vulcan sgcu coal to liquids case study

8/10/2019 Vulcan SGCU Coal to Liquids Case Study

http://slidepdf.com/reader/full/vulcan-sgcu-coal-to-liquids-case-study 1/24

Refinery Process Stream Purification Refinery Process Catalysts Troubleshooting Refinery Process Catalyst Start-Up / Shutdown

Activation Reduction In-situ Ex-situ Sulfiding Specializing in Refinery Process Catalyst Performance Evaluation Heat & MassBalance Analysis Catalyst Remaining Life Determination Catalyst Deactivation Assessment Catalyst PerformanceCharacterization Refining & Gas Processing & Petrochemical Industries Catalysts / Process Technology - Hydrogen Catalysts /

Process Technology – Ammonia Catalyst Process Technology - Methanol Catalysts / process Technology – PetrochemicalsSpecializing in the Development & Commercialization of New Technology in the Refining & Petrochemical Industries

Web Site: www.GBHEnterprises.com

GBH Enterprises, Ltd.

SYNGAS CONDITIONING UNIT FEASIBILITY

CASE STUDY: COAL-TO-LIQUIDSCase Study: #0953616GB/H

Process Information Disclaimer

Information contained in this publication or as otherwise supplied to Users isbelieved to be accurate and correct at time of going to press, and is given ingood faith, but it is for the User to satisfy itself of the suitability of the Product forits own particular purpose. GBHE gives no warranty as to the fitness of theProduct for any particular purpose and any implied warranty or condition(statutory or otherwise) is excluded except to the extent that exclusion isprevented by law. GBHE accepts no liability for loss, damage or personnel injurycaused or resulting from reliance on this information. Freedom under Patent,Copyright and Designs cannot be assumed.







COAL-TO-LIQUIDS FEASIBILITY PROJECT CASE STUDY

HT SHIFT REACTOR CATALYST SPECIFICATION

Process Specification

This process duty specification refers to a Syngas Conditioning Unit whichutilizes HT Shift reaction technology on a slip stream of raw gas to produce arecombined gas stream with a H2

:CO ratio of 1.57:1. This is an importantconsideration as the Shift reactor is not required to minimize CO at outlet, andthis specification refers to the expected performance that can be achieved in asingle stage reactor scheme.

The Syngas Conditioning Unit is part of a proposed coal-to-liquids complex inwhich synthesis gas is produced by gasification of coal for downstream

processing in a Fischer Tropsch reactor and Hydrocracker unit.

Basis of Design

It is anticipated that a number of parallel trains will be required, in view of thelarge gas flow rate considered. For the purposes of this preliminary design, amaximum allowable catalyst bed diameter of 4.5m should be assumed. Thecatalyst vendor will advise the number of parallel trains required on this basis.

Our client has requested that we evaluate two competing options for integration

of the HT Shift reactor into the overall coal-to-liquids process scheme. Twoalternative HT Shift feedstock options are therefore under consideration, (A andB) as defined below (Sour and Sweet shift respectively):







End of Run Conditions

A B

Inlet Temperature (o

191C) 196Inlet Pressure (bara) 40 40

Inlet Flow (kmol/h) 159,208 151,439

Feedstock Composition

Component Mol % Mol %HydrogenCarbon MonoxideCarbon DioxideNitrogenMethaneEthane

AmmoniaH2SHCNCOS

Water

13.0033.662.700.500.04-0.010.070.010.01

50.0

13.7535.600.100.500.05-----

50.0







Catalyst Performance Specif ication

END of Run CO Slip (Dry

basis)

10%

max

10%

max Available Pressure Drop(bar)

0.25 (1) 0.25 (1)

Run Life (years) 2 2

1. Vessel nozzle to nozzle

The feedstock specification provided refers to end of run conditions based on thefollowing assumed approach to equilibrium:

Feedstock A (Sour Shift): 40°C approach to equilibrium at EORFeedstock B (Sweet Shift): 24°C approach to equilibrium at EOR

At SOR conditions, the flowrate of raw gas that bypasses the Shift reactor will beincreased to ensure that the H2

:CO ratio of the recombined gas can bemaintained at 1.57:1.

