GAS TO LIQUIDS (GTL) FROM METHANE FUELED SYNTHESIS GAS.pdf

GAS TO LIQUIDS (GTL) FROM METHANE FUELED SYNTHESIS GAS VIA THE FISCHER-TROPSCH
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Required: “CONFIDENTIAL” Sh.a. DESIGNED: Arberor MITA Design group“MITA T.M.D.S.G”
TECHNOLOGY REPORT
PLANT DESIGN AND
PROPOSAL
PLANT "GAS TO LIQUIDS (GTL) FROM
METHANE FUELED SYNTHESIS GAS VIA THE
FISCHER-TROPSCH REACTION"
Summary
The proposed Fischer-Tropsche Methane Gas to Liquids Plant has been evaluated for
feasibility of design, environmental impact, safety concerns, and economic viability. This
includes the design of a grass roots Fischer-Tropsch Reactor sequence and a grass roots

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Summary
The proposed Fischer-Tropsche Methane Gas to Liquids Plant has been evaluated for
feasibility of design, environmental impact, safety concerns, and economic viability. This
includes the design of a grass roots Fischer-Tropsch Reactor sequence and a grass roots
product separation process. The plant outlined in this report has low environmental impact
and a high carbon efficiency converting 81.6% of the carbon in the feed to products. The
proposed plant has been designed to maximize efficiency and profit and to minimize
environmental impact and safety hazards.
The Syngas and Hydro-Isomerization Unit have already previously been designed, and the
Fischer-Tropsche reactor unit and product separations have been designed in this report to
integrate with these existing processes through heat integration and recycle utilization. This
design package features a waste heat boiler unit to recover heat generated from the syngas
unit. In addition, a CO2 recovery and compression unit has been designed for CO2
separations from syngas. The recovered CO2 will be recycled and the remainder will be piped
to the methane reservoirs for storage. The reactor is designed to maximize carbon monoxide
conversion and diesel selectivity. The proposed reactor design is capable of 95+% conversion
of carbon monoxide to alkanes. The reactor unit design includes a cooling and energy
recovery system to recover the heat generated by the reaction. Furthermore, the product
separation unit is designed to minimize the amount of unnecessary product feed to the HI
unit. All of the unit designs proposed have been heat integrated to reduce steam and cooling
water utilities. To maximize carbon efficiency and profit, the GTL plant must incorporate a tail
gas recycle from the distillation and HI units back to the syngas unit. Without the
incorporation of the tail gas recycle the plant would not be economically viable. The GTL plant
has been designed to utilize self produce utilities to supplement a large portion of the utilities
cost to operate the plant. The plant also utilizes a self produced tail gas to supplement for the
natural gas feed, allowing for a higher obtainable yield. Realistic and adequate process
controls have been designed to maximize process consistency for the Fischer Tropsche unit in
the attached plant proposal.
A full economical analysis was conducted on the GTL plant, and was determined to be highly
cost effective. The total capital investment for the GTL was determined to be $356M. Through
economic analysis the Net Present Value of the proposed plant designed is determined to be
$240M. The DCFROR was determined to be approximately 18%. Furthermore, the payback
period for this design is 3.35 years with a 15 year plant life. The annual net profit is expected
to be $146M. Also, a Monte-Carlo Simulation was generated, which yields a 57.5% probability
that the plant will make money based on a 10% interest rate. The economic and
environmental analysis yield promising results for the success of this plant; however, the plant
success is dependent highly upon the fluctuation in crude oil market price. Although this
preliminary design considered many environmental, safety, and health concerns posed by the
GTL plant, these matters should be considered further prior to implementation of the final
design. Our design firm finds the proposed Gas to Liquids plant design to be a profitable
venture with low environmental impact, and further design and evaluations efforts should be
made.
Table of Contents
1. List of Tables Figures and Equations.................................................iv
2. Project Premesis.............................................................................vii

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Table of Contents
1. List of Tables Figures and Equations.................................................iv
2. Project Premesis.............................................................................vii
3. Introduction....................................................................................1
4. Summary.........................................................................................3
5. Discussion.......................................................................................7
5.1 Unit 100: Synthesis Gas..................................................................7
5.2 Unit 200: Waste Heat Recovery.......................................................9
5.2.1 Boiler Design...............................................................................9
5.2.2 Turbine Design..........................................................................13
5.3 Unit 300: FTR Feed Purification.....................................................14
5.3.1 Water Knockout and Compression..............................................14
5.3.2 CO2 Absorption.........................................................................14
5.3.3 Solvent Regeneration.................................................................15
5.3.4 CO2 Solubility and Selexol Flow Requirement...............................16
5.3.5 McCabe-Thiele Analysis..............................................................16
5.3.6 Abosrber Diameter Calculation....................................................19
5.4 Unit 400: Fischer-Tropsche Synthesis Reactors...............................21
5.4.1 Pressure Drop...........................................................................23
5.4.2 Heat Transfer............................................................................27
5.4.3 Catalyst Weight.........................................................................30
5.4.4 Recycle.....................................................................................34
5.4.5 Design Summary.......................................................................35
5.4.6 Unit 400 Design.........................................................................35
5.5 Unit 500: FTR Product Separation..................................................38
5.5.1 Liquid Product Separation...........................................................39
5.5.2 Gaseos Product Sepeartation......................................................40
5.5.3 Unit Outputs..............................................................................42
5.5.3.1 Main Products.........................................................................42
5.5.3.2 Side Products.........................................................................44
5.5.4 Product Storage.........................................................................45
5.5.5 Seperations Equipment Design...................................................45
6. Process Controls.............................................................................51
6.1 Unit 200......................................................................................51
6.2 Unit 300......................................................................................52
6.3 Unit 400......................................................................................53
6.4 Unit 500......................................................................................55
7. Environmental Summary.................................................................66
7.1 Waste Treatment..........................................................................66
7.2 Carbon Capture and Storage ........................................................67

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6.1 Unit 200......................................................................................51
6.2 Unit 300......................................................................................52
6.3 Unit 400......................................................................................53
6.4 Unit 500......................................................................................55
7. Environmental Summary.................................................................66
7.1 Waste Treatment..........................................................................66
7.2 Carbon Capture and Storage ........................................................67
7.3 Environmental Efficiency...............................................................73
8. Safety Summary.............................................................................74
8.1 Standard Safety...........................................................................74
8.2 Unit 400 Safety............................................................................75
8.3 Process Control Safety..................................................................76
8.4 Explosion Proof Equipment............................................................76
8.5 Flammability Limits.......................................................................77
8.6 Ventilation and Deluge System......................................................80
9. Equipment Sizing...........................................................................82
9.1 Pump Sizing.................................................................................82
9.2 Compressor Sizing........................................................................82
10. Economics...................................................................................84
10.1 Equipment Cost..........................................................................84
10.2 GTL Plant Utilities.......................................................................86
10.3 Cash Flow Analysis.....................................................................88
10.4 Monte Carlo Simulation...............................................................89
11. Conclusions and Recommendations...............................................92
11.1 Unit 100.....................................................................................92
11.2 Unit 200.....................................................................................92
11.3 Unit 300.....................................................................................93
11.4 Unit 400.....................................................................................93
11.5 Unit 500.....................................................................................94
11.6 Overall.......................................................................................95
11.7 Final Conclusion.........................................................................95
12. References...................................................................................96
Appendix A: Procces Flow Diagrams with Stream Summaries......I to XXVII
Appendix B: Equipment Summary Tables.......................................I to IX
Appendix C: Economic Summary and Outputs..............................I to XIV
Appendix D: Calculations & Computer Simulation Outputs.........I to XXVIII
Project Premises
1. Standard conditions are at 15 °C and 1 atm.
2. At standard conditions for this project, = 1 kg mol.
23.691

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Project Premises
1. Standard conditions are at 15 °C and 1 atm.
2. At standard conditions for this project, = 1 kg mol.
23.691
3. All C4 molecules produced are straight chain butane for simulation purposes.
4. Shell-side of the reactor operates isothermally.
5. Pressure drop per tray in distillation columns is assumed to be 690 Pa (Brannan, 2005).
6. ChemCAD provide accurate physical property data for processes streams and
components.
7. Assume all tray column stage efficiencies are 85%. (Walas, 1990).
8. Liquid turbines act as reverse pumps; therefore, they can be sized as pumps with a
negative value for pressure drop.
9. Assume Tk-601 is heated by steam coil that is incorporated into the direct and indirect
costs of the tank taken into account by the multiplying factor of 4.8 in the total capital
costs. Also, assume that the cost of the steam utility is accounted for in the steam utility
usage of the HI unit already preset, as the actual steam consumption of the tank heater
will be negligible. The tank heater is assumed to have been designed to maintain the tank
at a minimum temperature of 138 deg C to prevent solidification of the Distillate/Wax.
10. Simulated flash in ChemCAD of a stream containing the FTR reaction products at the
outlet temperature and pressure by the Anderson-Shulz-Flory distribution properly reflects
the phase separation of reactor products leaving the reactor.
11. We will assume that the cooling water tower and associated pumps for supplying the
process with water exist as part of the company’s accessible utility grid and will not be
costed or sized as part of the scope of this project. Any additional operating costs
associated with the load placed on this equipment by our plant is assumed to be
incorporated into the material purchase price.
12. Utilities for compressors and turbines were calculated on a turbine/compressor basis,
not on the drive itself to better keep track of individual equipment production/
consumption. Drive efficiencies as well as compressor/turbine efficiencies were taken into
account during the utility calculation and the sizing of the equipment
13. For heat exchangers: (1) assume shellside temperature varies linearly with length; (2)
Physical properties remain constant; (3) assume 2.5% of the heat is dissipated to the
surroundings; (4) All heat exchanger with now phase change will be fixed head shell and
tube exchanger, where as heat exchangers with phase change will be floating heat shell
and tube exchangers.
14. Overall heat transfer coefficients from Table 3 and 4 in Brannan’s Rules of Thumb for
Chemical Engineers were assumed to be accurate estimations.
15. The calculations done by hand and the outputs of ChemCAD are proven to be within
the acceptable 95% confidence limit.
16. Shipping and logistics costs are incorporated in the additional operating costs taken as
3% of the total capital investment.
17. For all pumps, a pipe length of 30 m. was estimated for pressure drop
accommodation.
18. Efficiencies of large centrifugal compressors, 45000-75000 m3/hr at suction, are
76-78% (Turton, 2007).
19. For large machines, electric drive efficiency is greater 85-95% (Turton, 2007).
20. Electric drives are available for services up to 15000 kW (Turton, 2007).
21. Steady-state design may not require all utilities. Transient start-up requirements for
preliminary design may be ignored.
22. Gas expanders for power recovery may be justified at capacities of several hundred

