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TRANSCRIPT
1
Plant Process Design and
Economic Analysis
of a
Proposed Process of a Gas to Liquid Plant
(GTL) from Methane Fueled Synthesis Gas via
the Fischer-Tropsch Reaction
2
Table of Contents
List of Figures ................................................................................................................................. 4
List of Tables .................................................................................................................................. 4
List of Equations ............................................................................................................................. 5
Executive Summary ........................................................................................................................ 7
Introduction ..................................................................................................................................... 9
Summary ....................................................................................................................................... 10
Discussion ..................................................................................................................................... 12
Process Description ................................................................................................................... 12
Equipment Design ..................................................................................................................... 14
Syngas Unit............................................................................................................................ 14
Air Separation unit................................................................................................................. 17
Fischer-Tropsch Reactor ........................................................................................................ 17
Hydro-Isomerization Unit...................................................................................................... 23
Product Separations ............................................................................................................... 23
Material and Energy Balance Discussion.................................................................................. 31
Safety and Environmental Summary......................................................................................... 32
Unit Control and Instrumentation Description .......................................................................... 36
Conclusions ................................................................................................................................... 37
Recommendations ......................................................................................................................... 38
Project Premises and Assumptions ............................................................................................... 39
References ..................................................................................................................................... 41
Appendix ....................................................................................................................................... 44
Appendix A.1 – Process Flow Diagrams .................................................................................. 44
Appendix A.1.1 – Synthesis Gas Unit....................................................................................... 45
Appendix A.1.2 – Fischer-Tropsch Reactor Unit ..................................................................... 50
Appendix A.1.3 – Diesel and Wax Separation Unit ................................................................. 51
Appendix A.1.4 – Naphtha Separation Unit ............................................................................. 52
Appendix A.2 –Economics Summary ....................................................................................... 69
Appendix A.3 – Equipment Sizing Summary ........................................................................... 72
Appendix A.4 – Computer Simulation Outputs ........................................................................ 79
3
Synthesis Gas Unit MATLAB Design Function ................................................................... 79
Synthesis Gas Unit MATLAB Design Output ...................................................................... 83
Fischer-Tropsch Reactor MATLAB Design Function .......................................................... 85
Fischer-Tropsch Reactor MATLAB Design Function .......................................................... 92
Appendix A.5 – Supporting Graphs .......................................................................................... 94
Appendix A.6 – Sample Calculations ..................................................................................... 102
4
List of Figures
Figure 1. Plant Process Flow Diagram ......................................................................................... 44
Figure 2. Synthesis Gas Unit PFD ................................................................................................ 45
Figure 3. ASPEN Diagram for the Syngas Unit Heat Exchangers ............................................... 46
Figure 4. Fischer-Tropsch Reactor Unit PFD ............................................................................... 50
Figure 5. Diesel and Wax Separation Unit PFD ........................................................................... 51
Figure 6. Naphtha Separation Unit PFD ....................................................................................... 52
Figure 7. ASPEN Diagram for Diesel and Wax Separation Unit ................................................. 53
Figure 8. ASPEN Diagram for Naphtha Separation Unit ............................................................. 54
Figure 9. Sensitivity Analysis of the Plant.................................................................................... 78
Figure 10. Syngas Temperature Optimization .............................................................................. 94
Figure 11. Syngas Water to Methane Ratio Optimization ............................................................ 94
Figure 12. Reactor Temperature Optimization ............................................................................. 95
Figure 13. Reactor Cost Optimization .......................................................................................... 95
Figure 14. Graph of Conversion vs. Catalyst Weight in a Single Tube...................................... 100
Figure 15. Graph of Percent of Pressure Drop vs. Weight of Catalyst ....................................... 100
Figure 16. Graph of Temperature vs. Catalyst Weight of a Single Tube ................................... 101
Figure 17. Graph of Molar Flow Rates of Carbon Monoxide, Hydrogen, and Water vs. Catalyst
Weight in a Single Tube ............................................................................................................. 101
List of Tables
Table 1. Economic Summary .......................................................................................................... 8
Table 2. Syngas Effluent and Feed Streams 1-DIOXIDE ............................................................ 47
Table 3. Syngas Effluent and Feed Streams FTR-MP2OUT ........................................................ 48
Table 4. Syngas Effluent and Feed Streams O2TOSYN-TOSYN ............................................... 49
Table 5. Stream Table for Separation Streams 1-6 ....................................................................... 55
Table 6. Stream Table for Separation Streams 7-12 ..................................................................... 56
Table 7. Stream Table for Separation Streams 13-18 ................................................................... 57
Table 8. Stream Table for Separation Streams 19-24 ................................................................... 58
Table 9. Stream Table for Separation Streams 25-30 ................................................................... 59
Table 10. Stream Table for Separation Streams 31-36 ................................................................. 60
Table 11. Stream Table for Separation Streams 37-42 ................................................................. 61
Table 12. Stream Table for Separation Streams 43-48 ................................................................. 62
Table 13. Stream Table for Separation Streams 49-AIR .............................................................. 63
Table 14. Stream Table for Separation Streams FuelAir-Waste H2O .......................................... 64
Table 15. Stream Table for Separation Streams FuelAir-Waste H2O .......................................... 65
Table 16. Stream Table for Heat Integration Cool1-1 to Cool2-3 ................................................ 66
Table 17. Stream Table for Heat Integration Cool2In to LP1Out ................................................ 67
Table 18. Stream Table for Heat Integration MP1In and MP1Out............................................... 68
5
Table 19. Economic Summary Breakdown .................................................................................. 69
Table 20. Utilities and Credits Breakdown ................................................................................... 70
Table 21. Products Breakdown ..................................................................................................... 70
Table 22. Inflation Effects on Profitability ................................................................................... 71
Table 23. Energy Efficiency ........................................................................................................ 71
Table 24. Economics Summary .................................................................................................... 71
Table 25. Equipotent Costs Summary ......................................................................................... 72
Table 26. Heat Exchanger Sizing Summary, Heat Exchangers 1-8 ............................................. 73
Table 27. Heat Exchanger Sizing Summary, Heat Exchangers 9-16 ........................................... 73
Table 28. Flash Drum Sizing Summary ........................................................................................ 74
Table 29. Decanter Sizing Summary ............................................................................................ 74
Table 30. Distillation Column Sizing Summary ........................................................................... 75
Table 31. Compressor Sizing Summary ....................................................................................... 75
Table 32. Fired Heater Sizing Summary....................................................................................... 75
Table 33. Refrigeration Summary ................................................................................................. 76
Table 34. Reactor Sizing Summary .............................................................................................. 76
Table 35. Syngas Unit Process Control Summary Table ............................................................. 77
Table 36. Fischer-Tropsch Reactor Control Summary Table ....................................................... 77
Table 37. Distillation Control Summary Table ............................................................................. 77
Table 38. Flash Drum Control Summary Table ............................................................................ 78
Table 39. Syngas Flow Rates for Temperature Optimization...................................................... 96
Table 40. Syngas Cost Optimization for Temperature Optimization ........................................... 96
Table 41. Syngas Flow Rates for Water to Methane Ratio Optimization .................................... 96
Table 42. Syngas Costs for Water to Methane Ratio Optimization .............................................. 97
Table 43. Material Balance for Syngas and FTR .......................................................................... 98
Table 44. Material Balance for Syngas and FTR 2 ....................................................................... 99
List of Equations
Equation 1. Steam Reforming Equilibrium Kinetics .................................................................. 102
Equation 2. Water-Gas Shift Reaction Equilibrium Kinetics ..................................................... 102
Equation 3. Anderson-Shulz-Flory (ASF) equation ................................................................... 102
Equation 4. Selectivity of Methane ............................................................................................. 102
Equation 5. Selectivity of Ethane, Propane, and Butane ............................................................ 102
Equation 6. Design Equation for a Packed-Bed Reactor ............................................................ 103
Equation 7. Fischer-Tropsch Rate Equation ............................................................................... 103
Equation 8. Ergun Equation to Model Pressure Drop in Packed-Bed Reactors ......................... 103
Equation 9. Temperature Gradient Calculation .......................................................................... 104
Equation 10. Thickness for Cylindrical Shells ........................................................................... 104
Equation 11. Thickness for Torispherical Heads ........................................................................ 104
6
Equation 12. Volume of a Torispherical Head ........................................................................... 104
Equation 13. Cost of Steel for a Pressure Vessel ........................................................................ 104
Equation 14. Maximum Vapor Velocity for a Flash Drum ........................................................ 105
Equation 15. Number of Tubes in a Shell ................................................................................... 105
7
Executive Summary
The use of a gas to liquid plant to produce diesel and naphtha from methane is a viable
operation to bring to market newly discovered natural gas deposits in a remote location.