Data requested from GBHE

Catalyst type/sizeCatalyst bed volume and ordered volumeRecommended bed dimensions and number of parallel trains requiredMaximum gas flowrate through each bedOutlet gas composition and pressure drop at start and end of run conditionsPerformance GuaranteesCatalyst Cost







CTL FEASIBILITY CASE STUDY - RESULTS

FEEDSTOCK A – SOUR SHIFT CATALYST: VIG SGS201

Preliminary calculations have shown that the Shift Sections must consist ofseveral parallel trains. A number of 5 trains results in reasonable Bed Volumes.Please note all Flow Rates stated below are per train. The following Cases wereconsidered:

Case: Client Original GBHE Required

Feed Flow Rate (kmol/h, wet): 31841.80 40579.28

Feed Flow Rate (Nm3 713697.6/h, wet): 909543.9

Feed Flow Rate (kmol/h, dry): 15920.80

Feed Flow Rate (Nm3

356848.8/h, dry):

Steam Flow Rate (kmol/h): 15920.80 26802.69Steam Flow Rate (Nm3 356848.8/h): 600755.6

Steam/CO Ratio (mol/mol: 1.485 2.500

Steam/Dry Gas Ratio (mol/mol): 1.000 1.684

Catalyst Bed Diameter D (m): 4.50

Feed Composition (mol%, dry):HydrogenCarbon MonoxideCarbon DioxideNitrogen

MethaneNH3H2SHCNCOSTOTAL

26.0067.325.401.00

0.080.020.140.020.02 (= 200 ppmv)100.00

Inerts (mol%, dry) 1.28

H2S/COS Ratio (mol/mol): 7.00

Inlet Operating Pressure (bara): 40 (= 38.987 barg)

Interbed Cooling: Heat ExchangerInlet Operating Temperature (

o300C)

Exit CO Slip (mol%, dry) 10.00







The Client’s original design is not acceptable due to both the Steam/CO Ratioand the Steam/Dry Gas Ratio specified being too low. For an Inlet OperatingPressure First Reactor of 40 bara and an Inlet CO Level exceeding 40 mol%(dry) a Steam/CO Ratio of at least 2.500 mol/mol is required, which then resultsin a Steam/Dry Gas Ratio of 1.684 mol/mol, which is acceptable. These valuesfor the Steam/Dry Gas Ratio and the Steam/CO Ratio are required to avoid therisk of methanation. The recommended steam level will not cause any desulfidingof the catalyst at the specified inlet H2S Level.

Case: GBHE EOR GBHE SOR

Feed Flow Rate (kmol/h, wet): 40579.28Feed Flow Rate (Nm

3909543.9/h, wet):


Feed Flow Rate (Nm3

356848.8/h, dry):Expected Exit Operating

Temperature (

o495.0

C):

510.0

Expected Catalyst BedPressure Drop (bar):

4.70 (2years) 3.61 (2 years)

Guaranteed Catalyst Life(Years):

2 N.A.

Gas Exit Composition (mol%,dry):Inerts:H2COCO2

TOTAL

0.8451.3510.0037.81100.00

0.8252.677.0239.49100.00

Exit Steam: Dry Gas Ratio(mol/mol):

0.7646 0.7168

Guaranteed Maximum CO Slip(mol%, dry):

10.00 N.A.

Expected Maximum CO Slip(mol%, dry):

N.A. 7.02

Catalyst Bed Volume (m3 87.80):


Catalyst Bed Height H (m): 5.52

Number of Parallel Trains

Required (-):

5

Amount of Catalyst to beOrdered (m3

5 x 1.03 x 87.80 = 452.2):







FEEDSTOCK B – SWEET SHIFT: VULCAN SERIES VSG-F101

Preliminary calculations showed that the original case (provided by client)produced an Exit Operating Temperature of the first reactor of around 560

o

C,which is far too high. As such the following two bed option was considered.