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17. For all pumps, a pipe length of 30 m. was estimated for pressure drop
accommodation.
18. Efficiencies of large centrifugal compressors, 45000-75000 m3/hr at suction, are
76-78% (Turton, 2007).
19. For large machines, electric drive efficiency is greater 85-95% (Turton, 2007).
20. Electric drives are available for services up to 15000 kW (Turton, 2007).
21. Steady-state design may not require all utilities. Transient start-up requirements for
preliminary design may be ignored.
22. Gas expanders for power recovery may be justified at capacities of several hundred
horsepower; otherwise any pressure reduction in the process is done with throttling
valves (Turton, 2007).
23. A typical construction period of 2 years has been assumed (Turton, 2007) with the
capital being divided 60/40 between the first two years, respectively.
24. It is assumed that during year one, only 85% of the total expected revenue from
sales can be accounted for due to “teething” problems after initial startup of the facility
(Turton, 2007).
25. It is assumed that the catalyst will be changed four times over the life of the plant in
addition to the initial cost of the catalyst. These costs have been turned into an annuity to
be incurred each year as a raw material for the purposes of the economic analysis. The
catalyst changes are assumed to occur at year 0 (startup), 4, 8, and 12.
26. It is assumed that the salvage value of the plant is near zero, and is relatively
negligible in the scope of the project economics.
27. The working capital is assumed to be 15% of the total capital investment (4.8Cp0).
28. An average discount rate of 10% is used for the discounted cash flow analysis
(Robertson, 1999).
29. It is assumed that the cost of the land, waste heat recovery, CO2 recovery and
compression, the storage of materials, and transport of material to and from the SU and
HI units is incorporated into the total $500M annual capital recovery costs of the HI and
SU units.
30. Assume COMd during turnaround years does not change. Even though turndown may
reduce some manufacturing costs, many expenses remain fixed and other turnaround
expenses may be incurred. It is assumed that these differences offset and COMd remains
constant for the life of the project.
31. Revenue from sales during turnaround years is assumed to be the same as other
years because of accumulated inventory in our existing storage tanks. The additional
capacity designed into the existing equipment allows for increased throughput or recycle
rate to counteract any losses incurred during the turnaround period.
32. It is assumed that there is no degradation of catalyst activity between change outs.
Therefore, product distribution from the FTR reactors is relatively constant.
33. It is assumed, based on the reaction kinetics provided, that no CO2 is produced as a
product of the Fischer-Tropsche synthesis reaction due to the highly efficient cobalt
catalyst used in the plant.
34. It is assumed, for distillation column design purposes, that C21+ pseudocomponents
were as volatile as the most volatile component in the array (i.e., the vapor pressure at a
certain temperature and pressure of C21-C25 material is equal to the vapor pressure of
straight chain C25 at that temperature and pressure).
35. The following assumptions were made in order to apply the McCabe-Thiele analysis
method to the design of CO2 absorbers utilizing Selexol solvent: a. The heat of absorption
is considered to be negligible.
b. Temperature changes in the column are negligible.
c. The solvent has a very low vapor pressure.
d. The gas containing the solute, CO2, is insoluble in the solvent.
e. The mass flow rates of the carrier gas and the solvent are constant.

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straight chain C25 at that temperature and pressure).
35. The following assumptions were made in order to apply the McCabe-Thiele analysis
method to the design of CO2 absorbers utilizing Selexol solvent: a. The heat of absorption
is considered to be negligible.
b. Temperature changes in the column are negligible.
c. The solvent has a very low vapor pressure.
d. The gas containing the solute, CO2, is insoluble in the solvent.
e. The mass flow rates of the carrier gas and the solvent are constant.
36. 70% flooding is assumed for CO2 absorber design.
37. Constant molal overflow is assumed for distillation column design.
38. Assume 75% flooding in Unit 500 columns.
39. Pressure drop across heat exchangers was assumed to 0.35 bar for normal fluid flow
and 0.1 bar for boiling applications.
40. It is assumed that any Selexol lost to the process is recovered in water knockouts
downstream.
41. Multiplying the calculated gas phase pressure drop by 0.1 accounts for pressure
losses due to the conversion of gas to liquids.
42. The stoichiometric coefficients of hydrogen and carbon monoxide are 2 and 1,
respectively, based on the average expected Fischer-Tropsche reaction.
43. It was assumed that a 25 mm spacing triangularly packed tube arrangement
correlates with the heat transfer correlation provided in the project package.
44. It was assumed, for pressure drop calculations, that the flow of gases through the
reactor tubes is fully developed turbulent flow.
45. The gas flowing through the reactor was modeled as an ideal gas for pressure drop
calculations.
46. It was assumed that water has minimal fouling affect on the ultra-stable cobalt
catalyst developed by the company.
47. It was assumed that since a methane pipeline exists nearby, a CO2 most likely also
exists within a reasonable distance.
48. It was assumed that a natural methane reservoir is located nearby to the plant
location that is consistent with geological characteristics desired for a CO2 storage site.
49. M = 106 for all unit designation purposes.
50. The K-value and enthalpy models used to simulate the processes were as follows:
a. Pre-FTR unit operations: i. K-value: Ideal Gas
ii. Enthalpy: SRK
b. Local Thermodynamic settings for CO2 absorbers were: i. K-value: PSRK Solvent
Package
ii. Enthalpy: SRK
c. Post FTR unit operations: i. K-value: Peng-Robinson
ii. Enthalpy: Peng-Robinson
3. Introduction
As crude oil is the world’s most predominant source of fuel, and the world’s energy
demand continues to increase, the drive for alternative energy sources has begun to grow

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3. Introduction
As crude oil is the world’s most predominant source of fuel, and the world’s energy
demand continues to increase, the drive for alternative energy sources has begun to grow
rapidly. Therefore, liquid hydrocarbon fuels from synthetic gas have proven to be a viable
fuel production via Fischer-Tropsche synthesis (FT). Fischer-Tropsche synthesis is an
extremely viable process, where natural gas and coal reserves are abundant, and has
been an alternative to crude oil synthesis since the early 20th century. The survival of
Fischer-Tropsche plants is highly dependent on the price and availability of crude oil. High
yields of gasoline, excellent quality diesel fuel, or high value linear α-olefins are controlled
by the downstream process of the FTR .
Industrial operation of the Fischer-Tropsche synthesis involves five steps: (1) synthesis
gas manufacture; (2) gas purification by removal of water and carbons; (3) synthesis of
hydrocarbons; (4) condensation of liquid products and recovery of gasoline from product
gas; and (5) fractionation of synthesis products. A simplified block flow diagram of this
process can be seen below in figure 1.1.
Figure 1 Simplified GTL Block Flow Diagram
Coal, peat, biomass, and/or natural gas are converted into synthesis gas, which is
predominantly composed of CO and H2. This synthesis gas is then converted to a multi-
component mixture of hydrocarbons. Due to the absence of sulfur in the feedstock to the
FT process and low concentrations of aromatic compounds, the fuels that are produced
are extremely high purity and high quality. With the FT synthesis method cetane numbers
can be as high as 75. Waxes, alcohols, solvents, ketones, α-olefins, propene, and ethane
are some of the many valuable products that can be tailored by the FT process in
combination with upgrade process.
In the FT synthesis reaction, hydrocarbons and water are produced by the net reaction
shown below in Figure 2. This reaction occurs at highly controlled temperatures, which are
selective towards specific hydrocarbons.
H2 + CO = H2O + (CH2)nH2
Figure 2: Overall Chemical Reaction
The reaction above is accompanied by and ultra-stable Cobalt-base FTR catalyst. The rate
of the reaction follows the Langmuir-Hinshelwood equation and parameters. The
selectivity towards hydrocarbons follows the Anderson-Shulz-Flory (ASF). These
parameters and mechanism will be discussed in there reactor design section of this report.
After the synthesis gas has been processed in the Fischer-Tropsche Reactor train, the wax