Production of liquid fuels from a variety of carbon sources via synthesis gas production and
Fischer-Tropsch synthesis has been used in a variety of operations since the early 20th
century.
This design and economic analysis investigates the commercial viability of constructing a gas to
liquid plant to convert 500 MSCF/Day of natural gas to diesel and naphtha via synthesis gas
production and Fischer-Tropsch synthesis.
Synthesis gas production and Fischer-Tropsch synthesis involve the use of many
potentially harmful substances that could endanger employees and the environment if handled
improperly. It is important to ensure that all equipment design and plant layout are developed
with safety and environmental safeguards to protect employees and the environment in the event
of an emergency. All plant personal should be properly trained in the proper use of personal
safety equipment and the safe operation of equipment within the facility.
Synthesis gas production occurs at high temperatures to react methane with steam to
produce synthesis gas. The resulting synthesis gas is composed primarily of carbon monoxide
and hydrogen. The synthesis gas is then fed into a Fischer-Tropsch reactor where a
polymerization reaction takes place to form longer chain hydrocarbons. These longer chain
hydrocarbons contain LPG, naphtha, diesel, and petroleum waxes which are subsequently
separated to be sold commercially.
The proposed plant design has the potential of producing 14.6 million barrels of naphtha
and 10.5 million barrels of diesel per year. Recent trends indicate that liquid fuel prices will
continue to rise which suggests that the conversion of natural gas to liquid fuels will continue to
be viable during the lifespan of the plant. The overall efficiency of the plant is 91% and recovers
99.5% of the diesel and 98.8% of the naphtha produced in the FTR. The economic analysis
suggests that the cash flow payback period will be 2.36 years and has a DCFRR of 42.2%.
8
Table 1. Economic Summary
A majority of the fixed costs for the project stem from the cost of the reactors while a
majority of the manufacturing costs are associated with separating oxygen for synthesis gas
production. Prior to the construction of the plant an investigation should be completed to
improve the conversion in reactor. If the conversion can be increased the volumetric flow
though the separation would be dramatically decreased as well as increase the efficiency of the
plant. Also, prior to construction it should be investigated that the plant’s utility grid can handle
the increase in demand associated with this process or if new facilities need to be constructed.
Plant Operation Life (years) 15
Total Capital Investment ($) 1,020,494,941.82$
Total Revenue ($/year) 2,920,420,070.88$
Total Manufacturing Costs ($/year) 2,375,953,263.52$
Net Profit ($/year) 364,792,760.93$
Cash Flow ($/year) 432,825,757.05$
DCFRR (%) 42.20%
Cash Flow Payback Period (years) 2.36
Return on Investment (%) 35.75%
Net Present Worth of Plant ($) 2,684,267,901.25$
Economics Summary
9
Introduction
The recent discovery of a remote natural gas deposit has prompted an investigation into
the feasibility of a gas to liquid plant to bring these resources to market. Recent trends indicate
that rising gasoline and diesel prices in combination with abundant natural gas produces a
favorable environment for such a facility. Production of liquid fuels from a variety of carbon
sources via synthesis gas production and Fischer-Tropsch synthesis has been used in a variety of
operations since the early 20th
century. This report was compiled to determine the feasibility of a
gas to liquid plant to convert 500 MSCF/Day to liquid fuels via synthesis gas and Fischer-
Tropsch synthesis.
10
Summary
The proposed gas to liquid plant is designed to create valuable liquid fuels from remote
natural gas deposits via synthesis gas production and Fischer-Tropsch synthesis. Synthesis gas
production is done using steam reforming to produce carbon monoxide and hydrogen from
methane and water. This reaction is endothermic and requires a partial oxidation of methane to
maintain the desired reaction temperature. The proposed design converts 500 MSCF/Day of
natural gas to synthesis gas which is then fed into a packed-bed Fischer-Tropsch reactor where
an ultra-stable cobalt catalyst is used to polymerize long chain hydrocarbons from synthesis gas.
The product dispersion of the reaction is dependent on temperature where lower temperatures
encourage the formation of longer chain hydrocarbons. This reaction is extremely exothermic
and requires close monitoring of the reactor temperature to prevent a runaway reaction and
maintain the desired reaction temperature.
The syngas unit operates at 1950°F to increase the conversion of methane to synthesis
gas and decrease the amount of oxygen necessary to maintain the reactor temperature. The
resulting synthesis gas is used to preheat the fed to the synthesis gas unit and create steam for use
in other parts of the plant. Synthesis gas is then fed into a Fischer-Tropsch reactor that operates
at 425°F to encourage the formation of longer chain hydrocarbons. The reactor temperature is
maintained by varying the pressure on the shell side of the reactor to change the boiling
temperature of the cooling water. Products from the reactor include a wide range of
hydrocarbons including LPG, naphtha, diesel, petroleum wax, and unconverted carbon monoxide
and hydrogen. These products are then fed into a separation system that uses a series of flash
drums, distillation columns, strippers, absorbers, and decanters to separate the various products
to the desired purities. A summary of the equipment specifications can be viewed in Table 25 in
Appendix A.3.