Case GBHE Design

Feed Flow Rate (kmol/h, wet): 36345.36

Feed Flow Rate (Nm3

814644.9/h, wet):

Minimum Acceptable CatalystBed Volume for Each Reactor(m

3

101.8

)(Order: 2 x 5 x 1.03 x 101.8 =1049)


Catalyst Bed Height H (m): 6.40H/D 1.42

Feed Flow Rate (kmol/h, dry): 15143.90Steam Flow Rate (kmol/h): 21201.46

Steam/Dry Gas Ratio (mol/mol): 1.40

Feed Composition (mol%, dry):HydrogenCarbon MonoxideCarbon DioxideNitrogenMethaneNH3H2S

HCNCOSTOTAL

27.5071.200.201.000.100.00

0.000.000.00100.00

Case B First Reactor EOR

Bed 1

Expected Exit Operating Temperature(o

531.4C):

Expected Catalyst Bed Pressure Drop(bar):

1.4 (2 years)

Guaranteed Catalyst Life (Years): 2







Gas Exit Composition (mol%, dry):CH4H2CO

CO2N2TOTAL

0.0749.3419.61

30.270.70100.00

Exit Steam: Dry Gas Ratio (mol/mol): 0.677

Guaranteed Maximum CO Slip (mol%,dry):

19.61

Bed 2

Optimized Inlet Operating Temperature(o

385.5C):


422.3C):

Expected Catalyst Bed Pressure Drop(bar): 1.5 (2 years)

Guaranteed Catalyst Life (years): 2Gas Exit Composition (mol%, dry):CH4H2COCO2N2TOTAL

0.0654.786.7737.760.62100.00


Guaranteed Maximum CO Slip (mol%,dry):

6.77

Case B First Reactor SOR

Bed 1


541.5C):


1.4 (SOR)

Gas Exit Composition (mol%, dry):CH4H2

COCO2N2TOTAL

0.0653.51

9.7836.010.64100.00








Expected Maximum CO Slip (mol%, dry): 9.78Bed 2

Optimized Inlet Operating Temperature(o

364.7C):

Expected Exit Operating Temperature(o 397.3C):


1.3 (SOR)

Gas Exit Composition (mol%, dry):CH4H2COCO2N2TOTAL

0.0655.704.5739.040.61100.00

Exit Steam: Dry Gas Ratio (mol/mol): 0.466Expected Maximum CO Slip (mol%, dry): 4.57

The only way to reduce the Catalyst Bed Pressure Drops for both sour and sweetshift will be to increase the Catalyst Bed Diameter. The choice of a largerMaximum Allowable Catalyst Bed Diameter may imply local manufacture of thereactors.

One way to reduce the Catalyst Bed Volumes required would be to use radialreactors. At the same time GBHE could then consider the use of mini catalyst,

because the Catalyst Bed Pressure Drops in radial reactors tend to be lower thanin the corresponding axial reactors. The use of mini catalyst in axial reactors isclearly out of the question because of the unacceptably high Catalyst BedPressure Drops that would result.







APPENDIX







HTS Section (Sweet Shift)

Starting Point : Preliminary calculations have shown that the HTS Section(Sweet Shift) must consist of several parallel trains and that due to equilibriumrestrictions each train needs to have 2 HTS reactors in series called HT1 andHT2, respectively (with intercooling) in order to be able to achieve the requiredCO Slip. A number of 5 trains results in reasonable Bed Volumes. In thisdocument only the Sweet Shift option will be considered (Case B). Please noteall Flow Rates stated below are per train (unless specifically mentionedotherwise). The following Cases were considered:

1/Cases Considered

Type of Shift Life InletOperatingPressureFirstReactor(bara)

InletOperatingTemperatureFirst Reactor(oC)

Case BSOR

Sweet Shift SOR Expected 40 310*

Case BEOR

Sweet Shift 2 YearsGuaranteed

40 310*

*In order to minimize the Exit Operating Temperature of the first HTS reactor theminimum allowable Inlet Operating Temperature was chosen for this reactor. TheInlet Operating Inlet Temperature of the second reactor was optimized, becausein the second reactor the exotherm was much less severe than in the first HTSreactor.