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The reaction above is accompanied by and ultra-stable Cobalt-base FTR catalyst. The rate
of the reaction follows the Langmuir-Hinshelwood equation and parameters. The
selectivity towards hydrocarbons follows the Anderson-Shulz-Flory (ASF). These
parameters and mechanism will be discussed in there reactor design section of this report.
After the synthesis gas has been processed in the Fischer-Tropsche Reactor train, the wax
(C 11+) is hydro-cracked and hydro-isomerized to yield a middle distillate boiling range
product. The hydro-isomerization stage employs a commercial hydro-cracking catalyst in a
trickle flow reactor where 100% of the wax distillate. The products from the Hydro-
Isomerization unit include methane, ethane, liquefied petroleum gas (LPG), naphtha, and
diesel.
As crude oil supplies continue to diminish, Large International companies as well as
several private investors are beginning to build larger, environmentally efficient facilities to
produces fuels to blend with fuels produced from crude fuels and decrease crude oil usage
exponentially. As economies come out of the economic recession, “Natural gas is
suspected to increase. Natural gas demand is only suspected to increase as fossil fuel
demands increase and crude oil synthesis does not have the capacity to supply.
A proposal has been made to build a Fischer-Tropsche unit (FTR) as part of a planned
GTL plant. The design also includes all post separations facilities, as well as a steam and
power production facility. The design is focused on safe, environmentally efficient, and
thermally integrated process with optimal and adequate capital and operating cost
utilization. The FTR unit easily integrates optimally and economically into the existing
design of the Syngas and Hydro-Isomerization units, which have already been designed.
An overview can be found in the summary section of this report followed by a detailed
discussion of each unit operation design.
4. Summary
The objective of this proposal is to outline the preliminary grass roots design package of a
Fischer-Tropsche Reaction unit (FTR) and associated Fischer-Tropsche synthesis products
separation facility as part of a planned Gas to Liquids (GTL) production facility by J&D
Engineering. Overall objectives of the design require that the facility be planned with
specific focus on safety, environmental impact, and thermal efficiency through heat
integration. Additionally, the designed units must be completely integrated with the
already designed Synthesis Gas Unit and Hydro-Isomerization Unit and other designed or
existing facilities required for efficient diesel, naphtha, and liquefied petroleum gas (LPG)
production. The specific objectives of the preliminary design package are:
• To utilize appropriate equipment, instrumentation, and process control strategies to
eliminate any environmental, health, or safety hazards within the plant and to the
surrounding area
• To implement an effective recycle strategy to minimize waste and prevent unnecessary
venting or flaring of material
• To determine optimum operating conditions to maximize the conversion of the mass of
carbon in the feed into finished fuel products
• To perform the appropriate economic evaluation to provide evidence of the development
of an optimum liquid fuel production system
• To implement a realistic and sufficient process instrumentation and control strategy to
improve process safety and efficiency and minimize plant personnel requirements.
By going above and beyond the overall and specific objectives and proposal requirements
listed above, Mita Technological and Mechanical Design Studies Group has ensured a Gas

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• To determine optimum operating conditions to maximize the conversion of the mass of
carbon in the feed into finished fuel products
• To perform the appropriate economic evaluation to provide evidence of the development
of an optimum liquid fuel production system
• To implement a realistic and sufficient process instrumentation and control strategy to
improve process safety and efficiency and minimize plant personnel requirements.
By going above and beyond the overall and specific objectives and proposal requirements
listed above, Mita Technological and Mechanical Design Studies Group has ensured a Gas
to Liquids facility design that meets all company standards for safety, environmental
impact, and profitability.
The GTL plant will be comprised of a series of units that have each been designed with a
specific purpose. The unit numbers and descriptions can be seen in Table 1 below.
Unit Number Description
100 Synthesis Gas Unit
200 Waste Heat Recovery Unit
300 FTR Feed Purification Unit
400 Fischer-Tropsche Reaction Unit
500 FTR Product Separation Unit
600 Hydro-Isomerization Unit
700 Air Separations Plant (Outside Contract)
The unit identification numbers and descriptions may be used interchangeably throughout
this discussion. Please refer back to this table for clarification of the unit number if
necessary.
The GTL plant will utilize a 588 m3/hr feed of clean methane from a newly located,
remote source, which will be combined with high pressure steam (HPS), carbon dioxide,
purified oxygen, and recycled tail gas to be converted into synthesis gas in an autothermal
reformer. The feed preparation equipment and the autothermal reformer will make up
Unit 100, the Synthesis Gas unit. The purified oxygen will be purchased through a
prearranged contractual agreement with a nearby Air Separations Plant. The steam and
CO2 feeds will be comprised entirely of material recycled from downstream processes to
reduce the cost of raw materials and minimize waste output. The autothermal reformer
converts the methane into a mixture of hydrogen gas and carbon monoxide in a 2:1 molar
ratio, along with a significant amount of water and carbon dioxide.
The effluent leaving the autothermal reformer will be at a very high temperature and will
first been charged to a heat recovery system prior to being purified and sent to the FTR
Unit.
The syngas from Unit 100 will pass through a waste heat boiler and a series of steam
generators in order to recover the heat produced in the exothermic partial oxidation of
methane occurring in the autothermal reformer upstream. This equipment will make up
Unit 200. High pressure saturated steam produced in this unit will be recycled to the feed
of the Synthesis Gas unit and will be used as the final preheat media for the FTR feed.
Other high and low pressure steam produced in this unit will be returned to the company’s
utility grid for credit. The waste heat boiler will produce superheated high pressure steam
that will be utilized in a series of steam turbines to generate electricity. The low pressure
saturated steam exiting the last turbine is also returned to the utility grid for credit. The
hot synthesis gas will also be crossed with several downstream process streams prior to
being cooled the final degree with cooling water. The steam generators, turbines,
exchangers, and associated boiler feed water equipment are all considered part of Unit
200, the Waste Heat Recovery unit.
The cooled synthesis gas leaving Unit 200 contains a large amount of water and CO2,
which will be removed prior to the FTR unit to reduce reactor size requirements and
increase the concentration of reactants in the unit. The bulk of the water will be removed

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that will be utilized in a series of steam turbines to generate electricity. The low pressure
saturated steam exiting the last turbine is also returned to the utility grid for credit. The
hot synthesis gas will also be crossed with several downstream process streams prior to
being cooled the final degree with cooling water. The steam generators, turbines,
exchangers, and associated boiler feed water equipment are all considered part of Unit
200, the Waste Heat Recovery unit.
The cooled synthesis gas leaving Unit 200 contains a large amount of water and CO2,
which will be removed prior to the FTR unit to reduce reactor size requirements and
increase the concentration of reactants in the unit. The bulk of the water will be removed
in a preliminary flash drum, after which the syngas will be compressed to 69 bar. The
syngas is the sent through two parallel CO2 absorbers, which utilize Selexol solvent to
effectively removes more than 99% by weight of the CO2 from the syngas stream. The
solvent will be regenerated in a three stage flash sequence, after which it is cooled and
recycled to the tower. The bulk of the CO2 removed is recycled back and utilized as a feed
to the syngas reactor. The remainder is pressurized and cooled in preparation for injection
into a CO2 pipeline. The CO2 will then be transported to nearby methane reservoirs for
underground storage or to be utilized by the methane well as pressure sweeping fluid. The
CO2 produced by the GTL plant can be utilized by the methane wells to help maintain and
maximize the wells.
The reactor unit is designed to process over 75,000 kmol/hr of material at 98%
conversion of carbon monoxide. This is accomplished through 8 identical reactors
arranged in 4 trains with 2 reactors in series per train. Each reactor will be 6 m in
diameter and 18 m in length. The reactors will house almost 16,716 40 mm tubes. The
reaction will be conducted at 230 deg C and 30 atm. Cooling water will be injected into
the reactor shell and converted to saturated steam at 220 deg C by the heat generated by
the reaction. The reactor controls are designed to maintain isothermal operation in the
reactor to obtain a consistently high conversion. The primary safety concerns in the
reactor unit are a runaway reaction and fire. A secondary cooling system should be
installed along with a deluge and monitor system to prevent fires and runaway reactions.
The diverse hydrocarbon product leaving the FTR unit will be divided between a gas phase
and a liquid phase. The liquid phase will be subjected to a series of depressurization,
flashing, and substantial preheating before being distilled to separate diesel and heavier
material from naphtha. The naphtha stream overhead from this column is then separated
from LPG and tail gas in the next distillation column. All off gases from the liquid
separation sequence are recycled and compressed in the gas separation sequence. The
gas product from the FTR Unit is cooled and flashed to remove entrained heavy
hydrocarbons before undergoing a refrigerated flash to remove the bulk of the methane in
the gas phase. The light gas stream is then charged to the first column in the gas
separation sequence, where ethane and lighter material is separated from LPG and any
entrained naphtha. The next distillation column will achieve the final butane/pentane split
to yield a pure LPG and light naphtha product. The light naphtha product from this tower
can be combined with the naphtha product from the liquid phase separation as a single
product. The diesel and heavier wax material recovered in the first phase of the liquid
separation are used as the feed stream to Unit 600, the Hydro-Isomerization Unit. In this
unit, the material heavier than diesel will be converted to tail gas, LPG, naphtha, and
diesel, which can then be sold as a finished product or recycled for use elsewhere in the
plant.
The control loops incorporated in the GTL plant include simple feedback, cascade, and
ratio controllers. The heat exchangers which utilize cooling water or steam as a heat
exchange media are cascade controlled to incorporate some feed forward control into the
loop to reduce lag time. The two heaters in the process are also cascade controlled to
reduce lag time. Branch points and feed junctions are ratio controlled to maintain a certain
flow ratio to insure consistent
process conditions. The tower separations are controlled by feedback temperature control.
The majority of the heat exchangers in the process have process streams on both shell
and tube sides. The outlet temperatures are controlled by manipulating the flow of a

