiz, P.E.

Fluor Process Engineer and Lecturer Department of Chemical Engineering and Materials Science

University of California, Irvine

Subject: Completion of the Upgrade Project: Reconfiguration of the E&C Refinery Co. Column

We are pleased to submit to you the report on “Upgrade Project: Reconfiguration of the E&C Refinery Co. Crude Distillation Column in Irvine.”

This report utilized the knowledge and experience that we have accumulated throughout our studies from the Chemical Engineering Curriculum; specifically, the Chemical Engineering Design series that provided much insight on how process engineers conduct their work. This senior design project was to imitate real-life scenarios, and our team treated it as such.

Within this paper includes our findings and upgrade recommendations of the equipments involved in the crude distillation process. These recommendations are to satisfy the process rate of 265,000 BPSD per request from the E&C Refinery Co, Irvine. If any clarification is needed, we will happily be able to answer your questions.

Thank you for passing your knowledge to us,

David Nguyen

Cara Chan

Justin Chung

Jackie Medina

Caitlin Lim

Executive Summary

This document is a proposal to the E&C Refinery Company for a chemical plant with the

capacity of processing 265,000 BPSD of crude. The proposal details the following: equipment

sizing for distillation columns, heat exchangers, pumps, and vessels are outlined including

calculations; the cost estimates of the plant include the capital cost of purchasing equipment

and ISBL fixed capital costs; risk assessments are considered based on FMEA guidelines; the

environmental impact is discussed concerning air pollutants, process hazards, and how to

mitigate them. The main saleable products from the plant are naphtha, kerosene, turbosene

and diesel. The proposed heat integration is designed to conserve energy by first utilizing the

heat in the product streams to preheat the raw crude before using utility to obtain the desired

operating temperature of 708 °F. The estimated fixed capital investment is $62 million which will

be recovered in 10 days of the plant operating at full capacity. The ISBL fixed capital costs are

estimated at $326 million which can be recovered within 2 months when at full capacity.

Table of Contents Introduction

1: Process Description and Feed Specification

1.1 Process Description

1.2 Feed Specification

2: Product Rate and Properties

3: Heat Integration

4: Heat Exchangers

4.1 The Required Heat Exchangers for the Preheat Process

4.2 Product Temperature Specification and Utility Usage

5: Pump Calculations

6: Column Sizing

6.1 Wall Thickness

7: Vessel Sizing

7.1 Raw Crude Charge Drum Dimensions

7.2 Crude Flash Drum Dimensions

7.3 Vessel Wall Thickness

8: Estimated Cost of Equipment

8.1 Estimated Plant Cost

8.2 Estimated Return on Investments

9: Risk Assessment and Adverse Health Effects

10: Environmental Impacts

11: Failure Mode Effect Analyses (FMEA) on Crude Flash Drum

Conclusion

References

Appendices

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Introduction

E&C Refinery Company plans to redesign its refinery in Irvine, CA, to process 265,000

BPSD of petroleum. Currently, the Crude Distillation Unit (Unit 1) operates at 100,000 BPSD

and cannot account for the remaining 165,000 BPSD. Therefore, a new unit will be required to

fully process the 265,000 BPSD of raw crude oil.

The main crude unit in the refinery uses distillation to separate the crude oil’s

components into useful, marketable products. This project will focus on the atmospheric

distillation column and support equipment (piping, preheaters, desalter, vessels, and pumps).

This project aimed to design the different equipment components to optimize the process of

refining the petroleum and determining cost-effective designs; however, safety and

environmental precautions also need to be considered to meet standard guidelines in protecting

civilians and the environment.

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1: Process Description and Feed Specification

1.1 Process Description

The raw crude oil will be heated, partially vaporized and then fed into the atmospheric

tower. The vapors from the tower will be fractioned off at defined boiling points to produce a

variety of products. Lighter products such as Naphtha are recovered at low temperatures at the

top of the distillation unit. Mid-range products such as Turbosene and Kerosene are sent from

the atmospheric tower to side strippers where they are recovered. Heavier products, such as

residual fuel oil, are recovered at high temperatures from the bottom of the unit.

The hot final products leaving the distillation columns will be contacted with the cold feed

through a heat exchanger system. This is to conserve energy by taking advantage of heat

already present in the system. This also limits the use of utilities such as water and air in the

process of heating the crude and cooling the products. The amount of shells required and the

streams of interest for this heat exchange system are available in the heat integration.

The feed and products are moved from one equipment to another by using single stage

centrifugal pumps. The overall process can be seen in Appendix A.

1.2 Feed Specification

The feed will be a mixture of 2 different crudes: crude A and crude B. The operating

conditions and flowrate are presented in Table I.

Table I: Feed Operating Conditions

Stream Operating Condition Value

Crude A Temperature 68 F

Pressure 5 PSIG

Rate 107,250 BPSD

Average Specific Gravity 0.8628

Crude B Temperature 65 F

Pressure 5 PSIG

Rate 57,750 BPSD

Average Specific Gravity 0.9277

3

Other known specifications for the feed such as true boiling point (TBP) distillation data, lightend

composition, Nickel content, Vanadium content, basic Nitrogen content, Sulfur content,

Neutralize number (ANEU), Asphaltene content, and kinematic viscosity are available in

Appendix A.