Plant safety and environmental stewardship are major concerns taken into account in the
design of this plant. All plant employees must be trained in the proper use of personal safety
equipment and emergency response procedures. To protect the environment the venting of
hydrocarbons should be avoided during normal operations and any hydrocarbons that are not
used in the final products should be burned in a fired heater to produce steam or electricity for
the plant.
11
This plant is capable of producing 14.6 million barrels of naphtha and 10.5 million
barrels of diesel per year. The separation process recovers 99.5% of the diesel and 98.8% of the
naphtha produced in the FTR and has an overall plant efficiency of 91%. Total capital
investment for the process was determined to be $1 billion by multiplying the total equipment
costs by a Lang factor of 4.8. The resulting DCFRR for the process is 42.2% with a cash flow
payback period of 2.4 years and a return on investment of 35.75% per year.
12
Discussion
Process Description
The recent discovery of a remote natural gas deposit has prompted the investigation of
utilizing gas to liquid (GTL) technology to bring these resources to market. This technology
uses a synthesis gas unit to turn 500 MSCF/Day into synthesis gas that will be converted into
liquid fuels and other products using the Fischer-Tropsch reaction. The synthesis gas production
unit uses steam reforming in combination with a controlled partial oxidation of methane to
produce synthesis gas to be fed into the Fischer-Tropsch reactor. In steam reforming, steam and
methane create an equilibrium reaction to produce carbon monoxide and hydrogen. In order to
increase the temperature of the syngas unit to the desired temperature, a controlled amount of
oxygen is fed into the reactor to partially combust the methane into carbon monoxide and water.
There is an additional water-gas shift equilibrium reaction that combines carbon dioxide and
hydrogen to produce carbon monoxide and water. These reactions are shown below in reactions
1, 2, and 3.
Steam Reforming Reaction
⇔
Partial Oxidation Reaction
⇒
Water-Gas Shift Reaction
⇔
These reactions are extremely temperature dependent and a change in temperature can
dramatically change the resulting amounts of carbon monoxide and hydrogen. This makes
controlling the oxygen feed rate and thus the partial oxidation of reaction extremely important in
controlling the conversion for the syngas unit. The synthesis gas effluent is then cooled and fed
into a packed bed Fischer-Tropsch reactor with a cobalt based catalyst to create liquid fuels. The
Fischer-Tropsch reaction is a chaingrowth reaction that polymerizes alkane building blocks to
create longer chain alkanes (Yates & Satterfield, 1991). These reactions are shown below in
reactions 4 and 5
(1)
(2)
(3)
13
Overall Fischer-Tropsch Reaction
( ) ⇒ ( )
Fischer-Tropsch Mechanism
⇒ ⌊ ⌋
The Fischer-Tropsch reaction is also very temperature dependent and requires close
monitoring of the reactor temperature in order to achieve the desired product distribution. Lower
temperatures encourage the formation of longer alkanes including diesel and petroleum wax,
while higher temperatures encourage the formation of methane and short chain hydrocarbons
including liquefied petroleum gas (LPG) and naphtha. The temperature of the reactors is
controlled by boiling pressurized water at the desired ambient temperature to control the
reaction. The resulting reactor effluent is then fed into the separation unit where the various
products are separated.
(4)
(5)
14
Equipment Design
Syngas Unit
The synthesis gas production unit uses steam reforming in combination with a controlled
partial oxidation of methane to produce synthesis gas to be fed into the Fischer-Tropsch reactor.
In steam reforming, steam and methane create an equilibrium reaction to produce carbon
monoxide and hydrogen. This reaction is very endothermic and requires high temperatures and a
large amount of energy to maintain this reaction. To operate the syngas unit at higher
temperatures beyond the maximum preheater temperatures of 1000⁰F, a partial oxidation of
methane is used to increase the temperature to an operating temperature of 1950⁰F. The amount
of oxygen fed into the reactor is closely controlled to ensure that the desired temperature is
maintained. This reaction is irreversible and it is assumed that all of the oxygen fed into the
reformer is consumed to produce carbon monoxide and water in the reaction. There is an
additional water-gas shift equilibrium reaction that combines carbon dioxide and hydrogen to
produce carbon monoxide and water. These reactions are shown below in reactions 1, 2, and 3.
Steam Reforming Reaction
⇔
Partial Oxidation Reaction
⇒
Water-Gas Shift Reaction
⇔
The synthesis gas production unit is designed to convert 500 MSCF/Day of methane into
carbon monoxide and hydrogen. Prior to being fed into the syngas unit, the feed is preheated
using a series of heat exchangers with the syngas unit effluent. These heat exchangers increase
the temperature of the feeds from 100⁰F to 1000⁰F and reduce the heating requirements for the
process. The alternative to using a heat exchanger would be to preheat the steam using fuel gas
and steam or having to combust more methane using the partial oxidation to control the syngas
unit temperature. A fired heater will still be necessary for transient states and startup procedures.
However during normal operations the heat exchangers will be sufficient to properly preheat the
feeds and will not need any additional utilities. Increasing the amount of methane being partially
(1)
(2)
(3)
15
oxidized was also not pursued because the oxygen separation accounts for approximately 60% of
the variable costs for the process and was avoided to minimize the oxygen required by the syngas
unit. Separate preheat heat exchanger are provided for the oxygen and methane feeds. This is to
ensure that the methane does not prematurely combust in the heat exchangers and cause a hazard
during the operation of the plant.
The resulting syngas unit effluent composition was determined by developing a
MATLAB file that solves a series of equations to determine the carbon dioxide molar feed rate,
and the extents of the three reactions occurring in the reformer. The MATLAB code can be
viewed in Appendix A.4. This program solves an energy balance around the syngas unit to
determine the necessary extent of reaction two, the partial oxidation reaction and thus the feed
rate of oxygen, to maintain the desired temperature. It was also assumed that the desired effluent
molar ratio of hydrogen to carbon monoxide should be equal to two to produce the desired
hydrocarbons. The remaining two equations used in the program were equating the equilibrium
kinetics to the effluent concentrations of the products.
A range of syngas unit operating conditions were considered by varying the temperature,
pressure and steam to methane feed ratio. The temperature was varied between 1600⁰F and
1950⁰F while the steam to methane feed ratio was varied between 0.5 and 1.5. This optimization
can be viewed in Appendix A.5 in Figures 10 and 11 and Tables 39-42.
Due to the constraints of the equilibrium kinetics used for the steam reforming and water-
gas shift reaction, the reactor pressure could not be varied with pressure because the units that
were provided did not include pressure dependence (Wang & Michael, 2009). This means that
the partial pressures of the components could not be used to determine the syngas unit effluent
because the kinetics made it only dependent on temperature. During the design and optimization
of the syngas unit it was assumed that the lowest operating pressure would be the best for
conversion. In the equilibrium there are more moles of products being formed than reactants in
the steam reforming reaction and the water-gas shift reaction had an equal amount of moles for
the both products and reactants. This means that a lower pressure would promote the formation
of the desired products and increase the conversion of methane in the steam reforming reaction,
while a higher pressure would promote the formation of methane and would not affect the water-
gas shift reaction.