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plant.
The control loops incorporated in the GTL plant include simple feedback, cascade, and
ratio controllers. The heat exchangers which utilize cooling water or steam as a heat
exchange media are cascade controlled to incorporate some feed forward control into the
loop to reduce lag time. The two heaters in the process are also cascade controlled to
reduce lag time. Branch points and feed junctions are ratio controlled to maintain a certain
flow ratio to insure consistent
process conditions. The tower separations are controlled by feedback temperature control.
The majority of the heat exchangers in the process have process streams on both shell
and tube sides. The outlet temperatures are controlled by manipulating the flow of a
stream that bypasses the equipment. All control valves have been designed to fail safe to
minimize the consequences of failure. The primary goal of the control design is to control
the process in the simplest way possible which minimizes lag.
The primary safety concerns in the GTL plant are fires and explosions. Several layers of
safety must be put into place ranging from the design of the process to the plant
emergency response systems encompassing safe control systems, critical alarms and
human intervention, safety instrumented functions, physical protection, and post-release
physical protection. All equipment used in the process must be appropriately classified
explosion proof equipment. Analyzers must be implemented to detect flammable vapor
concentrations approaching the lower flammability limits. The GTL plant will be an open
air plant which provides adequate natural ventilation to avoid the concentration of
flammable vapors. Proceeding GTL plant designs should compound on the safe design
proposed in this preliminary design package.
A complete economic analysis of the plant was performed, which included the costing of
all equipment designed in Units 400 and 500 and determination of the annual cost of
manufacturing and annual revenue from finished product sales. A cash flow analysis was
performed to determine the net present value of the facility at the most likely conditions
of operation. A Monte-Carlo analysis was also completed to quantify the risk associated
with investment in the GTL facility based on certain variable economic parameters. The
Monte-Carlo simulation yielded a net present value (NPV) cumulative probability curve that
showed that the investment has a 57.5% probability of returning a positive NPV at the
end of the plant life. Other economic parameters, such as discounted payback period
(DPBP), present value ratio (PVR), and total capital investment, are shown Table 2 below.
Net Present Value $240M
Discounted Payback Period 3.35 years
Present Value Ratio 1.69
Total Capital Investment $356M
Annual Revenue $1573M
Annual Cost of
Manufacturing
$1427M
Table 2: Economic Analysis Summary
This analysis provides evidence in support of the continuation of this project to the next
design phase and ultimately the GTL facility’s construction. The following discussion
provides an in depth look at the preliminary design of the plant and makes
recommendations regarding the next steps in the design process and post construction
plant operation.
5. Discussion
5.1 Unit 100: Synthesis Gas
The first unit in the Gas to Liquids (GTL) plant is the Synthesis Gas Unit, which uses a

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recommendations regarding the next steps in the design process and post construction
plant operation.
5. Discussion
5.1 Unit 100: Synthesis Gas
The first unit in the Gas to Liquids (GTL) plant is the Synthesis Gas Unit, which uses a
clean methane feed of 588 m3/hr to produce syngas, which consists of predominately
carbon monoxide and hydrogen gas, in an autothermal reformer. The methane will be fed
into the Syngas unit along with high pressure steam and carbon dioxide (CO2), where it
will be preheated in a fired heater, H-101, before being mixed with a purified oxygen
stream and charged into the autothermal reformer. The autothermal reformer utilizes
three parallel reactions to complete the production of syngas. These reactions are the
partial oxidation of methane, the steam reformation of methane, and the water-gas shift
reaction, which are represented, respectively, as:
The partial oxidation of methane runs to completion in the reformer; the heat from this
extremely exothermic reaction will provide the energy necessary to run the other two
endothermic reactions. Oxygen will be fed at a sub-stoichiometric molar ratio to prevent
the complete combustion of the methane. The oxygen supplied for this reaction is 99%
pure and will be purchased from a nearby Air Separations Plant under predetermined
contractual terms.
Steam is fed to the system to fuel the reformer to fuel the reformation reaction, where
methane is converted into the desired products, carbon monoxide (CO) and hydrogen gas
(H2). This reaction runs to equilibrium in the autothermal reformer, not to completely like
the partial oxidation reaction. The steam rate must be at a minimum of 0.5 mol/mol of
methane fed to prevent coking of the feed. However, the 3:1 molar ratio of H2 to CO is
not the stoichiometric ratio at which the reactants are consumed in the downstream
Fischer-Tropsche reactors. For this reason, CO2 is added to the feed stream in the
appropriate molar ratios to drive the water-gas shift reaction back toward products. This
equilibrium shift results in an increase in CO production and a decrease in H2 production
until the desired molar ratio for the FTR Unit, 2 mole H2 per mole CO, is achieved.
Although the design of the Synthesis Gas Unit is not within the scope of the project, the
efficient operation of the unit to obtain the desired FTR unit feed and the successful
integration of the unit with the rest of the GTL facility were goals of the complete design
package. The operating temperature and pressure of the autothermal reformer were
specified as an operable range of 870-1065°C and 20-35 bar, respectively. To determine
the best conditions at which to operate the unit, including molar ratios of steam, oxygen,
and carbon dioxide to methane, the Synthesis Gas Unit was modeled in ChemCAD. The
autothermal reformer was modeled using a Gibbs reactor, which utilizes the individual
component properties in the ChemCAD library to determine the most stable form of the
products based on the reduction of Gibbs free energy. The feed heater H-101 and gas
expander J-101 were placed before the reactor in order to control the reactor temperature
and pressure. Please refer to the process flow diagram of the unit in for a detailed
representation of Unit 100.
The model of Unit 100 was then used to perform a sensitivity analysis to determine the
operating temperature and pressure at which the utility costs of the unit would be
minimized. The feed ratios may then be adjusted to maximize production of the FTR
synthesis reactants. The results of the sensitivity study are present in this material. The
final conclusion regarding the operating temperature and pressure was to minimize both
to reduce the overall utility consumption of the unit.
The final molar ratios of CO2, steam, and O2 to the methane being fed to the reactor were
determined after the decision was made to recycle 118 m3/hr of tail gas from the FTR

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representation of Unit 100.
The model of Unit 100 was then used to perform a sensitivity analysis to determine the
operating temperature and pressure at which the utility costs of the unit would be
minimized. The feed ratios may then be adjusted to maximize production of the FTR
synthesis reactants. The results of the sensitivity study are present in this material. The
final conclusion regarding the operating temperature and pressure was to minimize both
to reduce the overall utility consumption of the unit.
The final molar ratios of CO2, steam, and O2 to the methane being fed to the reactor were
determined after the decision was made to recycle 118 m3/hr of tail gas from the FTR
Product Separation Unit into the Synthesis Gas Unit feed stream. This recycle stream
contains approximately 85% methane by weight, with the balance being made up of CO,
H2, and light hydrocarbon gases. The recycle of this material increases throughput in the
unit and also increases the overall conversion of the methane feed into saleable products.
The steam and the supplemental CO2 fed into the unit are also being recycled from the
Waste Heat Recovery Unit and the CO2 Absorption and Compression Units, respectively.
These recycles will prevent the need to purchase high pressure steam and CO2 as feed
materials during steady state operation. The molar ratios chosen to maximize CO and H2
production in the autothermal reformer were determined via a trial and error analysis in
ChemCAD. The feed ratios relative to methane and molar flow rates are shown below in
Table 3.
Material Flow Ratio (Relative to
CH4)
Molar Flow [kgmol/hr]
CH4 (Including Recycle) 1.000 28900
High Pressure Steam (495 °F,
600#)
0.750 21675
O2 (99 mol % pure) 0.549 77262
CO2 (99.5 mol % pure) 0.266 770
The autothermal reformer effluent leaving the synthesis gas unit contains the appropriate
2:1 ratio of hydrogen gas to carbon monoxide, at about 43 mol % H2 and 21.5 mol % CO.
The remainder of the stream is comprised of approximately 23 mol % H2O, 10 mol %
CO2, and the balance being mostly methane with trace amounts of other light gases. For
more detailed stream composition and property data, please refer to the Stream Summary
Tables in Appendix A. The reformer effluent will exit the reactor at approximately 895 °C
and be sent to Unit 200 for recovery of the waste heat and Unit 300 for purification prior
to being fed to the Fischer-Tropsche synthesis reactors in Unit 400.
5.2 Unit 200: Waste Heat Recovery
Coupled with the FTR and post-separations design, is the design of a waste heat boiler
system to recover the heat from the synthesis gas before being fed to the Fischer-
Tropsche Reactors. This will maximize the over thermal efficiency of the plant and provide
extensive heat integration. The boiler will produce superheated steam, which will be
turned into electrical energy through multistage turbine generators. The steam and power
system designed for this plant can be seen in Appendix A.
5.2.1 Boiler Design:
Synthesis gas from the Syngas unit (SU) will be initially fed into a 1-1 tube pass, floating
head, kettle re-boiler (E-301 ) where heat will be exchanged with boiler feed water at a
rate of 70000 kW. The heat duty for this exchanger as well as all other exchanger
iscalculated from equation 4 below.
Here the duty (dQ) is calculated with average specific heats. The specific heats (Cp) were
obtained by multiplying the compound mole fractions by the compounds specific heat
value. For several cases the Cp equation must be left inside the integral because specific