2: Product Rate and Properties

The main process for the crude distillation system is to take raw crude A and crude B oil

and separate into distant products or “cuts” based on the boiling point properties. Off gas,

Naphtha, jet/turbosene, kerosene, diesel, gas oil, and atmospheric resistance were the distant

products that was created from the raw crude oil. Naphtha, jet/turbosene, kerosene, and diesel

were tested under the ASTM D86 method and must meet the specification as shown in Table II.

Table II: Required product specification

STREAM ID CRUDEA CRUDEB

2C002 OR

2C002A 2C001C 2C005A1 2C007C

DESCRIPTION UNITS RAWCRUDEA RAWCRUDEB

MIXED CRUDE WITH

WATER

RAW CRUDE AFTER PRH-1

RAW CRUDE AFTER PH-2

RAW CRUDE AFTER PH-3

TOTAL STREAM FLOW, MOLAR lb-mol/h 6280 2720 1580 9000 10100 7920

FLOW, MASS M lb/hr 1.35 0.782 2.25 2.13 2.15 2.03

TEMPERATURE °F 68.0 68.0 295 295 325 440

PRESSURE psig 5.00 -1.29 238 238 158 232 VAPOR PHASE FLOW, MOLAR lb-mol/h n/a n/a n/a n/a n/a n/a

FLOW, MASS M lb/hr n/a n/a n/a n/a n/a n/a STD VAPOR RATE M ft3/h 1.59 0.02 6.59 0.0470 2.85 0.0444

SPECIFIC Btu/lb-°F n/a n/a n/a n/a n/a n/a

4

HEAT (CP)

VISCOSITY cP n/a n/a n/a n/a n/a n/a LIQUID PHASE FLOW, MOLAR lb-mol/h 6280 2720 15800 9000 10100 7920

FLOW, MASS M lb/hr 1.35 0.782 2.25 2.13 2.15 2.03 FLOW, VOLUME gal/min 3140 1690 5610 5350 5490 5380 SPECIFIC HEAT (CP) Btu/lb-°F 0.436 0.406 0.582 0.557 0.576 0.618

DENSITY lb/ft3 53.6 57.7 50.1 49.7 48.9 47.0

VISCOSITY cP 6.27 22.8 0.648 1.28 0.922 0.715 SURFACE TENSION dyne/cm 29.0 32.2 32.5 19.4 21.1 14.9

STREAM ID 2C100 2C400 2C510B 2C600B 2C710B 2C800B 2C900

DESCRIPTION UNITS OFFGAS NAPHTHA PRODUCT

Turbosene PRODUCT

KEROSENE PRODUCT

DIESEL PRODUCT

GAS OIL PRODUCT

ATMOSPHERIC RESID

TOTAL STREAM

FLOW, MOLAR lb-mol/h 312 3510 1370 1030 708 319 1830

FLOW, MASS M lb/hr 0.0192 0.362 0.217 0.221 0.190 0.106 1.02

TEMPERATURE °F 128 128 396 535 519 610 695

PRESSURE psig 1.85 1.85 8.75 10.3 11.4 12.4 13.9

VAPOR PHASE

FLOW, MOLAR lb-mol/h 312 n/a n/a n/a n/a n/a n/a

FLOW, MASS M lb/hr 0.0192 n/a n/a n/a n/a n/a n/a STD VAPOR RATE M ft3/h 0.147 1.57 0.544 0.285 0.00420 0.00232 0.0207 SPECIFIC HEAT (CP) Btu/lb-°F 0.429 n/a n/a n/a n/a n/a n/a

VISCOSITY cP 0.00809 n/a n/a n/a n/a n/a n/a

LIQUID PHASE

FLOW, MOLAR lb-mol/h n/a 3510 1370 1030 708 319 1830

FLOW, MASS M lb/hr n/a 0.362 0.217 0.221 0.190 0.106 1.02 FLOW, VOLUME gal/min n/a 1040 665 680 533 296 2490 SPECIFIC HEAT (CP) Btu/lb-°F n/a 0.538 0.634 0.673 0.653 0.681 0.684

DENSITY lb/ft3 n/a 43.3 40.8 40.5 44.4 44.8 50.9

VISCOSITY cP n/a 0.355 0.218 0.216 0.352 0.341 1.27 SURFACE TENSION dyne/cm n/a 19.0 12.1 10.3 14.0 13.3 17.1

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Table III: List of Product Properties from PRO/II Simulation

By using PRO/II simulation, listed in Table III, the product rate and properties were determined

and compared to the required specifications.

3: Heat Integration

The raw crude needs to be preheated to 708°F before reaching the main distillation

column. To accomplish this, the crude feed would go through four separate networks of heat

exchangers. The exchangers were labeled as Preheat 1, Preheat 2, Preheat 3, and HXR-

1. Preheat 1 and 2 was used to heat the crude to a specific temperature, 289°F and 310°F

respectively. The final temperature of Preheat 3 was adjustable because HXR-1 provided the

remainder duty (Q) necessary to reach the final temperature of 708°F. This feed would be ready

for retrieval by the distillation column. Figure A depicts the heat integration graph.