16
In steam reforming, steam and methane create an equilibrium reaction with carbon
monoxide and hydrogen. At higher temperatures the equilibrium kinetics encourage the
formation of carbon monoxide and hydrogen, however this reaction very endothermic with a heat
of reaction of 209.22 kJ/gmol (89,972.08 BTU/lbmol) (Wang & Michael, 2009). At high
temperatures the amount of carbon monoxide and water formed in the reaction is reduced. The
water-gas shift reaction is normally exothermic with a heat of reaction of -42 kJ/gmol (-
17631.1581 BTU/lbmol) (Wang & Michael, 2009). However additional carbon dioxide is fed
into the syngas unit to produce more carbon monoxide and water, making this reaction
endothermic in the reformer. It was assumed that the syngas unit operated isothermally and the
only energy input into the unit came from the two equilibrium reactions and the partial oxidation
reaction occurring within the reformer. The equilibrium kinetics for these reactions is shown in
Appendix A.6 in Equations 1 and 2.
During normal operation, additional water is fed into the syngas unit at a ratio of 0.5
moles of water to every mole of methane fed into the reactor. This is done to prevent any carbon
coking that might occur and protects the unit from requiring additional heat if the carbon coke
accumulated. The additional water also serves to drive the equilibrium kinetics for the steam
reforming reaction to form more carbon monoxide and hydrogen. Additional carbon dioxide is
also fed into the syngas unit to encourage the formation of carbon monoxide and water in the
water-gas shift reaction. The carbon dioxide feed rate is closely controlled to maintain the water-
gas shift reaction to maintain a carbon monoxide to hydrogen ratio of two.
The syngas unit was optimized by trying various operating conditions and plotting the
cost of the utilities against the temperature. Oxygen was the most expensive utility due to the
amount of energy and cooling water necessary to separate the oxygen from air at a cost of
approximately $2/lbmol. The second biggest cost was carbon dioxide at approximately
$0.15/lbmol and the cost of steam was the smallest at approximately $0.09/lbmol. These graphs
can be seen in Appendix A.5 in Figures 12 and 13.
An operating temperature of 1950°F and a water to methane ratio of 0.5 proved to be the
most economical conditions for the syngas unit. At higher temperatures the amount of carbon
monoxide and hydrogen formed from the methane is dramatically increased compared to
conditions at lower temperatures. This shift in the kinetics requires less energy inputs from the
17
partial oxidation and reduces the oxygen requirements for the reactor. Also, the higher
temperature requires a lower feed rate of carbon dioxide to achieve the desired hydrogen to
carbon monoxide ratio. At these temperatures more carbon dioxide is used in the water-gas shift
reaction to form more carbon monoxide and hydrogen for the same flow rate of methane. This
temperature also has the highest conversion of methane to synthesis gas and consumes more
water in the reaction. This reduces the amount of carbon dioxide and water exiting the syngas
unit makes the subsequent Fischer-Tropsch reactor smaller and reducing cost of the reactors.
Higher water to methane ratios offered a reduced oxygen flow rate however, the additional cost
of steam and carbon dioxide proved to be more expensive.
Air Separation unit
Oxygen for the syngas unit is purchased from a third party oxygen plant for $100 per
short ton of oxygen. The oxygen supplied is 99mol% oxygen and 1mol% nitrogen.
Additionally, the plant is required to provide either electricity of 600 psig steam to power the
compressors used in the separation. It was determined that the steam option was most
economical due to the ability to recycle the steam condensate to produce more steam. 400 gpm
per short ton per day of cooling water is also required by the oxygen separation and is the major
cost for the separation.
Fischer-Tropsch Reactor
Fischer-Tropsch can be used in combination with synthesis gas production to create
liquid fuels from either solids including coal, gases such as methane or other carbon sources such
as biomass. The Fischer-Tropsch reactor used in this design is a packed-bed reactor with an
ultra-stable cobalt based catalyst to polymerize the carbon monoxide and hydrogen into longer
chain hydrocarbons and water (Yates & Satterfield, 1991). This reaction occurs though the
addition of alkane building blocks to an existing chain to create longer chain alkanes (Wang &
Michael, 2009). This reaction is shown in reaction 4 and the mechanism is shown in reaction 5.
Overall Fischer-Tropsch Reaction
( ) ⇒ ( )
Fischer-Tropsch Mechanism
⇒ ⌊ ⌋
(4)
(5)
18
The alkane selectivity and product distribution of the Fischer-Tropsch reaction is
dependent on temperatures. This product distribution is dictated by the Anderson-Shulz-Flory
(ASF) equation to model the distribution for the C5+ carbon alkanes while the selectivities of
methane though butane is dictated as a function of temperature. The ASF equation was specially
developed to characterize the chain growth and product distribution of alkanes in the Fischer-
Tropsch reaction. The ASF equation can be viewed in Equation 3 in Appendix A.6. The
methane though butane selectivity equations can be viewed in equations 4 and 5 in Appendix
A.6.
Since the ASF equation was not used for all of the components and the selectivity of the
C1 to C4 hydrocarbons were lower than what would be predicted by the ASF distribution, the
selectivities of the products did not sum to one. In order to correct this imbalance the sum of the
selectivities was normalized to equal one by multiplying the selectivities by a correction factor.
The selectivities were then divided by the number of carbons in that component to get the
selectivity into units of per mol of carbon monoxide fed.
The product distribution from the packed bed reactor is determined by the inlet
temperature of the reactor. A temperature range of 390⁰F to 450⁰F was examined to determine
what temperature would yield the most desirable product distribution. To determine the most
profitable operating conditions the values of the products were calculated based on the selectivity
for one mole of carbon monoxide fed. For this optimization it was assumed that the subsequent
separation operated perfect, except for the waxy components. It was assumed that a basis of
30% diesel and 70% wax by weight was fed into the hydro-isomerization unit to determine the
value of the waxy components. This optimization is shown in Figure 12 in the Appendix A.5.
Based on the optimization, the maximum potential profit from the reactor is at an
operating temperature of approximately 400⁰F, however at such lower temperatures the price of
the reactors necessary to achieve the desired conversion becomes extremely expensive compared
to higher reactor temperatures. To compare the price of a reactor to the reactor temperature, the
price reactors at various temperatures were plotted against the corresponding temperature. The
price of the reactors was determined by using the weight of the steel to price the reactors and
assuming 15 year straight line depreciation (Peters, Timmerhaus, & West, 2003). The price of
catalyst was also factor in by dividing the amount of catalyst needed for the four year life and
19
dividing the price over a four year span. This comparison can be seen in figure 13 in the
Appendix A.5.