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5.2.1 Boiler Design:
Synthesis gas from the Syngas unit (SU) will be initially fed into a 1-1 tube pass, floating
head, kettle re-boiler (E-301 ) where heat will be exchanged with boiler feed water at a
rate of 70000 kW. The heat duty for this exchanger as well as all other exchanger
iscalculated from equation 4 below.
Here the duty (dQ) is calculated with average specific heats. The specific heats (Cp) were
obtained by multiplying the compound mole fractions by the compounds specific heat
value. For several cases the Cp equation must be left inside the integral because specific
heats vary greatly for gases. However, liquid Cp’s can be assumed to be constant, which
allows for a 95% confidence to be obtained. Furthermore, the area to achieve the
required heat transfer is calculated with the log mean temperature difference method. The
equation for this calculation can be seen below.
where
U = overall heat transfer coefficient
Q = Heat Duty, kW
A = Area to achieve required heat transfer
ΔTLM= log mean temperature difference
＝
Q ⋅
⋅
U A ΔTLM
The process for sizing heat exchangers is intensive and specific to the application. An
example of this calculation was done by hand, however, due to the large number of
exchangers the remaining exchanger duties as well as areas (A). All assumptions made for
these sizing calculations can be found in the premises section of this report.
High pressure saturated stream (HPS) is extracted off exchanger E-201 to be used as a
feed to the SU unit. The only demand for HPS is for the feed into the syngas unit, which is
produced by E-201.The synthesis gas enters this exchanger at 880 oC and 22 bar, and is
cooled to 660 oC. This requires an approximate heat transfer area of 2500 m2.
From there, synthesis gas passes through a waste heat boiler where superheated steam is
produced. Superheated steam is produced by heating saturated steam at constant
pressure, so the temperature rises above the saturation temperature. Both equations
listed above are used for the boiler sizing. Superheated and Saturated steam tables are
used from Abbot (2005). Below in Figure 3 is the heat curve for boiler B-201.
Figure 3. Waste Heat Boiler Heat Curve.
The heat profile for heat exchangers is important to generate and understand to ensure
there is no temperature cross. If there is a temperature cross, multiple passes are
required to achieve the required heat transfer. The red plot on the temperature profile
represents the profile for the heating of BFW. Initially, the slope rapidly increases then
levels out as the phase change occurs. The length of the flat horizontal line represents

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The heat profile for heat exchangers is important to generate and understand to ensure
there is no temperature cross. If there is a temperature cross, multiple passes are
required to achieve the required heat transfer. The red plot on the temperature profile
represents the profile for the heating of BFW. Initially, the slope rapidly increases then
levels out as the phase change occurs. The length of the flat horizontal line represents
how much energy is required for the BFW to change phases into steam. Then the slope
constantly increases until the steam is heated to the required temperature. This verifies
that one tube pass is sufficient for this design
Most boilers contain a steam and mud drum along with a fire side and water side of the
boiler. However, the main difference in this boiler design is the steam and mud drum have
been removed from the boiler and tubes are passed through the mud drum, in which the
process stream (synthesis gas) will flow. This boiler recovers waste heat and does not
require a combustion chamber. This water tube boiler design was chosen due to the high
thermal efficiencies of water tube boilers. A diagram of a simple waste heat boiler is
shown below in Figure 4.
Figure 4 Waste Heat Boiler
Sizing of waste heat boilers is quite an involved procedure; however for a quick estimate
the method described for waste heat boilers in Branan (2005). The waste heat recovery
boiler (B-201) is designed to process 1816.000 kg/hr, this includes room for a 10%
increase in throughput and steam production. For the preliminary design, the process
stream of synthesis gas will enter the concentric 1,000 50 mm OD (46 mm ID) high

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Sizing of waste heat boilers is quite an involved procedure; however for a quick estimate
the method described for waste heat boilers in Branan (2005). The waste heat recovery
boiler (B-201) is designed to process 1816.000 kg/hr, this includes room for a 10%
increase in throughput and steam production. For the preliminary design, the process
stream of synthesis gas will enter the concentric 1,000 50 mm OD (46 mm ID) high
temperature alloy steel tubes at 660 oC and 21 bar pressure where it will be cooled to 325
oC and 21 bar pressure. The gas will be cooled by 99.9% pure boiler feed water (BFW) at
127.5 oC and 41 bar pressure. The heat cure for the boiler is shown above in Figure 3. For
simulation purposes a 1-1 tube pass, countercurrent, kettle style re-boiler with a floating
head was selected to represent the boiler. The steam and power system was simulated
using ChemCAD software. The outputs for this simulation can be found in Appendix A.
Based on the quick estimation method it was determined that a boiler with a heat transfer
area of 1880 m2 is needed to processes the synthesis gas. Using the formula for surface
area it has been calculated that a boiler with a diameter of 12 m and a height of 23 m will
be sufficient to achieve the desired heat transfer. This boiler will be able to support a heat
duty of 317000 kW. Due to the extremely high temperatures of this steam system,
insulation is necessary to ensure maximum thermal efficiency of boilers, exchangers, and
turbines. There are numerous types of insulation for high temperature applications such
as refractory or gypsum plaster. To determine the best insulation vendor quotes would be
obtained to determine the most economical means to improve the thermal efficiency
(Bergman, 2007).
As heat is transferred from the synthesis gas to the BFW at a heat duty of 440000 kW,
superheated steam is produced at 315 oC and 41 bar. Due to the high temperature and
pressure of the synthesis gas, the materials of construction (MOC) will be a carbon steel
alloy (CS AISI 1020) UNS# G10200 that can withstand temperatures upward of 1515 oC
and 3450 bar(Perry, 2008 ). The boiler will be insulated in addition to the piping and
distribution system. Insulating materials such as calcium silicate, mineral fiber, fiberglass,
perlite, and cellular glass are recommended. Information for these insulating materials is
provided by The American Society for Testing and Materials and the North American
Insulation Manufacturers Association.
The plant is assumed to already have a water purification system designed with carbon
filter beds to remove dissolved solids, heavy metals and silica, as well as resin deionizers
that can meet the throughput demand of the boiler and other steam generating process
equipment. The water is assumed to be treated with softeners and biocides to remove
minerals and organic compounds. The purity of the water will be tested with a total
dissolved solids (TDS) analyzer. The analyzer will be connected to high concentration
alarm. In the case that high dissolved solids enter, the boiler the analyzer will send a
signal to the blow-down flow control to increase the mass flow rate.
As well as total dissolved solids, the presence of oxygen at high temperatures induces
corrosion and should be removed with a unique piece of equipment know as a deaerator.
Here oxygen is stripped from clean recycled condensate with low pressure steam. The
condensate is then sent to the boiler feed water pumps to return clean water back to the
steam drum discussed above. The feed water oxygen concentration can be reduced to
levels ranging from 7 to 40 parts per billion (ppb). To decrease the O2 concentration even
further, deaerating heaters are suggested to be installed. Also, due to changing
downstream steam demands, a clean condensate tank should be installed on the return
line before the deaerator to ensure a steady flow rate. The condensate storage tank
should have a level indicator and transmitter attached as well as a conductivity
transmitter. A rapid change in conductivity indicates a tube rupture in the boiler and if
returned condensate to the boiler will foul and/or plug the down comers.

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condensate is then sent to the boiler feed water pumps to return clean water back to the
steam drum discussed above. The feed water oxygen concentration can be reduced to
levels ranging from 7 to 40 parts per billion (ppb). To decrease the O2 concentration even
further, deaerating heaters are suggested to be installed. Also, due to changing
downstream steam demands, a clean condensate tank should be installed on the return
line before the deaerator to ensure a steady flow rate. The condensate storage tank
should have a level indicator and transmitter attached as well as a conductivity
transmitter. A rapid change in conductivity indicates a tube rupture in the boiler and if
returned condensate to the boiler will foul and/or plug the down comers.
The boiler feed water enters the boiler in the upper drum known as the steam drum then
travels down the right side down comers where it enters the mud drum. Here in the mud
drum is where the phase change and all the energy transfer occurs. The steam then
travels up the risers back into the steam drum where it passes through BFW. The BFW
acts as a scrubber where suspended water molecules produce superheated steam that
rises to the top of the steam drum. It then leaves the boiler through a fail-open control
valve to the main high pressure steam (HPS) header. The boiler will also have a safety
pressure release valve that will be rated for 60 bar. The steam header is where steam is
distributed to HPS users and to the Turbine Generators (J#1). The steady state blowdown
on the waste heat boiler is 3% of the total mass of steam produced. This blowdown water
is the same temperature and pressure as the boiler water. Therefore, the boiler feed
water contains thermal energy and will be used as cooling/make up water after being heat
integrated with a process stream in the plant to recover thermal energy and flashed to
recover low pressure steam (LPS) to be used in the deaerator.
Since passing steam through a turbine and a pressure release valve (PRV) are the only
two ways to reduce the pressure of a gas stream, turbines are favorable for power
recovery. As part of the waste heat boiler, a stream driven multi-stage automatic
extraction turbine is to be installed and rated for 50000 kW. This is #1J which will only
process steam. For simulations purpose the turbines were modeled with gas expanders
and steam traps where modeled with flash drums. In the process simulation J-201, J-202,
V-201, V-202, and the drives associated with represent one steam driven turbine and
generator.
5.2.2 Turbine Design
The turbine installed in this plant should include a governor. Blade (rotor) speed control is
highly important. Without it an over speed trip may occur causing the nozzle that controls
the steam flow to the turbine to close rapidly. This can cause the turbine to rapidly
accelerate until the rotor spins itself into a catastrophic failure. Since safety is our number
one priority in this design, all turbines will have a governor for speed control installed on
them and the proper standard operating procedures will be implemented to ensure a safe
work environment.
For a preliminary estimation of the turbine size, it was assumed that turbines act as
reverse compressors, and the turbines here sized with the same equations as for a
compressor. This assumption proved reasonable as the power calculations where within
95% confidence of the simulation outputs. When designing a turbine system the main
focus is to make sure the steam never drops below its superheated temperature and
become saturated. The steam flows through the turbine blades extremely fast. Water
droplets traveling in saturated steam vapors will cause pitting on the surfaces of the
turbine blades producing drag which will decrease the blades speed and in turn decrease
power production significantly. Many diverse areas of technology and discipline are
required to design a reliable, efficient steam turbine.
The Fischer-Tropsche synthesis reactors produce saturated steam in the shell side. This is
low pressure saturated steam, which is then superheated to 222oC and 23 bar, and sent
to a steam turbine (#2 J). #2 turbine generator is represented in the simulations and
process flow diagrams by two gas expanders and the steam traps are represent by flash