2/Catalyst Used: VULCAN SERIES VSG-F101

3/Client Requirements:

Required Guaranteed Maximum Catalyst Life(years):

2

Required Guaranteed Maximum CO Slip EOR(i.e. 2 years, mol%, dry):

10.00

Required Guaranteed Total MaximumCatalyst Bed Pressure Drop for Each Train(EOR (bar):

0.25*

Required Maximum Allowable Catalyst BedDiameter (m):

4.5*







*Preliminary calculations have shown that the Required Guaranteed TotalMaximum Catalyst Bed Pressure Drop for Each Train cannot at all be achievedwith the Required Maximum Allowable Catalyst Bed Diameter (Unless the Clientwould accept a ridiculously large number of parallel trains, say 10 to 15, anumber of 5 appears to be the maximum number that is still consideredacceptable by our Clients in general). A larger Maximum Allowable Catalyst BedDiameter should be considered. The choice of a larger Maximum AllowableCatalyst Bed Diameter may imply local manufacture of the reactors.

4/Feed Flow Rates per Train and Feed Composition Inlet First Reactor;Minimum Acceptable Catalyst Bed Volume for Each Reactor and OtherReactor Dimensions

Case B1* B2 B3

Feed FlowRate (kmol/h,wet):

30287.80 33316.58 36345.36

Feed FlowRate (Nm

3678870.7

/h,wet):

746757.8 814644.9

Minimum AcceptableCatalyst BedVolume forEach Reactor(m

3

84.9

)**:

(Order: 2 x 5 x1.03 x 84.9 = 874)

93.3(Order: 2 x 5 x1.03 x 93.3 = 961)

101.8(Order: 2 x 5 x1.03 x 101.8 =1049)

Catalyst BedDiameter D(m):

4.50***

Catalyst BedHeight H (m):

5.34 5.87 6.40

H/D (-)****: 1.19 1.30 1.42Feed FlowRate (kmol/h,dry):

15143.90

Steam FlowRate (kmol/h):

15143.90 18172.68 21201.46

Steam/Dry GasRatio(mol/mol):

1.00 1.20 1.40







FeedComposition(mol%, dry):Hydrogen

CarbonMonoxideCarbon DioxideNitrogenMethaneNH3H2SHCNCOSTOTAL

27.5071.20

0.201.000.100.000.000.000.00100.00

*Case B1 is the original case as submitted by the Client. However, preliminarycalculations showed that this original case produced an Exit OperatingTemperature of the first reactor of around 560

oC, which is far too high. This

temperature needs to be brought down to at least 540o

C in accordance with ourdesign rules. The only way to do this is by increasing the Steam/Dry Gas Ratio.This may not be achievable. The Client will probably also have to faceconsiderable material of construction problems/issues in view of the ratherdemanding temperature regime.

**These Catalyst Bed Volumes were calculated using in accordance with ourdesign rules (Proprietary).

*** Required Maximum Allowable Catalyst Bed Diameter.

****The H/D values found are all greater than 0.6 and are therefore acceptableaccording to our design rules.

*****This Inlet CO Level is very high and makes the overall duty very demanding.







Results:

The Total Wet Flow Rate throughout each train does not change. For Cases BSOR and B EOR an Initial WGS Activity Factor was chosen (High TemperatureShift Catalyst Activity for Partial Oxidation Plants), Proprietary.