A balance between reactor size and potential product was determined to be around a
reactor temperature of 425⁰F. At this temperature the reactor has the potential to produce an
additional $108 million per year in potential product compared to the highest operating
temperature of 450⁰F. This temperature was chosen because the slope of the revenue per year
decreases as it approaches the optimum temperature. Also, the estimated yearly cost of the
reactor for this temperature was less than half of the cost of a reactor at the optimum
temperature. At the optimum operating temperature of 400⁰F the additional product value
produced is approximately equal to the cost of the new reactors and does not justify the
additional cost.
Modeling of the temperature, conversion, pressure drop, and product accumulation for
the reactor was done using a MATLAB program to model a single tube within the reactor. It
was assumed that all of the tubes within the reactor system would behave similarly the modeled
tube. The MATLAB file solved for three separate differential equations using ode15. The
conversion for a single tube was modeled using the design equation for a packed-bed reactor
(Fogler, 2008). This equation is shown in Equation 6 in Appendix A.4.
The reaction rate for the conversion of synthesis gas to liquid fuels using the Fischer-
Tropsch model was determined to be best fit by the Langmuir-Hinshelwood rate equation. The
rate equation is shown below in Equation 7 in Appendix A.6. It was assumed that the vapor
phase within the reactor contains all of the hydrogen, carbon monoxide, water, methane, and
nitrogen components and 70% of the moles for C2+ hydrocarbons for every mol of CH4 made.
To account for this formation of liquid products within the reactor is was assumed that the total
pressure drop within the reactor was 1.5 times the gas phase pressure drop. It was also assumed
that the density and viscosity of the components was constant though the reactor. The pressure
drop within the reactor was modeled using the Ergun equation because it is analytical equation to
determine the pressure drop in packed-bed reactors (Fogler, 2008). This equation is shown below
in Equation 8 in Appendix A.6.
20
The temperature gradient within the tubes in the reactor was modeled using an energy
balance (Fogler, 2008). This equation is shown in Equation 9 in Appendix A.6. It was assumed
that the shell side of the reactor operated isothermally and that all of the energy from the reactor
was used to vaporize water in the shell of the reactor. The vaporization of water allows the
ambient temperature of the reactor to be assumed constant and that there is no external
temperature gradients.
Since the temperatures of the components within the tubes were not constant, the
MATLAB file was designed to incorporate heat capacity calculations that would vary with the
internal temperature of the tube. This would allow for more accurate analysis of the reactor
design to account for potential problems that could arise from limitations in the heat transfer of
the tubes. The Fischer-Tropsch reaction is extremely exothermic with a heat of reaction of
70,200 BTU per lbmol of carbon monoxide. If the temperature in the shell of the reactor is not
maintained at the desired operating temperature there is a serious risk of a runaway reaction.
The cooling water for the reactor is preheated to a saturation liquid and pressurized to the
desired pressure before being fed into the reactor. Condensed steam was chosen as the cooling
water because it is already preheated near the desired temperature and is 99.9% pure water. This
is to ensure that the shell of the reactor will not corrode or accumulate any solids that might
hinder the heat transfer capability of the reactor. The boiling temperature of the water is
controlled by changing the pressure inside of the shell. In case of emergency, the pressure of the
shell can be lowered to reduce the boiling temperature of the reactor and cold water can be fed
in. This will cause the water to boil at a lower temperature, cooling the reactor and stopping the
reaction. Once the cooling water is vaporized it can then be used as process steam in other
processes.
Heat transfer within the reactor is a major concern, during the optimization of the reactor
a variety of tube sizes were investigated to determine which tube size could provide an
economical surface area to volume ratio. A larger tube diameter would reduce the number of
tubes necessary to contain the desired amount of catalyst, but offers poor heat transfer from the
tube. A small tube diameter allows for greater heat transfer and temperature control however,
the amount of catalyst within a tube is limited and thus requires a larger number tubes. A 1.25
inch diameter tube was chosen because this offers a balance between heat transfer area and
21
catalyst volume. The larger heat transfer area allows the ambient temperature of the cooling
water to be higher and help facilitate the reaction when it slows down near the end of the reactor.
Due to the higher concentrations of carbon monoxide and hydrogen at the inlet of the reactor, the
reaction proceeds much quicker than at the end of the reactor where there are fewer reactants
present. This can be seen in Figure 14 in the Appendix A.5.
An ambient temperature of 405.1⁰F was chosen because it allows for a buffer region
before the reaction becomes a runaway reaction while maintaining a high enough temperature to
encourage the reaction to be completed within a reasonable amount of catalyst. This temperature
also keeps the reactor within the 390⁰F to 450⁰F operating range where the methane selectivity is
applicable. Since approximately 45% of the conversion occurred within the hotter portion of the
reactor and the remaining 45% of the conversion occurred in the cooler portion of the reactor, it
was assumed that the product distribution would not be affected by the fluctuations of
temperatures within the reactor. It was then acceptable to assume that the inlet temperature of
the reactor dictated the product distribution from the reactor. The temperature profile of the
reactor can be seen in figure 16 in Appendix A.5.
Typical conversions for Fischer-Tropsch range between 11-70% for a bench scale reactor
(Yates & Satterfield, 1991). The conversion within the full-scale reactor was truncated at
approximately 91% to account for any possible improvements in conversion that might be
associated with scale. However complete conversion is possible based on the reaction kinetics
however, this was avoided to ensure that the subsequent separations could handle possibly large
flows of carbon monoxide and hydrogen though the system. Additionally it was assumed that
the desired molar ratio of hydrogen to carbon monoxide was two however, in actuality it will
need to be slightly higher than two to account for the additional hydrogen molecules at the ends
of the alkanes. This would make hydrogen the limiting reactant for the Fischer-Tropsch reaction
and there would be unreacted carbon monoxide present in the product stream.
An operating pressure of 300 psig was chosen for the Fischer-Tropsch Reactor to
eliminate the cost of compressing the vapors from the syngas unit to a higher pressure. It was
determined that the operating pressure of the Fischer-Tropsch reactor did not have a major effect
on the conversion and that operating the synthesis gas unit at a lower pressure would be more
cost effective for the process. The pressure drop within the reactor is a major concern due to the
22
decrease in moles from the polymerization of the Fischer-Tropsch reaction and the condensation
of the larger hydrocarbons. The maximum allowable pressure drop was assumed to 50 psi per
reactor. This maximum pressure drop limited the length and the maximum vapor velocity
though the reactors. To minimize the pressure drop within the reactors, the reactors were
arranged in parallel to increase the surface area of the tubes and decrease the superficial gas
velocity though the tubes. Since there was only a 50 psi pressure drop though the reactor,
further pressurization was not necessary in the subsequent separations to separate the heavier
hydrocarbons from the lighter components.