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95% confidence of the simulation outputs. When designing a turbine system the main
focus is to make sure the steam never drops below its superheated temperature and
become saturated. The steam flows through the turbine blades extremely fast. Water
droplets traveling in saturated steam vapors will cause pitting on the surfaces of the
turbine blades producing drag which will decrease the blades speed and in turn decrease
power production significantly. Many diverse areas of technology and discipline are
required to design a reliable, efficient steam turbine.
The Fischer-Tropsche synthesis reactors produce saturated steam in the shell side. This is
low pressure saturated steam, which is then superheated to 222oC and 23 bar, and sent
to a steam turbine (#2 J). #2 turbine generator is represented in the simulations and
process flow diagrams by two gas expanders and the steam traps are represent by flash
drums. There turbine/gas expanders and liquid turbines were sized using the method
described in Branan (2005). Based on the calculations that can be found in Appendix D
the turbine was designed to be rated for 100000 kW. This allows for a 10% increase in
throughput that can be achieved as the plant is expected to expand over its life.
Liquid Turbines are used in different units of this design to let down the pressure.
Turbines produce power; where as, pressure reducing valves do not. Liquid turbines are
pumps that run in reverse as a turbine and recover some of the available power. The
liquid turbines used would have to be sized using pump curves by a specific manufacturer.
Tapered reducers would have to be installed on the suction side of the liquid turbin. This
technology has been proven effective in an air separations plant. These turbines where
sized using the same formulas and logic as pumps.
5.3 Unit 300: FTR Feed Purification
The synthesis gas leaving Unit 200 is sent directly to the FTR Feed Purification unit, Unit
300, prior to being charged to the FTR unit to be converted into products. Unit 300 was
designed with the objective of removing approximately 98% of the water and greater
than 96.5% of the CO2. The CO2 removal method chosen was the Selexol process, which
uses a physical solvent consisting of a mixture of variations of the dimethylether of
polyethylene glycol to absorb CO2 from the syngas. The purification process occurs at
high pressure to improve the solubility of CO2 in the Selexol solvent. The purification
sequence is arranged in three main stages: water knockout and compression, CO2
absorption, and solvent regeneration.
5.3.1 Water Knockout and Compression
Prior to being compressed, the syngas exiting the final cooler in Unit 200, E-205, will be
flashed in V-301 to remove the bulk of the water. The initial compression of the syngas in
C-301 will increase the pressure from 19 psig to 70 psig. The pressurized feed will then be
cooled to 65 °C in E-301 by heat exchange with the purified FTR feed leaving the
purification process. The feed will then be split evenly and charged to the two Selexol
absorbers, T-301 A/B. The two trains for CO2 Absorption have been designed identically;
for this reason, only train A will be discussed. It should be assumed that all discussion
applies to both trains.
5.3.2 CO2 Absorption
The use of the Selexol process for acid gas removal from natural gas and syngas is
common practice in the chemical process industry. The advantages of using Selexol
include solvent stability, high CO2 loading capacity, and a high affinity for water that helps
improve the water removal from the process stream as well. From an environmental
standpoint, the Selexol solvent has proven to be non-toxic and biodegradable.
Additionally, since only CO2 is being removed from the process, solvent regeneration only
requires multi-stage flashing, unlike some other processes that require solvent strippers to
drive off absorbed CO2. Flashing requires much less energy than a stripping tower; thus,
Selexol was also found to be the most economic, environmentally and process friendly

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5.3.2 CO2 Absorption
The use of the Selexol process for acid gas removal from natural gas and syngas is
common practice in the chemical process industry. The advantages of using Selexol
include solvent stability, high CO2 loading capacity, and a high affinity for water that helps
improve the water removal from the process stream as well. From an environmental
standpoint, the Selexol solvent has proven to be non-toxic and biodegradable.
Additionally, since only CO2 is being removed from the process, solvent regeneration only
requires multi-stage flashing, unlike some other processes that require solvent strippers to
drive off absorbed CO2. Flashing requires much less energy than a stripping tower; thus,
Selexol was also found to be the most economic, environmentally and process friendly
CO2 absorption process available.
The CO2 absorber will be a large random packed column with a 6 m diameter, operated
at 69 psig. The packing chosen was 2 inch diameter 304 Stainless Steel Pall Rings. Metal
random packing is used in many absorption and stripping applications throughout the
chemical process industry and has proven to be an efficient method of removing
impurities from process streams. Turton confirmed that packed towers are ideal for
absorption application due to increased throughput and separation in comparison to tray
towers. Packing diameter was chosen based on a rule of thumb presented by Walas that
states for gas rates above 3400 m3/hr, 2 inch packing should be used (1990). The syngas
will be fed into the absorber on the last stage, given that the first stage is at the top of
the column. The lean, or having low CO2 concentration, Selexol solvent will enter the
absorber on the first stage at 13 °C, after being refrigerated in E-306 by exchange with
compressed propylene. The cost of this refrigeration will be discussed in the Utilities
section of this report. The treated syngas leaving the absorber was targeted to contain no
more than 3.5 wt% CO2, but the simulated absorption system actual yields a syngas
purity of less than 1 wt% CO2. The rich solvent containing the absorbed CO2 leaving the
bottom of the column will be regenerated through a series of decompression and flashing.
5.3.3 Solvent Regeneration
The regeneration sequence of the rich solvent will be completed in three flashes in V-303,
V-305, and V-306, a design which is outlined well in. The solvent pressure will be let down
to approximately half that of the absorber in J-301 prior to the first flash and let down
again to 300 psig in J-302 prior to the second flash. The vapor from these two flashes will
be recompressed in C-302 and C-303 to 1000 psig and recycled back to the feed stage.
This recycle method will reduce the losses of CH4, CO, and H2 in the CO2 product stream
because these substances are less soluble in Selexol than CO2 and will desorb first. The
liquid leaving V-305 let down to just above atmospheric pressure in J-303 and is heated to
32 °C through heat exchange with hot syngas from Unit 200. The final flash occurs in
V-306 where approximately 99.8 wt% of the absorbed CO2 is removed from the solvent.
The regenerated lean solvent is pumped by an array of parallel pumps, designated P-301
A-AF including spares, designed to handle the large mass flow required by the absorber.
There are eleven pumps operating in each absorption train in order to supply the
absorption columns with cool, lean solvent. A solvent make up flow of approximately 4.5
kg/hr will be injected into the lean Selexol stream prior to the P-303 pumps to account for
the quantity of solvent lost to the process and to counteract any degradation of the
existing solvent with time. The solvent make up rate was estimated by observing the
amount of solvent exiting Unit 300 in the CO2 and FTR Feed streams. The lean solvent
pumps supply the Selexol to E-306, where it will be cooled, as previously mentioned, prior
to being fed back into the absorber. The CO2 in the vapor stream leaving the final flash
drum of each train is recombined and piped to the CO2 compression section of Unit 400
for export.
5.3.4 CO2 Solubility and Selexol Flow Requirement

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kg/hr will be injected into the lean Selexol stream prior to the P-303 pumps to account for
the quantity of solvent lost to the process and to counteract any degradation of the
existing solvent with time. The solvent make up rate was estimated by observing the
amount of solvent exiting Unit 300 in the CO2 and FTR Feed streams. The lean solvent
pumps supply the Selexol to E-306, where it will be cooled, as previously mentioned, prior
to being fed back into the absorber. The CO2 in the vapor stream leaving the final flash
drum of each train is recombined and piped to the CO2 compression section of Unit 400
for export.
5.3.4 CO2 Solubility and Selexol Flow Requirement
The beginning of the Selexol absorber design process was based on the solubility of CO2
in Selexol. Chemstations provides Henry’s Law constants for CO2 solubility in Selexol,
which were cited from scholarly journals. The Henry’s Law constant, H, for CO2 at 25 °C
used for sizing of the absorbers was 3.6 MPa or 35.53 atm. The Henry’s Law constant at
25 °C was used as a conservative design parameter even though the Selexol is fed into
the column at 13 °C. Henry’s Law defines H as:
where yCO2 = mole fraction CO2 in the vapor phase
xCO2 = mole fraction of CO2 in the liquid phase
P = column pressure in bar
The mole fraction of CO2 in the syngas entering the column will be 0.094 moles of CO2
per mole of syngas, based on the ChemCAD simulation of the process. Also, the column
will be operated at a pressure near 69 atm. Through solution of Henry’s Law using these
process conditions and the Henry’s Law constant for CO2 in Selexol provided by
Chemstations, it was found that the mole fraction of CO2 in the Selexol leaving the
column will be approximately 0.14 moles of CO2 per mole of Selexol. This solubility
calculation yielded a targeted Selexol loading of 1.69 m3 CO2/kgmol Selexol, with a
Selexol molar flow rate equal to 65% of the vapor inlet molar flow. After the liquid flow
rate necessary to remove the desired quantity of CO2 was determined, analytical methods
were applied to determine the number of theoretical equilibrium stages, Nequil, height of
the theoretical stage, HETS, and column diameter.
5.3.5 McCabe-Thiele Analysis
To apply the McCabe-Thiele solution method to absorption processes, Wankat states that
several assumptions must be made regarding the nature of the system. The assumptions
necessary to apply this method are:
i) The heat of absorption is considered to be negligible.
ii) Temperature changes in the column are negligible.
iii) The solvent has a very low vapor pressure.
iv) The gas containing the solute, CO2, is insoluble in the solvent.
v) The mass flow rates of the carrier gas and the solvent are constant.
These assumptions allow a McCabe-Thiele diagram to be produced and analyzed in a way
very similar to that of a typical binary distillation. Instead of defining the axes of the
diagram as vapor and liquid mole fractions, however, the molar ratios of solute in the pure
carrier gas, YCO2, and in the pure solvent, XCO2, must be used. These molar ratios are
defined by the following equations:
These molar ratios were used as the y and x axes of the McCabe-Thiele plot.
As is typical for a McCabe-Thiele diagram, an equilibrium line and an operating line must
be plotted. The equilibrium line is assumed to be linear, with the y and x values being the
molar ratios described above, and the slope defined by a derivation of Equation 6:

REACTION
diagram as vapor and liquid mole fractions, however, the molar ratios of solute in the pure
carrier gas, YCO2, and in the pure solvent, XCO2, must be used. These molar ratios are
defined by the following equations:
These molar ratios were used as the y and x axes of the McCabe-Thiele plot.
As is typical for a McCabe-Thiele diagram, an equilibrium line and an operating line must
be plotted. The equilibrium line is assumed to be linear, with the y and x values being the
molar ratios described above, and the slope defined by a derivation of Equation 6:
The equilibrium line was defined as having a y-intercept of at the origin based on the
assumption that the lean solvent entering the first stage of the absorber will contain
essentially no CO2.
The operating line for the absorber is obtained by performing a mole balance around the
top of the column. The derived equation for the operating line presented in Wankat
(2007) is given as:
where Yn+1 = the molar ratio of CO2 to carrier gas at stage n+1
Xn = the molar ratio of CO2 to solvent at stage n
L = the constant molar flow of solvent in the column
G = the constant molar flow of carrier gas in the column
Y1 = the molar ratio of CO2 to carrier gas on the first stage
It should be noted that the y-intercept of this equation occurs at the molar ratio of CO2 to
carrier gas on the first stage because it is assumed that no CO2 is contained in the solvent
entering this stage, as previously mentioned. This operating line will be above the
equilibrium line, unlike in typical distillation diagrams. The McCabe-Thiele diagram
produced from these equations can be seen below in Figure 5.

REACTION
The McCabe-Thiele diagram was stepped off to determine the number of theoretical
stages in the absorber. It was found that the number of theoretical stages was
approximately 7. This value was checked using the Kremser equation and confirmed to be
correct within 10%. Details of this method are not included in this report since it was only
used as a check on the original design method.

REACTION
The McCabe-Thiele diagram was stepped off to determine the number of theoretical
stages in the absorber. It was found that the number of theoretical stages was
approximately 7. This value was checked using the Kremser equation and confirmed to be
correct within 10%. Details of this method are not included in this report since it was only
used as a check on the original design method.
Another important parameter that was determined using the McCabe-Thiele diagram was
the minimum acceptable solvent to carrier gas ratio, (L/G)min, to obtain the desired
syngas purity. The slope of the line plotted between the y-intercept of the operating line
and the point on the equilibrium line representing the molar ratios of CO2 in the gas and
solvent on the last stage of the absorber is equal to the minimum solvent to gas ratio
allowable. This value was found to be (L/G)min = 0.45. Therefore, it was concluded that
the calculated liquid to gas ratio of 0.65 was an acceptable value for the desired
separation.
A conservative estimate of the HETS = 1220 mm was assumed based on the rules of
thumb outlined in Turton. This value was used to calculate the height of the column based
on the number of stages calculated from the McCabe-Thiele analysis. Using this method,
the height of the column was estimated at approximately 9600 mm; an additional 3000
mm were added to the column as a rule of thumb also outlined by Turton. 1200 mm at
the top of the column are added to allow space for vapor disengagement, and 1800 mm
are added at the bottom of the column to account for liquid level. The pressure drop
across the packing was calculated to be around 7 kPa based on rules of thumb. These
sizing parameters were used to simulate and cost out the absorbers.
5.3.6 Absorber Diameter Calculation
The absorber diameter was determined using a generalized flooding correlation for packed
columns. The correlation uses a flow parameter, Flv, to determine the vapor flux, G’ [kg/
(s-m2)], in the column. The flow parameter is defined by the following equation:
where WL = the mass flow of the solvent in the column in kg/hr
WV = the mass flow of the carrier gas in the column in kg/hr
ρL = the density of the solvent in kg/m3
ρV = the density of the carrier gas in kg/m3
The calculated flow parameter was then used as the abscissa (x-axis value) in Figure 6.

REACTION
Figure 6: Flooding Correlation for Packed Towers (Wankat, 2007)
This figure is from Wankat’s Separation Process Engineering, 2nd ed. (2007); it is the
graphical representation of the previously mentioned flooding correlation for packed
columns. This graph was used to determine the ordinate (y-axis value) using the flooding
line shown in the figure. The ordinate value determined from the graph is at 100%
flooding. Typical flooding in packed towers is between 70% and 80% according to Wankat

REACTION
This figure is from Wankat’s Separation Process Engineering, 2nd ed. (2007); it is the
graphical representation of the previously mentioned flooding correlation for packed
columns. This graph was used to determine the ordinate (y-axis value) using the flooding
line shown in the figure. The ordinate value determined from the graph is at 100%
flooding. Typical flooding in packed towers is between 70% and 80% according to Wankat
(2007). An absorber flooding of 70% was assumed for design purposes. Therefore, the
ordinate value read from Figure 6 is multiplied by 0.70 to obtain the actual value used in
the diameter calculation. As can be seen in the figure, the ordinate is represented by the
equation below:
where F = packing parameter for specific packing and diameter [1/m]
ψ = the density of water divided by the density of the solvent
μ = the viscosity of the solvent in cP
gc = gravitational constant, [6.67408 × 10-11 m3 kg-1 s-2]
The packing factor for 2 inch metal Pall rings given by Wankat is 6 [1/m] (2007). Density
data and viscosity data for Selexol was provided by Edwards in his report for
Chemstations (2011) and is also built into the ChemCAD component database. Density of
the syngas entering the column was determined from the process simulation in ChemCAD.
Therefore, the only unknown in Equation 10 is the vapor flux, G’. The vapor flux was
solved for and used to calculate the cross sectional area, A, of the absorber via the
equation below:
Due to the imperfections associated with the packed tower correlation used in the area
calculation, the calculated area was increased by 32% as a safety precaution. The
calculated area per absorber, including safety correction, was approximately 26.5 m2. This
area was then used to calculate the tower diameter, which was found to be 5800m. A
summary of the design parameters presented in this discussion can also be found in the
Equipment Summary Tables in Appendix B.
5.4 Unit 400: Fischer-Tropsche Synthesis Reactors
The heart of every chemical plant is the chemical reactor. The gas to liquids plant is no
exception. The reactor converts carbon monoxide and hydrogen to alkanes via the
Fischer-Tropsche reaction. The Fischer-Tropsche reaction occurs by a carbon atom
attaching to the active site of the cobalt catalyst where it is reduced by hydrogen atoms.
The chain length increases as more carbons attach and are reduced by hydrogen atoms.
The alkane then separates from the catalyst. The average chain length is mainly affected
by the temperature of the reaction. The chain length has been correlated to the
temperature of the reaction by the Anderson-Shulz-Flory distribution, shown below in
Equation 12.
The alpha term is a function of temperature. As temperature increases the FT reaction
favors shorter alkanes and as temperature decreases the FT reaction favors longer
alkanes. For the scope of this project, the desired most product is diesel which is an
alkane of mid length ranging from C11 to C20 alkanes. The most important parameter to
consider when designing a reactor is temperature. In the case of the FT reaction both the