Case B1 EOR B2 EOR B3 EOR

Bed 1: HT1 (HTS)

Expected Exit OperatingTemperature (

o559.2

C):550.8 531.4

Expected Catalyst Bed PressureDrop (bar):

0.908 (2years)

1.148 (2years)

1.422 (2years)

Guaranteed Catalyst Life (Years): 2

Gas Exit Composition (mol%, dry):CH4

H2COCO2N2TOTAL

0.07

51.2615.0832.920.67100.00

0.07

51.8713.6533.750.66100.00

0.07

49.3419.6130.270.70100.00


0.344 0.460 0.677


15.08 13.65 19.61

Reduction Potential (-): 1.515 1.263 1.083

Bed 2: HT2 (HTS)

Optimized Inlet OperatingTemperature (

o 363.4C):

375.5 385.5


o411.5

C):413.0 422.3


0.872(2years)

1.161 (2years)

1.525 (2years)

Guaranteed Catalyst Life (years): 2Gas Exit Composition (mol%, dry):CH4H2CO

CO2N2TOTAL

0.0653.998.57

36.710.63100.00

0.0654.677.00

37.630.63100.00

0.0654.786.77

37.760.62100.00


0.269 0.375 0.497








8.57 7.00 6.77


Case B1 SOR B2 SOR B3 SOR

Bed 1: HT1 (HTS)Expected Exit OperatingTemperature (o

559.2C):

551.3 541.5


0.864 (SOR) 1.103 (SOR) 1.377 (SOR)

Gas Exit Composition (mol%, dry):CH4H2COCO2N2

TOTAL

0.0751.2715.0532.930.67

100.00

0.0752.5412.0734.670.65

100.00

0.0653.519.7836.010.64

100.00Exit Steam: Dry Gas Ratio(mol/mol):

0.344 0.440 0.539


15.05 12.07 9.78

Reduction Potential (-): 1.515 1.263 1.083Bed 2: HT2 (HTS)

Optimized Inlet OperatingTemperature (

o342.5

C):355.5 364.7


o397.4

C):398.2 397.3


0.754 (SOR) 1.004 (SOR) 1.303 (SOR)

Gas Exit Composition (mol%, dry):CH4H2COCO2N2TOTAL

0.0654.457.4737.360.63100.00

0.0655.205.7238.380.62100.00

0.0655.704.5739.040.61100.00


0.256 0.359 0.466


7.47 5.72 4.57








Comments

Cases B3 EOR and B3 SOR show an Expected Exit Operating Temperature ofaround 540

o

C. However, the extra steam required has increased the CatalystBed Pressure Drops.

A larger Catalyst Bed Diameter will be required to achieve the RequiredGuaranteed Total Maximum Catalyst Bed Pressure Drop for Each Train.For all Cases the value for the Reduction Potential R is so low (all lower than1.9), that over-reduction of the HTS can be safely ruled out.

Using optimized Operating Inlet Temperatures could only be applied for thesecond bed, because of the very high exotherm in the first bed.

Of course the steam content must always stay below dew point level. Under thecurrent conditions there is no risk of steam condensation (See table below).

The Required Guaranteed Maximum CO Slip EOR (2 years) of 10.0 mol% (dry)can be achieved.

One way to reduce the Catalyst Bed Volumes required would be to use radialreactors. At the same time we could then consider the use of mini catalyst,because the Catalyst Bed Pressure Drops in radial reactors tend to be lower thanin the corresponding axial reactors. The use of mini catalyst in axial reactors isclearly out of the question because of the unacceptably high Catalyst BedPressure Drops that would result.

Case InletOperatingTemperatureHT1 (

o

Inlet DewPoint HT1(

C)

o

InletOperatingTemperatureHT2 (

C)o

Inlet DewPoint HT2(

C)

oC)

B1 EOR 310 212.4 363.4 179.8B1 SOR 212.4 342.5 179.9

B2 EOR 216.8 375.5 188.8

B2 SOR 216.8 355.5 187.5

B3 EOR 220.3 385.5 199.9

B3 SOR 220.3 364.7 193.4







Sour Shift Section

Starting Point : Preliminary calculations have shown that the Sour Shift Sectionmust consist of several parallel trains. A number of 5 trains results in reasonableBed Volumes. In this document only the Sour Shift option will be considered(Case A). Please note all Flow Rates stated below are per train (unlessspecifically mentioned otherwise). The following Cases were considered:

1/Cases Considered

Type ofShift

Life Inlet OperatingPressure (bara)

Inlet OperatingTemperature(oC)

Case ASOR

Sour Shift SOR Expected 40 300

Case AEOR

Sour Shift 2 YearsGuaranteed

40 300

2/Catalyst Used: VULCAN Series VIG SGS201. The use of this catalyst isallowed here, because the Inlet Operating Pressure is only 40 bara.