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Figure A: Heat integration chart. Preheat 1 required 7 steps, Preheat 2 required 1 step, and Preheat 3 required 3 steps

Each Preheat utilized product streams from the distillation column (such as naphtha, diesel,

AGO, etc.) to preheat the crude feed. This is done to fully optimize the heating process to

reduce the need of using air or water utilities; essentially, this optimization would save E&C

Refinery Co. in Irvine the cost of production from utility usage.

Table IV shows the duty and temperature input/output for each stream utilized in the

preheating process. Notice in Table IV that some streams had gone through multiple passes of

specific preheats; these streams are denoted as Upper and Lower to signify upper limit

temperature and lower limit temperature. For example, HX-1G heat exchanger housed the initial

kerosene product which is set at the upper limit of 534.8°F. A duty of 32.3 MMBtu/h was

exchanged with the crude feed to provide a kerosene temperature output of 300°F. The

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adjusted temperature of the kerosene stream goes to HX-1B for further heat exchange with the

crude feed to cool the product to 100°F.

Table IV: Duty and temperature for each stream according to the heat integration chart. Hot streams are the product in each description and the cold stream is the crude feed

4: Heat Exchangers

4.1 The Required Heat Exchangers for the Preheat Process

Utilizing the heat integration graph, heat exchangers were designed as

appropriate. Figure B shows the flow diagram of the heat exchangers and Table V provides

the specifications of each heat exchanger. The overall heat transfer coefficient (U) and

correction factor (fT) was provided by E&C Refining Co. Log mean temperature difference

(LMTD) and shell surface areas were calculated with the following equations:

Eq. 1 LMTD = ∆𝑇𝐴−∆𝑇𝐵

ln (∆𝑇𝐴∆𝑇𝐵

)

Where: ∆𝑇𝐴 = 𝑇ℎ𝑜𝑡 𝑜𝑢𝑡 − 𝑇𝑐𝑜𝑙𝑑 𝑖𝑛

∆𝑇𝐵 = 𝑇ℎ𝑜𝑡 𝑖𝑛 − 𝑇𝑐𝑜𝑙𝑑 𝑜𝑢𝑡

Eq. 2 Surface Area (A) = 𝑄∗106

𝑈∗𝐿𝑀𝑇𝐷∗𝑓𝑇

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Table V: Each heat exchanger with the calculated surface area

Figure B: Flow diagram of the heat exchangers. Preheat 1 is boxed in blue, Preheat 2 boxed in red, Preheat 3 in grey, and Utilities Only in black. HXR-1 is omitted.

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The surface area for each heat exchanger had a limit of 5000 ft2; therefore, any stream that

needed a surface area greater than the limit would be divided evenly between multiple

exchangers. Kerosene lower (HX-1B), for instance, required a total area of 5590 ft2. Due to the

constraint, two heat exchangers in parallel would require 2795 ft2 each. Sample calculations of

this can be seen in the appendix. Also note that HX-3A required 17 separate heat exchangers

running in three series of six in parallel. Figure C and Figure D shows an example layout of

HX-1B and HX-3A.

Figure C: Lower kerosene product stream utilized in two heat exchangers in parallel

Figure D: Preheat 3 consist of 17 heat exchangers that utilized ATM residual product stream. This layout is favorable compared to utility use

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4.2 Product Temperature Specification and Utility Usage

Product streams such as AGO, turbosene, diesel, naphtha and diesel pumparound used

utility to meet the required temperature specification for each product. Provided in Table IV are

the mandatory temperatures for each product and utility usage:

Table VI: Type of utility used in each product to reach specific temperature goals

All applicable products used utilities after the preheating exchange to reach the temperature

goals except for AGO. In Figure B, AGO can be seen going through Preheat 2 first before

transferring to Preheat 1. In between the preheats will be a utility to provide a duty of 2.93

MMBtu/h with a temperature difference of 50°F. Diesel Pumparound and Naphtha only used

utilities to reach the necessary final temperature and was not used in the preheating process.

Once the preheating process was completed, the crude was at 480°F. The target temperature

was 708°F; therefore, a heater with a tag number of HXR-1 was utilized. Table VII provides the

heater’s specifications.

Table VII: Duty and temperature range of the heater

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5: Pump Calculations

Pumps assist in the movement of fluid from one location to the next. Without pumps, the

fluid will not reach its destination due to the frictional pressure loss from the pipes and

equipment [1]. To account for the pressure loss, the pumps are designed to provide additional

pressure to overcome the pressure loss. Determining the best specifications for each pump is

crucial in terms of compatibility and efficiency. The pump needs to be compatible to the type of

fluid and sized correctly for the process. If the pumps are not sized properly for the process, the

amount of pressure provided will be insufficient in terms of driving the fluid to its destination,

causing the pump to be inefficient. Certain fluids can corrode the pump’s material; also, the

viscosity of the fluid can affect the amount of power at the pump output.

In order to correctly choose the most suitable pump for the crude and product streams,

the Gobi pump calculation method was performed. Pump calculations, provided in Appendix B,

gathers information about the fluid properties and pressure drops for the streams and

equipment. The PRO/II program is used to determine the stream properties of the fluid as

shown in Table VIII.

Table VII: Stream Properties from PRO/II

Stream Name

2C001A2 2C007 2COLREFLUX 2C400

Stream Description

Raw Crude Flashed Crude

Reflux Naphtha Product

Temperature F 67.6 303.0 128 128

Pressure PSIG -1.29 30.0 1.85 1.85

Boiling Point Pressure at Stream T

psia 2.79 41.5 13.4 13.4

Flowrate kg-

mol/min 68.0 59.9 36.3 26.5

Liquid Act. Rate (vol)

gal/min 4830 5020 1430 1040

Liquid Std. Sp. Gr.