The catalyst being used for this process is an ultra-stable cobalt based catalyst that has
low deactivation rates. This catalyst has a packed bulk density of 0.8 g/cm3 (49.94 blm/ft
3) with a
void fraction of 0.4. The catalyst is assumed to be replaced every 4 years of operation and costs
$10/lbm to purchase. The choice of a cobalt based catalyst improves the alkane selectivity of the
reactor because this catalyst is not very active in the water-gas shift reaction and is thus less
likely to convert water to carbon dioxide compared to iron based catalysts (Yates & Satterfield,
1991).
It was determined that the Fischer-Tropsch Reactor unit requires a total of 43 reactors
arranged in parallel to control the pressure drop within the reactor and achieve the desired
conversion. These reactors are 20 feet in diameter and are 41 feet in length. Each reactor
contains approximately 9,552, 1.25 inch, 16 gauge tubes arranged in an equilateral triangle
configuration (McCabe, Smith, & Harriott, 2005). This configuration allows for the maximum
number of tubes to be placed within a shell while maintaining the desired spacing of 1 inch in-
between tubes. The minimum shell wall thickness was designed based on the ASME Boiler and
Pressure Vessel Code (Peters, Timmerhaus, & West, 2003). It was assumed that the reactors
would have a 3mm corrosion allowance, a joint efficiency of 0.85 for spot examined, electric
resistance weld, and be made from 316 stainless with a maximum allowable working stress for
operation up to 343°C (Peters, Timmerhaus, & West, 2003). The minimum shell wall thickness
was calculated to be 4 inches (Peters, Timmerhaus, & West, 2003). ASME torispherical heads
were chosen for the ends of the reactor to reduce the costs of the reactor (Peters, Timmerhaus, &
West, 2003). Sample calculations for this are shown in Equations 10-12 in the Appendix A.6.
23
Stainless steel was chosen for the material of construction in order to prevent corrosion to
the system during operation. Even though a majority of the products produced during Fishcer-
Tropsch are n-alkanes there is potential to form reactive alkenes that might harm the reactor.
The reactor was priced using the weight of steel within the reactor (Yates & Satterfield, 1991).
The price of carbon steel was estimated using a correlation for the price of a pressure vessel,
however this correlation is only effective in a weight range from 400 to 50,000 kg (Peters,
Timmerhaus, & West, 2003). For these reactors the price of steel was assumed to be at the
higher end of this range with the price of steel for 50,000 kg of steel. The price of carbon steel
was then multiplied by a cost factor of three to account for the higher price of stainless steel
(Peters, Timmerhaus, & West, 2003). A pressure factor of 2.4 was also applied to the cost of the
reactor to account for the pressures of boiling water in the shell (Peters, Timmerhaus, & West,
2003). The final cost of a single reactor was determined to be $2.6 million dollars per reactor or
$112.9 million for the entire system of reactors.
Hydro-Isomerization Unit
The hydro-isomerization unit isomerizes paraffins and petroleum wax is converted into
LPG, naphtha, and diesel. The hydro-isomerization unit converts 100% of the greater than
700°F boiling point material to lighter components. This catalyst is very selective to the greater
than 700°F boiling point material and less is not selective to the less than 700°F boiling point
materials. However, the catalyst is sensitive to water and carbon monoxide and will deactivate
the catalyst. Prior to the hydro-isomerization unit all of the carbon monoxide and water is
removed from the petroleum wax. The hydro-isomerization unit has an overall selectivity of
1.0% wt% methane, 0.5wt% ethane, 3.5wt% propane, and the balance is diesel.
Product Separations
For simplicity the separation of this process was modeled assuming that all of the naphtha
and diesel behaved as a pseduocomponent within the separation. This assumption was made to
simplify the separation and because NRTL predicted that various hydrocarbons formed multiple
azeotropes with water. However, these azeotropes could not be independently verified in
literature and could have a drastic effect on the separation. Additionally, non-random two liquid
(NRTL) model was used to model the separation due to the large amount of water present in the
24
system. Prior to the construction of the proposed plant, a more rigorous investigation should be
completed to ensure that the separation will operate as intended.
Due to the lower operating temperature of the reactor, very little LPG was produced in
the reaction. The low quantity and value of LPG made it uneconomical to recover LPG from the
system due to the extreme condition necessary to separate LPG from methane, carbon monoxide
and hydrogen. All of the LPG produced by the plant comes from petroleum wax in the hydro-
isomerization unit. It is recommended that prior to the construction of the proposed plant that
further investigation of LPG recovery should be undertaken.
Diesel and Wax Separation
The heavier hydrocarbons that have boiling points greater than 350°F must be kept above
250°F prior to the hydro-isomerization unit to prevent wax crystallization within the equipment.
This fraction includes all of the C20+ hydrocarbons including diesel and petroleum wax. After
the Fischer-Tropsch reactor, the reactor effluent is immediately flashed at conditions similar to
the exiting reactor conditions to separate a majority of the heavier compounds. The reactor
effluent is adiabatically flashed at 408.5°F and 265.4 psia in a cluster of six flash drums. These
flash drums were sized assuming a 10 minute liquid hold up or the maximum vapor velocity to
determine the tank diameter. The height to diameter ratio was assumed to be 3. These equations
are shown in Equation 14 in the Appendix A.6 (Cheresources, 2008). Due to the large
volumetric flow rate in the system, many of the flash drum sizes needed to be broken into
smaller flash drums in order to have a realistic size. These flash drums were priced using the
price of steel similar to the sizing of the reactor. Again stainless steel was used to minimize
potential corrosion. A breakdown of the flash drum sizing and economics are shown in Table 28
in the Appendix A.3.
After this flash the liquid stream is combined with other heavy streams from other flash
drums in the vapor portion of the separation. The resulting stream is throttled down to 1 atm and
combined with more heavy products from the naphtha process streams. The resulting stream is
then preheated in a heat exchanger with the vapor stream from the first flash before is fed into a
vacuum distillation column. Steam ejectors are used in the vacuum distillation column to
maintain an operating pressure around 60mmHg. The use of a vacuum distillation allows the
waxy components to be separated from the diesel prior to being fed into the hydro-isomerization
25
unit without causing the waxy products to coke. This vacuum distillation completely separates
the diesel from the waxy components and increases the efficiency of the hydro-isomerization
unit so that only the waxy components are fed into the unit. The feed into the hydro-
isomerization unit is 100mol% petroleum wax. A process flow diagram for the diesel and wax
separation is shown in Figure 5 of Appendix A.1.
The first vacuum distillation column is 25 feet in diameter and has 14 stages with 2 foot
spacing in-between stages. Stainless steel was used for the construction material for the vacuum
distillation column to prevent possible corrosion. The reboiler is heated by a fired heater using
fuel gas in order to reach the temperatures necessary to separate the diesel from the waxy
components (Peters, Timmerhaus, & West, 2003). The condenser is cooled using cooling water
and was sized as a heat exchanger (Peters, Timmerhaus, & West, 2003). It was assumed that the
price of this distillation column would be three times as expensive as a distillation column
operating at atmospheric pressure. The price of this vacuum distillation column is $5.2 million.