3/Client Requirements:

Required Guaranteed Maximum Catalyst Life (years): 2

Required Guaranteed Maximum CO Slip EOR (i.e. 2 years, mol%,dry): 10.00

Required Guaranteed Total Maximum Catalyst Bed Pressure Dropfor Each Train EOR (bar):

0.25*

Required Maximum Allowable Catalyst Bed Diameter (m): 4.5*

*Preliminary calculations have shown that the Required Guaranteed TotalMaximum Catalyst Bed Pressure Drop for Each Train cannot at all be achievedwith the Required Maximum Allowable Catalyst Bed Diameter (Unless the Clientwould accept a ridiculously large number of parallel trains, say 10 to 15, anumber of 5 appears to be the maximum number that is still considered

acceptable by our Clients in general). A larger Maximum Allowable Catalyst BedDiameter should be considered. The choice of a larger Maximum AllowableCatalyst Bed Diameter may imply local manufacture of the reactors.







4/Oxygen and Olefins.

There are no oxygen or olefins present in the feed, which would have resulted inan additional exotherm and a change in the composition of the feed.

5/HCN

The catalyst hydrogenates 90% of the HCN present in the feed.

6/NH3

The NH3 Level present in the feed will not affect the performance of the catalyst.

7/Arsine

No mention is made of the presence of arsine. Arsine is a severe catalyst poison.In view of the fact that coal is the starting point in this process it would be wise todouble-check whether there is really no arsine whatsoever present in the feed ofthe Sour Shift Section.

8/Input Data Sour Shift Unit

Case: A1* A2**

Feed Flow Rate (kmol/h, wet): 31841.80 40579.28Feed Flow Rate (Nm3 713697.6/h, wet): 909543.9


Feed Flow Rate (Nm3

356848.8/h, dry):

Steam Flow Rate (kmol/h): 15920.80 26802.69

Steam Flow Rate (Nm3

356848.8/h): 600755.6Steam/CO Ratio (mol/mol: 1.485 2.500

Steam/Dry Gas Ratio (mol/mol): 1.000 1.684

Catalyst Bed Diameter D (m): 4.50***







Feed Composition (mol%, dry):HydrogenCarbon MonoxideCarbon Dioxide

NitrogenMethaneNH3H2SHCNCOSTOTAL

26.0067.325.40

1.000.080.020.140.020.02 (= 200 ppmv)100.00

Inerts (mol%, dry)****: 1.28

H2S/COS Ratio (mol/mol): 7.00

Inlet Operating Pressure (bara): 40 (= 38.987 barg)Interbed Cooling: Heat Exchanger

Inlet Operating Temperature (

o

300C)*****:Exit CO Slip (mol%, dry) 10.00

*Case A1 is the original case as submitted by the Client. Please note that boththe Steam/CO Ratio and the Steam/Dry Gas Ratio specified by the Client are toolow. For an Inlet Operating Pressure First Reactor of 40 bara and an Inlet COLevel exceeding 40 mol% (dry) a Steam/CO Ratio of at least 2.500 mol/mol isrequired, which then results in a Steam/Dry Gas Ratio of 1.684 mol/mol, which isacceptable. These values for the Steam/Dry Gas Ratio and the Steam/CO Ratioare required to avoid the risk of methanation. The recommended steam level willnot cause any desulfiding of the catalyst at the specified inlet H2S Level (it is

borderline though).