0.886 0.899 0.725 0.725

Liquid Act. Density

kg/m3 882 808 694 694

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Other important information for the pump is provided from the company's specification as shown

in Table IX.

Table IX: Pump Specification

Service Rated Flow

Minimum Flow

Suction Pressure

psig

Destination Pressure

psig

Destination Height

ft

Destination Line Loss

psi

Pump Efficiency

%

Motor Efficiency

%

Raw Crude Charge Pump

110% of

Normal 50% of Normal 10 250 45 20 75 90

Flashed Crude Charge Pump

110% of

Normal 50% of Normal

From Simulation 16 50 20 75 90

Reflux Pump

110% of


From Simulation

From Simulation 145 20 75 90

Naphtha Product Pump

110% of


From Simulation

Atmospheric Tank 25 20 75 90

Liquid Viscosity

cP 9.46 1.46 0.355 0.355

Stream Name

2C510B 2C600B 2C710B 2C800B 2C900

Stream Description

Jet Product

Kerosene Product

Diesel Product

Gas Oil Product

ATM Resid

Temperature F 396 535 519 610 695

Pressure PSIG 8.75 10.3 11.4 12.4 13.9

Boiling Point Pressure at Stream T

psia 20.3 21.8 22.9 23.9 25.4

Flowrate kg-

mol/min 10.4 7.82 5.36 2.41 13.9

Liquid Act. Rate (vol)

gal/min 665 680 533 296 2490

Liquid Std. Sp. Gr.

0.797 0.841 0.877 0.906 1.01

Liquid Act. Density

kg/m3 653 649 711 717 815

Liquid Viscosity

cP 0.218 0.216 0.352 0.341 1.27

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Jet/Turbo Product Pump

110% of


From Simulation


Kerosene Product Pump

110% of


From Simulation


Diesel Product Pump

110% of


From Simulation


Gas Oil Product Pump

110% of


From Simulation


Atmospheric Resid Pump

110% of


From Simulation


As seen in Table IX, the rated flow and the minimum flow is 110% and 50% respectively of the

normal flow for each of the pumps. The suction pressures, also known as the source pressures,

are the operating pressures of streams which can be found in Table VIII. The maximum source

pressure is assumed to be 50 psi for each pump calculation.

Besides the maximum source pressure, there are a few other assumptions made for the

meter loss, skirt height, and pump elevation. The meter loss was assumed to be 2 psi and the

line loss 1 psi. The skirt height was assumed to be 20 ft. Lastly, the pump elevation was

assumed to be 3ft for each pump calculation.

For the equipment loss, most values were obtained from the heaters. Based on the heat

integration, the amount of shells present affected the equipment loss. For each shell, the

pressure loss was 10 psi. Some streams are part of the utilities and are assumed to have

equipment loss from coolers. The pressure loss by the coolers was 5 psi.

There are a few criteria that must be addressed when handling pump calculations. There

are four criteria for the control valves: 15 psi minimum at normal flow, 10 psi minimum at rated

flow, 50% variable system drop at normal flow, and 25% variable system drop at rated flow. For

some pump calculation, the valves fulfilled all four criteria. The chosen control valve value was

the one that equated to a larger total discharge pressure. For majority of the pump calculation,

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the rated flow condition was used due to the criteria listed and contributed to a larger total

discharge pressure. Another criterion that must be followed is the safety margin for the NPSH

calculation. The safety margin is 10% of the NPSH calculation and must be equal or greater

than 3 ft. If the value is less than 3 ft, then the safety margin must be 3 ft.

6: Column Sizing

The plant consisted of 5 distillation columns in total: 1 main column and 4 side columns.

The main column will distill the crude oil, separating them into useful products:

AGO

Diesel

Naphtha

Kerosene

Turbosene

For instance, Naphtha is used in gasoline and as a petrochemical feedstock. In addition,

kerosene is used as jet fuel.

It is important to size the columns to estimate costs for economic analysis and determine

their measurements for the plant layout. Each column’s length, diameter, thickness, and number

of trays determined the size of the equipment.

6.1 Wall Thickness

To calculate for wall thickness, the equations specified by the ASME BPV Code for hoop

stress (Eq. 3) and for longitudinal stress (Eq. 4) were used.

Eq. 3:

Eq. 4:

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Where: t = wall thickness

Pi = internal pressure

Di = internal diameter

S = maximum allowable stress

E = welded joint efficiency

The values for wall thickness were calculated and compared (See Sample Calculations in

Appendix D). The larger of the two values for each column was chosen as the design wall

thickness. The results are summarized in Table X.

Table X: Column Dimensions

Column Length [ft] Diameter [ft] Thickness [in.] No. of Trays

Main 150 26.5 0.411 25

Turbosene 50 4.5 0.0625 2

Kerosene 50 9 0.127 2

Diesel 50 5 0.0732 3

AGO 50 5 0.0720 3

*Length information was given by E&C Co.

*Diameter values calculated from PRO/II

For more information for each column refer to column sheets Appendix C.