REACTION
The alpha term is a function of temperature. As temperature increases the FT reaction
favors shorter alkanes and as temperature decreases the FT reaction favors longer
alkanes. For the scope of this project, the desired most product is diesel which is an
alkane of mid length ranging from C11 to C20 alkanes. The most important parameter to
consider when designing a reactor is temperature. In the case of the FT reaction both the
selectivity of products and the reaction rate depend on temperature. Therefore, an
optimum operating temperature had to be evaluated which yields the appropriate
products at a reasonable conversion.
Modern chemical engineering design software such as ChemCAD cannot simulate the FT
reaction because of the wide array of products that are created in the reaction. The FT
reaction produces products from methane to C50+ alkanes. Approaching the problem of
simulating such a complicated reaction is a difficult task. The Polymath differential
equation solver along with chemical engineering intuition provided the necessary tools to
model the reaction.
The rate of disappearance of carbon monoxide and the selectivities of methane, ethane,
propane, and butane are known by the following equations:
Therefore, a system of differential equations for the flow of each of the compounds was
generated as a function of catalyst weight. A description of how the Polymath file works is
provided in the appendix. Although the flows of carbon monoxide, methane, propane,
ethane, and butane can be accounted for, the product flow is not described by the above
equations. Consequently, a carbon balance was performed. Since all of the carbons that
are dedicated to atoms C4 and below, the rest of the carbons must be dedicated to
alkanes with five or more carbons. A differential equation was generated to model the
molar flow of carbon atoms that were dedicated to alkanes with five or more carbon
atoms. This method was used in conjunction with a mass balance method. The mass flow
of all the compounds not included in the C5+ spectrum can be accounted for using the
given selectivity equations. Therefore, with the flow at the inlet to the reactor known, the
leftover mass flow rate must be dedicated to alkanes with five or more carbons. The
resulting mass flow rate can then be distributed by the Anderson-Shulz-Flory method, to
model the mass flow rate of each alkane. The mass and mole balances are only a small
step in accurately modeling the FT reaction. The temperature, pressure, and heat transfer
parameters are additional factors that must be taken into account.
The FT reaction is an extremely sensitive exothermic reaction. The risk of a runaway
reaction is a daunting threat which requires a reactor with excellent heat transfer. The
reaction was first modeled assuming an isothermal reaction to determine an operating
temperature. Using the Polymath software a graph of the conversion of carbon monoxide
versus reactor temperature was generated. Figure7 below displays the conversion of
carbon monoxide as a function of reactor temperature. As shown in the graph, the reactor
temperature of 500 degrees K allows for a conversion of around 90%. Temperatures
below 500K would require excessive amounts of catalyst to achieve a conversion of 90%.
The FT reaction is known to be a highly exothermic reaction, so maintaining the reactor
temperature in the vicinity of 500K where the conversion is high requires excellent heat
transfer to remove the heat generated by the reaction.

REACTION
Figure 7: Conversion of Carbon Monoxide over a Range of Inlet Temperatures
5.4.1 Pressure Drop
Pressure was the next parameter that was evaluated. In past operations, the FT reaction
has been carried out at pressures ranging between 25-35 bar; therefore a chart was
generated displaying conversion of carbon monoxide as a function of inlet pressure. The
pressure characteristics of the FT reaction were also difficult to model because it is a
mixed phase reaction. Initially, all of the reactants are in the gas phase, but as the
reaction proceeds a portion of the gas is transferred to the liquid phase. Therefore, as the
reaction occurs the pressure in the reactor will drop as gas is converted to liquid. The
assumption was made that this phenomenon can be taken into account by multiplying the
calculated pressure drop by 1.5 for all gas phase conditions. This assumption is intended
to model the pressure drop due to gas being converted liquid. However, this assumption
only slightly simplifies the calculation. The FT reaction is shown below:
Normally, pressure drop is calculated by the physical properties of the inlet fluid and the
reactor properties along with the stoichiometry of the reaction. The Ergun equation was
applied to account for pressure drop. The Ergun equation has two important terms that
must be calculated. The alpha term (not to be confused with the alpha in the Anderson-
Shulz-Flory distribution equation) takes into account the physical properties of the fluid
flowing through the reactor and the physical properties of the reactor itself. The epsilon
term accounts for the stoichiometrics of the reaction. Although the physical properties of
the inlet fluid and reactor characteristics are known, the reaction stoichiometrics are not
straight forward. It was previously determined that the optimum ratio of hydrogen to
carbon monoxide is 2:1 because that is the ratio that they are consumed in the FT
reaction. Therefore, it was assumed for pressure drop calculations that the stoichiometric
coefficient of hydrogen is two, and the coefficient for carbon monoxide is one. The
reaction shows that for every mole of carbon monoxide consumed one mole of water in

REACTION
must be calculated. The alpha term (not to be confused with the alpha in the Anderson-
Shulz-Flory distribution equation) takes into account the physical properties of the fluid
flowing through the reactor and the physical properties of the reactor itself. The epsilon
term accounts for the stoichiometrics of the reaction. Although the physical properties of
the inlet fluid and reactor characteristics are known, the reaction stoichiometrics are not
straight forward. It was previously determined that the optimum ratio of hydrogen to
carbon monoxide is 2:1 because that is the ratio that they are consumed in the FT
reaction. Therefore, it was assumed for pressure drop calculations that the stoichiometric
coefficient of hydrogen is two, and the coefficient for carbon monoxide is one. The
reaction shows that for every mole of carbon monoxide consumed one mole of water in
the gas phase is produced. However, the second product of the reaction is unknown. The
reaction could produce any alkane from methane to C50+. Consequently, the reaction
stoichiometrics vary. The goal of this evaluation is to determine an epsilon that is
representative of the reaction stoichiometrics. Since hydrogen and carbon monoxide are
being consumed, they are assigned a coefficient of -2 and -1. Water is produced in the
gas phase by the reaction so it is assigned a value of +1. However, the stoichiometrics of
alkanes are unknown. Conveniently, the selectivities of the alkanes are provided in the
equation 13 shown above. The stoichiometrics of the alkane products can be derived from
the selectivity equations. Notice, the selectivities are shown with respect to carbon
monoxide. The selectivity of methane is represented by the rate of formation of methane
divided by the rate of disappearance of carbon monoxide. This equation has the units:
(moles methane formed)/(moles of carbon monoxide consumed). This is exactly what the
stoichiometric coefficient of methane is. Therefore, an equation for epsilon can be derived
from the given selectivities. The equation used for epsilon in the Polymath program was:
Recall that from the mole balance the rate of C5+ was determined by subtracting the
known rates of the other alkanes. The equation below shows how the selectivity of C5+
alkanes was calculated.
The constant three was used because, rCn, is the rate of formation of the alkanes C2, C3,
and C4. The average carbon number for these three alkanes is three. Recall that this
method only yields the selectivity of moles of carbon dedicated to C5+ alkanes rather
than the actual moles of C5+ alkanes present. Although this selectivity value is not
representative of the actual number of moles present in the reactor, it was determined
later that the pressure drop in the reactor can be minimized making the accuracy of the
Ergun equation an insignificant factor.
The alpha term was calculated using the properties of the catalyst, the fluid properties at
the inlet, and the properties of the reactor. To simplify the efforts of the Polymath
software program the reaction was scaled down to a single reactor tube. A spreadsheet
was made that divides the mass flow of the inlet stream down to a single tube for a given
diameter. It was assumed that the tubes were spaced one inch apart and that the tubes
were triangularly packed.
The Ergun equation was used to model the pressure drop in the reactor. The alpha term
has a vast array of parameters that must be known to calculate it. However, the
calculation can be simplified by assuming that flow of the gases in the reactor is fully
developed turbulent flow. This assumption is justified by the fact that the heat transfer
coefficient is also a function of the mass flow through the reactor. Therefore, the mass
flow will most likely be fully developed turbulent flow to achieve the best heat transfer
since heat transfer improves as the superficial mass velocity is increased. The terms of
the Ergun equation are shown below:
If fully developed turbulent flow is assumed, then the first term in the brackets can be
deemed insignificant and dropped because the “1.75 G” term dominates. This assumption
was made for the purposes of modeling the pressure drop in the FT reactors. This
assumption also eliminated the viscosity term of the calculation which further simplified
the effort. In addition, it was assumed that the gas entering the reactor behaves as an

REACTION
If fully developed turbulent flow is assumed, then the first term in the brackets can be
deemed insignificant and dropped because the “1.75 G” term dominates. This assumption
was made for the purposes of modeling the pressure drop in the FT reactors. This
assumption also eliminated the viscosity term of the calculation which further simplified
the effort. In addition, it was assumed that the gas entering the reactor behaves as an
ideal gas for the pressure drop calculations. The density of the gas was calculated using
the ideal gas law. The spreadsheet calculated the ideal gas density and the superficial
mass velocity for a given flow of inlet gas through a single reactor tube. Therefore, the
beta term and the alpha term were calculated in the spreadsheet based on the inlet
temperature, inlet pressure, inlet molar flow, reactor diameter, and tube diameter. Once
the alpha term was calculated it could be plugged in to the Ergun equation which was
coupled with the other differential equations in the Polymath software program. The
Ergun equation was used to calculate the pressure drop as shown below.
The “1.5” factor included in the equation is meant to model the amount of gas that is
being transferred into the liquid phase. This differential equation generates a term “y” for
a given amount of catalyst weight. The “y” term is the fraction of the actual pressure
divided by the initial pressure. Therefore, to find the actual pressure at given catalyst
weight, the “y” term can be multiplied by the inlet pressure.
The pressure drop calculation is a complicated and in depth effort, but it was simplified by
using excel spreadsheets that iterated the calculations without having to perform them by
hand. Once the Ergun equation was running correctly in the Polymath program, pressure
studies were performed for the reactors. The first chart generated was intended to
evaluate the effects of inlet pressure on the conversion of carbon monoxide in the
reactors. The conversion of carbon monoxide was observed over a range of inlet
pressures ranging from near zero to 50 atm. Figure 8 shows how inlet pressure affects
conversion for a given weight of catalyst. Initially, inlet pressure plays an important role in
the conversion, but as the pressure nears 20 atm the graph turns linear. Consequently,
the minimum value for pressure at the inlet that should be considered is 20 atm. Even
though conversion improves as the inlet pressure is increased, the reactor tube wall
thickness must be increased to accommodate the high reactor pressures. Also, the
pressure drop will be more significant for high pressures. Recall that the maximum
pressure drop that can be allowed per reactor is 3.5 bar. In conclusion, an inlet pressure
of 30 atm was chosen to raise conversion without violating the pressure drop per reactor
ceiling and to avoid thick tube walls.
5.4.2 Heat Transfer
Heat transfer is an extremely important factor to consider in the design of the FT
reactors. The FT reaction is extremely exothermic. As the reaction proceeds to
completion, energy is generated by the reaction that must be removed by the cooling

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