**Case A2 was actually used in the calculations.

*** Required Maximum Allowable Catalyst Bed Diameter.

****Inerts (mol%, dry) is defined as 100 – CO Level (mol%, dry) – CO2 Level(mol%, dry) – H2 Level (mol%, dry).

*****Preliminary calculations demonstrated that for this duty only 1 bed per trainis required.







10/Activity (Proprietary)

Results:

The Total Wet Flow Rate throughout each train does not change.

Downflow and axial reactors have been assumed.

The details are given below:

Case: A2 EOR A2 SOR

Feed Flow Rate (kmol/h, wet): 40579.28Feed Flow Rate (Nm

3909543.9/h, wet):


Feed Flow Rate (Nm

3

356848.8/h, dry):Expected Exit OperatingTemperature (o

495.0C):

510.0


0.90* (2 years) 0.75 (2 years)

Guaranteed Catalyst Life (Years): 2 N.A.

Gas Exit Composition (mol%, dry):Inerts***:H2COCO2

TOTAL

0.8451.3510.0037.81

100.00

0.8252.677.0239.49

100.00Exit Steam: Dry Gas Ratio(mol/mol):

0.7646 0.7168


10.00 N.A.


N.A. 7.02

Catalyst Bed Volume (m3 87.80):


Catalyst Bed Height H (m): 5.52

Number of Parallel Trains

Required (-):

5

Amount of Catalyst to be Ordered(m3

5 x 1.03 x 87.80 = 452.2):







*The only way to reduce the Catalyst Bed Pressure Drop will be to increase theCatalyst Bed Diameter. The choice of a larger Maximum Allowable Catalyst BedDiameter may imply local manufacture of the reactors

Comments

Cases A2 EOR and A2 SOR show an Expected Exit Operating Temperature ofaround 500o

C, which is acceptable.

A larger Catalyst Bed Diameter will be required to achieve the RequiredGuaranteed Total Maximum Catalyst Bed Pressure Drop for Each Train. Anormal value would have been 0.1 – 0.2 bar. The values found here are far toohigh.

Of course the steam content must always stay below dew point level. Under thecurrent conditions there is no risk of steam condensation, because the dew pointat the inlet is only 224.1o

C.

The Required Guaranteed Maximum CO Slip EOR (2 years) of 10.0 mol% (dry)can be achieved.

One way to reduce the Catalyst Bed Volumes required would be to use radialreactors. At the same time we could then consider the use of mini catalyst,because the Catalyst Bed Pressure Drops in radial reactors tend to be lower thanin the corresponding axial reactors. The use of mini catalyst in axial reactors isclearly out of the question because of the unacceptably high Catalyst BedPressure Drops that would result.

Desulfiding of the catalyst is not expected to occur. However, margins of only 16(Case A2 EOR) and 5 (A2 SOR)

oC, respectively, were calculated, whereas we

would normally recommend a margin of 25o

C. If the H2S Inlet Level drops therewill certainly be a desulfiding problem. It is borderline.

A Catalyst Bed Height of 5.52 m is acceptable.

Linear velocities above 0.35 m/s have never been applied.







For Cases A2 EOR and A2 SOR, respectively, linear velocities of 1.04 and 1.05m/s, respectively, were found. These linear velocities are 1.35 (Case A2 EOR)and 2.19 (Case A2 SOR) times the fluidization threshold.

In down flow mode the linear velocity must be below 1.5 times the fluidizationthreshold to prevent milling of catalyst at the top of the bed. The only way toreduce the linear velocity is by choosing a larger Catalyst Bed Diameter.

At an Inlet Operating Pressure of 40 bara GBHE would have recommended aWet Space Velocity of 8132 h

-1. In the above 2 Cases a Wet Space Velocity of

10539 h-1

was found. Again this is pretty borderline (well worse, really). On theother hand GBHE may have been a bit too cautious here. Again the only way toimprove this situation is by choosing a larger Catalyst Bed Diameter.

vulcan sgcu coal to liquids case study

Documents