7: Vessel Sizing

The vessels also need to be sized to generate cost estimates for economic analysis and

determine the general vessel proportions for plant layout purposes. Two of the four vessels in

the plant design were sized: the raw crude charge drum and crude flash drum. The other two

vessels were desalters and will be obtained from a desalter package. The raw crude charge

drum was oriented horizontally and will separate the raw crude and water. The crude flash drum

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was oriented vertically and will separate the desalted crude vapor and liquid phases. This

section reports the steps and results for determining the vessel’s dimensions and stream

specifications.

Each vessel’s volume, length, and diameter were determined to specify the vessel’s

shape and size. The typical range of 2 to 5 for the length to diameter ratio was used to

determine an appropriate length and diameter for each vessel.

7.1 Raw Crude Charge Drum Dimensions

The raw crude charge drum’s dimensions were determined using a horizontal cylindrical

segment relationship and the following specifications provided by the E&C Refining Co.:

Horizontal drum

3 feet from HHLL to top of drum

10 minute hold up time

5 minute surge time

Alarms for HHLL and LLLL

12.510 ft3/s flow rate generated from Pro II

The volume of the liquid in the vessel was calculated using the following equation:

Eq. 5:

Where: Q = flow rate

tsurge = surge time

thold up = hold up time

tHHLL = time allowance for HHLL (1 minute)

tLLLL = time allowance for LLLL (1 minute)

The horizontal cylindrical segment equation was used to derive a relationship between vessel

length and diameter (Eq. 6):

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Eq. 6: Where: D = vessel diameter

L = vessel’s tangent to tangent length

V = volume of the liquid in the vessel

Using Eq. 6 with trial and error, the length was calculated by estimating the diameter until the

length to diameter ratio was greater than 2 to less than 5 range (see Sample Calculations in

Appendix D). Table XI summarizes the final dimensions.

Table XI: Final dimensions of the raw crude charge drum

Diameter [ft] Length [ft] Volume [ft3]

17.0 64.0 14,500

*For more information, refer to vessel sheets in Appendix C

7.2 Crude Flash Drum Dimensions

The crude flash drum dimensions were determined using the Souder’s and Brown

correlation and the following specifications provided by the E&C Refining Co.:

16 inch inlet nozzle

10 minute hold up time

5 minute surge time

Alarms for HHLL and LLLL

No demister pad

Pro/II generated the flowrate, density, viscosity, and pressure estimates of each outlet stream.

Table XII summarizes the results

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Table XII: Crude Flash Drum outlet stream properties from PRO/II

Tag No. Phase Flowrate [ft3/s] Density [lb/ft3] Viscosity [cP] Pressure [psig]

2C006 Vapor 122 0.283 0.0108 28.9

2C007 Liquid 11.2 50.5 1.47 28.9

The minimum area and diameter were determined using Table XII and the following equations:

Eq. 7:

Eq. 8:

Eq. 9:

Eq. 10:

Where: P = operating pressure

vt = terminal velocity

ρL = liquid density

ρG = vapor density

The liquid height at HHLL was calculated using Eq. 11.

Eq. 11:

To determine the total height of the vessel, the following allowances for the vessel were added

to the liquid height:

1.00ft for disaster

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3.67ft above liquid height

2.67ft above liquid height

0.500ft for the bottom

The total height was the addition of all the allowance and the calculated liquid height at HHLL.

Using Equation 11 with trial and error, the length was calculated by estimating the diameter

until the length to diameter ratio was within the 2 to 5 range (see Sample Calculations in

Appendix). Table XIII summarizes the final dimensions.

Table XIII: Final dimensions of the crude flash drum

Diameter [ft] Length [ft] Volume [ft3]

15.0 72.5 12,800

*For more information refer to vessel sheets in Appendix C

7.3 Vessel Wall Thickness

To calculate for wall thickness, the equation specified by the ASME BPV Code for hoop stress

(Eq. 3) and for longitudinal stress (Eq. 4) was used. The wall thickness was calculated similarly

to the column wall thickness in 6.1.

Table XIV: Vessel wall thicknesses

Tag. No Vessel Wall Thickness [in.]

V-1 Raw Crude Charge Drum 1.88

V-3 Crush Flash Drum 0.247

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8: Estimated Cost of Equipment

Equipment required for plant operations were priced using Eq. 12 [2] with constants obtained

from Table 7.2 compiled by Gavin Towler and Ray Sinnott [2].

Eq. 12: 𝐶𝑒 = 𝑎 + 𝑏𝑆𝑛

Where: 𝐶𝑒 is the cost of equipment on a U.S. Gulf Coast basis, January 2010

𝑎 and 𝑏 are cost constants found in Table 7.2

𝑆 is the size parameter, for example, mass, flowrate, duty, surface area, etc.

𝑛 exponent for the type of equipment.

Yearly indexes for equipment were obtained from the Chemical Engineering Plant Cost Index [3]

to estimate a current price. Since the plant is to be built in Irvine, the West Coast location factor

obtained from Table 7.7 [2] was utilized. The indexes and location factors are presented in

Table XV where it can be applied to Eq. 13.