A summary of the distillation column sizing and costing can be seen in Table 30 in the Appendix
A.3.
The distillate from the first vacuum distillation column contains primarily diesel and a
small amount of naphtha. Naphtha is separated from the diesel in a second vacuum distillation
column to increase the purity of the diesel and recover the naphtha. The naphtha in the distillate
is then pressurized to atmospheric pressure in a compressor and combined with naphtha from the
naphtha recovery process. Diesel from the second vacuum distillation is 100mol% diesel. This
vacuum distillation column is 25.8 feet in diameter and has 8 stages with 2 foot spacing in
between. The reboiler is heated using 600 psig steam from the plant while the condenser is
cooled using cooling water. This second vacuum distillation column costs $2.3 million to
purchase.
Naphtha Separation
The naphtha separation involves the removal the water and vapors from the naphtha.
After the Fischer-Tropsch reactor effluent is flashed, the vapor stream is used to preheat the feed
into the first vacuum distillation column and the feed to the sixth flash drum. This stream is then
further cooled to 365°F by vaporizing medium pressure steam condensate to produce steam to be
used in the plant. The vapors are then flashed again in a second adiabatic flash at 265.4 psia.
26
This second flash drum removes a portion of the residual diesel and is combined with the feed to
the vacuum distillation columns to be recovered. After this second flash, the vapor is cooled to
260°F in a heat exchanger with low pressure steam condensate to produce low pressure steam
and then flashed a third time in a decanter and flash drum to remove water from the vapor
stream.
The vapors from this flash drum are cooled with cooling water to 250°F and adiabatically
flashed and decanted in the fourth flash drum to recover any diesel that is present in the system.
The water from the decanter is combined with other waste water streams to be purified. Vapors
from the fourth flash drum are cooled to 95°F using cooling water and adiabatically flashed and
decanted in the fifth flash drum to separate the vapors from the waste water and naphtha. The
waste water phase is combined with other waste water steams to be purified while the naphtha is
refrigerated to 10°F and then used as an absorbent to remove naphtha from the vapor steam in an
absorber. This absorber recovers an additional 14,000 lbm per hour of naphtha that would
otherwise be used as fuel gas. The resulting naphtha stream is then decanted and mixed with
other naphtha streams. A process flow diagram for the naphtha separation is shown in Figure 6
of Appendix A.1.
The water recovered from the third flash drum is combined with other waste water
streams to be purified while the naphtha and diesel from the liquid stream is heated in a heat
exchanger with the vapors from the first flash to 381°F and then adiabatically flashed to
atmospheric pressure. This sixth flash drum separates the diesel from the naphtha and water.
The diesel is then fed into the diesel recovery and the naphtha vapors are then cooled in a heat
exchanger with cooling water to 235°F and adiabatically flashed in the seventh flash drum. The
liquid stream from this flash drum is fed into the diesel recovery and the naphtha vapors are
mixed with the other naphtha streams.
Naphtha from the diesel and wax separation is mixed with the naphtha from the seventh
flash drum and from the decanter. The resulting stream is then cooled in a heat exchanger with
water to 95°F and adiabatically flashed and decanted. This purifies the naphtha stream to
98.6mol% naphtha. The vapor stream from the flash drum still contains a small portion of
naphtha and is then compressed to 265.4 psia and recycled back into the absorber to recover the
naphtha.
27
Waste Water and Fuel Gas Purification
The vapor stream from the fifth flash drum contains primarily carbon monoxide,
hydrogen, methane, and C2-C4 hydrocarbons. This stream is then cooled using cooling water to
95°F and run through an absorber to remove a portion of the remaining naphtha. The scrubbed
vapors are then throttled to 1 atm and heated to 90°F using a heat exchanger with the waste water
stream. This stream is then fed into a stripper where the gas is used as a stripping agent to
remove volatile hydrocarbons from the waste water. The vapors leaving the stripper are burned
as a fuel gas to power a fired heater to produce steam or electricity.
Waste water produced in the plant is accumulated and then throttled to 1 atm. This water
is then cooled to 95°F in a heat exchanger with cooling water and fed into a stripper. The
stripper uses the fuel gas stream to removes a majority of the volatile hydrocarbons that might be
present in the waste water. It is important to remove these hydrocarbons before they are returned
to the plant’s cooling towers due to the possibility of these hydrocarbons vaporizing in the
cooling process. The resulting water stream is then fed into another stripper where air is blown
in using a blower to remove any remaining volatile hydrocarbons that might be present in the
water. This vapor stream is then used as air to be burned in a fired heater with the fuel gas. The
purified water is 99.4% pure water by volume and can be used as cooling water in the rest of the
plant.
Additional air maybe needed to power the fired heater because the concentration of fuel
gas to air is only 40%. The upper flammability limit of carbon monoxide and hydrogen are 75%
by volume and can be combusted using this volume ratio however, the upper flammability limits
for methane, ethane, propane, and butane are 15%, 12.4% 10.1% and 8.41% respectively. It is
important to ensure that all of the hydrocarbons are completely combusted to prevent any
possible hydrocarbon releases into the surrounding area.
Heat Integration
A majority of the heat produced by this process comes from the partial combustion of
oxygen in the synthesis gas production unit. This makes it difficult to find cold streams that need
to be heated within the process, so a majority of the heat integration in the system comes from
the production of steam that can be used in other processes. A summary of heat exchanger
equipment can be viewed in Tables 26 and 27 in Appendix A.3.
28
The feed to the synthesis gas production unit is preheated using a heat exchanger with the
effluent. This makes it so that a fired preheater is only needed during transient and startup
conditions. During normal operations the heat exchanger can provide all of the duty necessary to
preheat the feed.
Following the FTR, the reactor effluent is flashed in a flash drum where the vapor stream
is used to preheat the feed into the vacuum distillation column. This heat exchanger preheats the
feed to 410°F and partially vaporizes the feed. The heat exchanger reduces the duty on the
reboiler of the vacuum distillation column as well as the column diameter. After the vapor
stream is used to preheat the vacuum distillation column it is used to preheat the feed into the
sixth flash drum. This allows the subsequent flash to occur adiabatically and helps to separate
the naphtha from the diesel.
After the absorber the vapor stream is heated using the waste water stream. This heat
exchanger reduces the cooling water requirements to cool the waste water while also heating the
vapor stream. Heating the vapor stream makes the vapors a better scrubbing material for the
waste water and helps to vaporize the volatile hydrocarbons present in the waste water.
Partial condensers are used in both distillation columns to reduce the condenser duties
and raise the condenser temperatures on these columns. The lower condenser duties require less
cooling water and the condensers can be smaller due to the increased temperature difference.
Rather than separate the reactor effluent into multiple streams, the process is designed as one
separation unit to allow the distillation columns and other major pieces of equipment to be larger
to reduce the price of the equipment.