Eq. 13: 𝐶𝑜𝑠𝑡 𝑖𝑛 𝑦𝑒𝑎𝑟 2017 = 𝐶𝑜𝑠𝑡 𝑖𝑛 𝑦𝑒𝑎𝑟 2010 ∗𝐶𝑜𝑠𝑡 𝑖𝑛𝑑𝑒𝑥 𝑖𝑛 𝑦𝑒𝑎𝑟 2017

𝐶𝑜𝑠𝑡 𝑖𝑛 𝑦𝑒𝑎𝑟 2010

Table XV: Constants Applied to Calculated Prices

CEPCI January 2010 532.9

CEPCI January 2017 664.1

Location Factor: West Coast 1.07

Distillation column costs were estimated by adding the price of the pressure vessel shell and

valve trays. The size parameter required for the pressure vessel shell was its mass, so it was

priced based on its mass. The valve trays were priced based on diameter. The main column

and side column prices are presented in Table XVI.

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Table XVI: Cost of Distillation Column

The distillation column cost breakdown for shell and trays are presented in Table XVII

Table XVII: Column Cost Breakdown

The costs of vessels are presented in Table XVIII. The costs of shells were estimated based on

mass. The desalter package and atmospheric tower accumulator prices were provided by the

supplier.

Tag No. Equipment Cost Comments

Columns

TWR-1 Crude Column $1,548,149.68

Cost is for shell and trays only.

TWR-2 Turbosene Sidestripper $31,842.65


TWR-3 Kerosene Sidestripper $70,221.93


TWR-4 Diesel Sidestripper $37,654.24


TWR-5 AGO Sidestripper $38,109.15


Total $1,725,977.66

Column Shell Price Tray Price Number of Trays Total

Crude Column $835,210.23 $28,517.58 25 $1,446,868.86

Turbosene Sidestripper $29,338.18 $1,252.24 2 $29,759.50

Kerosene Sidestripper $62,406.00 $3,907.97 2 $65,627.97

Diesel Sidestripper $33,251.11 $1,467.71 3 $35,190.88

AGO Sidestripper $33,706.02 $1,467.71 3 $35,616.03

22

Table XVIII: Cost of Vessels

The cost of single stage centrifugal pumps are presented in Table XIX. The cost estimate was

based on the flowrate of the pump. Though the flowrates for the raw crude charge pump

exceeded the upper limit of the size parameter given in Table 7.2, extrapolation was considered

unnecessary because the function did not increase exponentially beyond the limits.

Table XIX: Cost of Pumps

Tag No. Vessels Cost Comments

V-1 Raw Crude Charge Drum $950,842.95 Cost is for shell only.

V-3 Crude Flash Drum $181,774.55 Cost is for shell only.

V-4 Desalter Package $3,745,000.00

V-5 Atmospheric Tower Accumulator $481,500.00

Total $5,359,117.50

Tag No. Pumps Cost Comments

P-1 Raw Crude Charge Pump

$90,401.98

Cost is for 2 pumps, one operating and one spare.

P-2 Flashed Crude Charge Pump

$21,822.57


P-3 Reflux Pump $44,461.2

0 Cost is for 2 pumps, one operating and one spare.

P-4 Naphtha Product Pump $38,769.7


P-5 Jet/Turbo Product Pump $32,967.9


P-6 Kerosene Product Pump

$34,272.20


P-7 Diesel Product Pump $30,866.2


P-8 Gas Oil Product Pump $26,959.6


P-9 Atmospheric Resid Pump

$59,567.48


P-10 Top PA Pump $374,500.


P-11 Turbosene PA Pump $214,000.


P-12 Diesel PA Pump $321,000. Cost is for 2 pumps, one operating and one

23

The cost of for the U-tube shell and tube heat exchangers are presented in Table XX. Heat

exchangers were priced were based on contact surface area. If the area required for a service

exceeded 5000 ft2 additional shells were purchased to divide the load. The number of shells

required was based on the heat integration. The streams that required air and water cooling

were priced based on a local cost factor of $10,000 per MMBtu/hr. The breakdown costs for

each stream is presented in Table XXI.

Table XX: Cost of Heat Exchangers

Tag No.

Heat Exchangers

Cost Comments

WC Cooling Water Cost

$1,021,776.00

Cost is to cool atm, turbo, diesel, diesel PA and naptha.

PH-1 Preheat 1 $18,333,243.54

PH-2 Preheat 2 $1,670,702.47

PH-3 Preheat 3 $33,143,981.09

Total $54,241,227.42

Table XXI: Heat Exchanger Cost Breakdown

Type Service Cost per Shell Number of Shells Cost

Water Cooling Atm Resid $29,262.00

Turbosene $149,312.00

Diesel $155,166.00

Diesel PA $634,317.00

Naphtha $53,719.00

Subtotal $1,021,776.00

Preheat 1 AGO $900,535.64 1 $900,535.64

00 spare.

Total $1,289,589.00

24

Kerosene $1,021,304.96 2 $2,042,609.92

Top PA $1,826,124.72 3 $5,478,374.15

Turbo PA $1,389,564.59 3 $4,168,693.76

Turbosene $965,550.42 2 $1,931,100.84

Diesel $1,835,785.06 1 $1,835,785.06

Kerosene $988,072.09 2 $1,976,144.17

Subtotal $18,333,243.54

Preheat 2 AGO $1,670,702.47 1 $1,670,702.47

Subtotal $1,670,702.47

Preheat 3 Atm Resid $1,949,645.95 17 $33,143,981.09

Subtotal $33,143,981.09

Total $54,169,703.10

The total cost summary is presented in Table XXII.

Table XXII: Total Cost Summary

Tag No. Equipment Cost Comments

Columns

TWR-1 Crude Column

$1,548,149.68 Cost is for shell and trays only.

TWR-2 Turbosene Sidestripper $31,842.66 Cost is for shell and trays only.

TWR-3 Kerosene Sidestripper $70,221.93 Cost is for shell and trays only.

TWR-4 Diesel Sidestripper $37,654.24 Cost is for shell and trays only.

TWR-5 AGO Sidestripper $38,109.15 Cost is for shell and trays only.

Vessels

V-1 Raw Crude Charge Drum $181,774.56 Cost is for shell only.

V-3 Crude Flash Drum $950,842.95 Cost is for shell only.

V-4 Desalter Package $3,745,000.

00

V-5 Atmospheric Tower Accumulator $481,500.00

Pumps

P-1 Raw Crude Charge Pump $90,401.98 Cost is for 2 pumps, one operating and one spare.

P-2 Flashed Crude Charge Pump $21,822.57


P-3 Reflux Pump $44,461.20 Cost is for 2 pumps, one operating and one

25

spare.

P-4 Naphtha Product Pump $38,769.71 Cost is for 2 pumps, one operating and one spare.

P-5 Jet/Turbo Product Pump $32,967.95 Cost is for 2 pumps, one operating and one spare.

P-6 Kerosene Product Pump $34,272.20 Cost is for 2 pumps, one operating and one spare.

P-7 Diesel Product Pump $30,866.21 Cost is for 2 pumps, one operating and one spare.

P-8 Gas Oil Product Pump $26,959.69 Cost is for 2 pumps, one operating and one spare.

P-9 Atmospheric Resid Pump $59,567.48 Cost is for 2 pumps, one operating and one spare.

P-10 Top PA Pump $374,500.00 Cost is for 2 pumps, one operating and one spare.

P-11 Turbosene PA Pump $214,000.00 Cost is for 2 pumps, one operating and one spare.

P-12 Diesel PA Pump $321,000.00 Cost is for 2 pumps, one operating and one spare.

Heat Exchangers

WC Cooling Water Cost $1,021,776.

00 Cost is to cool atm, turbo, diesel, diesel PA and naptha.

PH-1 Preheat 1 $18,333,243

.54

PH-2 Preheat 2 $1,670,702.

47

PH-3 Preheat 3 $33,143,981

.09

Total $62,544,387.26

8.1 Estimated Plant Cost

The plant cast was estimated by determining ISBL fixed capital cost of a plant (FCI)

based off of the equipment cost and delivery. By using Equation 14 [2] the ISBL fixed capital

cost can be obtained.

Eq. 14 𝐶 = 𝐹(Σ𝐶𝑒)

Where 𝐶 is the total plant ISBL capital cost including engineering costs

𝐹 is the installation factor or the Lang factor.

Σ𝐶𝑒 is the total delivered cost of all the major equipment items.

Table XXIII outlines the calculation of the total plant ISBL capital cost. This includes a delivery

fee that is 10% of the total equipment purchase cost.

26

Table XXIII: Total ISBL Capital Cost

Component Cost

Purchase Equipment $62,544,387.26

10% Delivery $6,254,438.73

Lang Factor 4.74 [1]

Fixed-Capital Investment $326,106,435.20

8.2 Estimated Return on Investments

Based on the market prices for the main saleable products naphtha, turbosene, and

kerosene, as shown in Table XXIV, the gross income can be calculated and a return on

investment time can be estimated.

Table XXIV: 2017 Market Value of Main Products

Product 2017 Market Value

Naptha $522 per metric ton

Turbosene $1.51 per gallon

Kerosene $1.58 per gallon

Diesel $2.97 per gallon

At the flowrates specified in the H&MB for the product streams the gross income per day for

each product stream is shown in Table XXV.

Table XXV Gross Income per Day at Full Capacity

Product Gross Income

Naphtha $2,055,418.84 per day

Turbosene $1,214,970.16 per day

Kerosene $1,222,172.66 per day

Diesel $1,866,266.91 per day

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With these rates of revenue, the capital costs of purchasing equipment can be recovered in 9.84

days and the ISBL fixed capital costs of the plant can be recovered in 51.28 days.

9: Risk Assessment and Adverse Health Effects

Occupational hazards are plentiful in raw crude distillation refineries which consist of

exposure, fires, employee fatigue, etc. In the case study of the Texas City Refinery explosion in

2005, the tragedy surmounted to understaffing, neglect of routine maintenance, and exposure to

hydrocarbon vapors [4]. The Chemical Safety Board (CSB) and Occupational Safety and

Health Administration (OSHA) enforce strict regulations to ensure refineries are safe for workers

and the public. All hydrocarbon products that will be produced in the E&C Refining Co. in Irvine

have chemical hazards and potential fire outbreaks.

The raw crude and its byproducts also causes skin and eye irritation. Long-term

exposures to any of the hydrocarbons are carcinogenic [5]. It is recommended that safety

goggles and protective gear are utilized if handling products directly. Inhalation of the

hydrocarbon vapors are also a safety issue. Breathing the vapors can cause dizziness,

headaches, and nausea. In rare cases, the vapors can render a worker unconscious. If

exposed to the product vapor, move to a well ventilated area and avoid long term exposure.

Raw crude and its byproducts are highly flammable. The hydrocarbon flash points can

be as low as 100°F. The distillation process will be handling product streams that are well

above this and extreme caution is needed. The refining process is in a closed system but leaks

can occur and ignite. Some ignition sources are heaters such as the HXR-1 heat exchanger,

sparks from pumps, and external fires [6]. In the case study of the Tosco Avon Refinery in 1999

by the US Chemical Safety and Hazard Investigation Board (2001), the ignition source was the

hot surface of the fractionator [7]. Personnel were attempting to isolate and repair a naphtha

leak from a pipe while the refinery was in operation. After numerous attempts, frustration and

impaired judgment ensued. A cut was made on the pipe, but it was not properly drained. The

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maintenance supervisor ordered to open a flange to drain the line, but naphtha suddenly

backfired and spilled out from the cut made previously. Naphtha was ignited by the fractionator

and resulted in 4 deaths.

Following guidelines by the CSB and OSHA is imperative for safety of personnel and

public alike. Proper protocols can be found in the General EHS Guidelines given by the

International Finance Corporation and on the OSHA website.

10: Environmental Impacts

The chemical plant can emit several toxic airborne pollutants. These typically include:

nitrogen oxides (NOX)

sulfur oxides (SOX)

carbon dioxide (CO2)

volatile organic compounds (VOCs)

particulates (PM10 and PM2.5)

Polluting greenhouse gas (e.g. CO2) induced global warming by trapping infrared radiation within

Earth’s surface, thus increasing the Earth’s temperature [1]. The three main pollutants consist of

NOX, SOX, and CO2. NOX emissions can be controlled using many methods. NOX can be

managed using selective catalytic reduction (SCR) with NH3. Although expensive, flue gas

scrubbing (FGD) can be applied to SOX to monitor its emissions. Amine scrubbing can also be

used to capture CO2 emissions.

The plant’s location poses human safety hazards as well. Being situated in Irvine, the plant is

near thousands of people. Oil refinery plants typically carry noisy machinery (e.g. pumps,

compressors, and turbines) to facilitate petroleum processing and mass distribution [2]. Due to

the refinery’s proximity to residences, noise pollution can impair residents’ hearing. Oil

refineries, as previously mentioned, carry highly flammable hydrocarbons that can lead to

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explosions. Improper handling of materials and machinery can consequently decimate the

surrounding area and lead to multiple deaths and injuries.

11: Failure Mode Effect Analyses (FMEA) on Crude Flash Drum

HAZOP of the crude flash drum was done to provide some of the hazards that can be

produced and affect the operations of the refinery.

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Conclusion

The design of the crude distillation unit for processing 165,000 BPSD was achieved. The

necessary equipment was designed appropriately. Three heat exchangers and a heater were

designed using a heat integration curve to get the product streams to the desired temperatures

while minimizing utility usage. Information about the pumps were specified to ensure the

procurement of sufficient pumps. A column was added and sized using PRO/II to separate the

components of the crude oil. The dimensions for the raw crude charge and crude flash drum

were determined using the horizontal cylinder relationship and souder and brown correlation

respectively. Overall, the estimated fixed capital investment is $62 million which will be

recovered in 10 days of the plant operating at full capacity. The ISBL fixed capital costs are

estimated at $326 million which can be recovered within 2 months when at full capacity. Those

working in the plant are at risk to health effects from exposure to the present chemicals and

injury from equipment malfunction. The risks can be reduced by following the CSB and OSHA

guidelines. The plant can release toxic pollutants such as SOx whose emission can be controlled

by implementing techniques such as flue gas and amine scrubbing. A HAZOP report is provided

to anticipate potentially hazardous scenarios.

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References

[1] C. Ruiz, "Process Hydraulics and Fluid Handling Equipment", UC Irvine, 2017.

[2] G. Towler and R. Sinnott, Chemical engineering design, 1st ed. Amsterdam: Butterworth-

Heinemann, 2013.

[3] Economic Indicators, WWW.CHEMENGONLINE.COM, 2017.

[4] G. Poje, I. Rosenthal, A. Taylor. Investigation Report Refinery Fire Incident: Tosco Avon

Refinery. US Chemical Safety and Hazard Investigation Board, 99-014-I-CA, 2001, p. 1-

70.

[5] U. NOAA Office of Response and Restoration, "CRUDE OIL, PETROLEUM,

[COMBUSTIBLE LIQUID LABEL] | CAMEO Chemicals | NOAA",

Cameochemicals.noaa.gov, 2017. [Online]. Available:

https://cameochemicals.noaa.gov/chemical/11018. [Accessed: 03- Jun- 2017].

[6] "OSHA Technical Manual (OTM) | Section IV: Chapter 2 - Petroleum Refining Process |

Occupational Safety and Health Administration", Osha.gov, 2017. [Online]. Available:

https://www.osha.gov/dts/osta/otm/otm_iv/otm_iv_2.html. [Accessed: 03- Jun- 2017].

[7] Investigation Report Refinery Explosion and Fire: BP. US Chemical Safety and Hazard

Investigation Board, 2005-04-I-TX, 2007, p. 1-341.

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