A majority of the heat integration was done in the form of steam generation thought out
the process. Following the syngas unit the synthesis gas needs to be cooled from 1950°F to
425°F. After the synthesis gas is used to preheat the feed, it is used to create high pressure and
low pressure steam in a series of heat exchangers. Following these heat exchangers an additional
1.9 million lbm of high pressure steam and 133.5 thousand lbm of medium pressure steam are
produced per hour. Additionally after the vapor stream from the first flash is used to preheat the
feeds into the vacuum distillation column and the sixth flash there is still enough heat to generate
35 thousand lbm of medium pressure steam per hour. Following the second flash drum, low
29
pressure steam condensate is used to cool the vapor steam to produce an additional 900 thousand
lbm of low pressure steam per hour.
Cooling water was also integrated into the system to reduce the size of the heat
exchangers used in to cool various streams. Cooling water is used to cool the streams that are
needed to be cooled to low temperatures first. After these streams are sufficiently cooled, the
same cooling water is used in later heat exchangers that allow for higher temperatures
differences in these heat exchangers. This increases the temperature difference in the low
temperature heat exchangers because of the much higher flow rate of cooling water and
decreases the surface area needed by these heat exchangers. The later heat exchangers have such
a large temperature difference already that the slight increase in the inlet temperature of the
cooling water does not affect the size of these heat exchangers.
Plant Economics
The expected service life for the plant is expected to be 15 years and has a total capital
investment of $1.02 billion. The total capital investment for the project was calculated by
multiplying the total equipment cost of $212 million by a Lang factor of 4.8. This Lang factor
accounts for any associated direct costs, indirect costs, and working capital. A summary of the
equipment costs can be viewed in Table 25 of Appendix A.2.
A stream factor of 0.92 was applied to the process to account for one month turn around
for catalyst replacement in the Fischer-Tropsch Reactors. The proposed plant design has the
potential of producing 14.6 million barrels of naphtha, 10.5 million barrels of diesel, and 4.2
million pounds of LPG per year. This amounts to an annual revenue of $2 billion with an
additional $879 million from credits. Variable costs including raw materials and utility costs for
the plant account for $1.8 billion per year. A breakdown of the plant economics is shown in
Table 19 of Appendix A.2.
Fixed costs for the plant include yearly operating expenses, depreciation, and the hydro-
isomerization unit and syngas unit capital recovery. The year operating expenses is expected to
be 3% of the total capital investment to cover fixed charges such as plant overhead costs,
administrative costs, distribution and marketing, and research and development. Depreciation
was calculated using straight line depreciation with no salvage value for the service life of the
30
plant. To cover the cost of purchasing the hydro-isomerization unit and syngas unit, the capital
recovery of this equipment was estimated to be $100 million and $400 million respectively. This
brought the total manufacturing costs to $2.4 billion. A summary of the utilities and credits for
various process units can be seen in Table 20 in Appendix A.2
A tax rate of 33% was applied to the gross profit to calculate a net profit of $365 million
per year. The cash flow was then calculated to be $433 million each year with a DCFRR of
42.2% and a cash flow payback period of 2.36 year. The plant offers a return on investment of
be 35.75% per year with a net present worth of the plant of $2.6 billion, assuming an 8% interest
rate. A summary of the Economics can be viewed in Table 24 of Appendix A.4
A sensitivity analysis for the plant was also calculated to determine the how the plant
economics would respond to variations in product price, capital investment, utility, and raw
material costs. It was determined that the product costs had the largest effect on the economics
of the plant. A decrease in produce price of 20% could severely hurt the profitability of the
plant. Utilities and raw materials accounted for a slight change in the profitability of the plant
but even with a 50% increase in the utility or raw material costs the plant will still be profitable.
The fixed capital remained relatively constant throughout the analysis. The sensitivity analysis
can be viewed in Figure 9 of Appendix A.2
Inflation was assumed to be a constant 3% per year for the lifespan of the plant. The
appreciation of the products and manufacturing costs helped improve the profitability of the
plant. However, inflation does not have an effect on the time value of money and does not affect
depreciation or the net present worth of the plant. A summary of the inflation can be viewed in
Table 22 in Appendix A.2.
31
Material and Energy Balance Discussion
Material and energy balances were performed though the development of the system to
ensure that the number of carbons and oxygen were constant though out the system. However in
the design of the Fischer-Tropsch reactor is was assumed that for every mole of carbon
consumed, two moles of hydrogen were consumed. This assumption is not accurate since in the
formation of alkanes two additional hydrogen molecules are necessary to complete the reaction.
To account this, a hydrogen balance was not used to measure the mass in this system. A
summary of the material and energy balances can be seen in the stream tables in Appendix A.1.
The carbon balances around the synthesis gas unit and the Fischer-Tropsch reactor close
to 99.15% while the oxygen balance closes to 100%. Also the carbon efficiency for the system
was calculated assuming that the composition of the naphtha and diesel streams within the
separation maintained the same composition as the FTR effluent. The carbon efficiency of the
system was determined by summing the mass of carbon in the finished products (LPG, Diesel,
and Naphtha) and dividing it by the mass of carbon in the methane feed. Though a carbon
balance the system’s carbon efficiency was determined to 91.3%. A majority of the carbon lost
from the system was burned in the fired heater as carbon monoxide, methane or LPG. The
separation process recovers 99.5% of the diesel and 98.8% of the naphtha produced in the FTR
and has a plant of 91%. If a higher conversion is possible, more liquid products could be
produce which would improve the carbon efficiency of the process. The material balances can
be viewed in Tables 43 and 44 in Appendix A.5 and a summary of the efficiency can be viewed
in Table 23 of Appendix A.2.
32
Safety and Environmental Summary
It is imperative that every precaution be taken to prevent possible leaks and ensure the
safe operation of the equipment and maintain the safety of employees, the community, and the
environment. This process involves many pieces of equipment that operate at high temperatures
and pressures, in the event of an emergency the plant should be immediately shut down and the
proper action should be taken. It is recommended that plant employees wear person protective
equipment when handling equipment and take the required safety training before working with
or near equipment.
The synthesis gas production unit has operates at a temperature of 1950°F and requires
special attention when performing maintenance and during normal operations. It should be
ensured that the area around the syngas unit is properly marked and cordoned off to prevent any
potential burns that might occur if an employee were to come into contact with the equipment.
The syngas effluent is cooled in a series of heat exchangers that should be properly insulated to
maintain energy efficiency and prevent potential hazards to employees.
The Fischer-Tropsch Reactor has the potential to cause a runaway reaction if the reactor
conditions are not properly maintained. Failsafe valves and close process controls are used to
maintain the reactor conditions and should be regularly inspected and maintained to ensure the
proper measurements are being made. Additionally, the shells and tubes of the reactor should be
inspected whenever necessary to clean away any potential buildup on the tubes that might disrupt
heat transfer of the reactor. In the event of an emergency, the pressure within the reactor should
be dropped to atmospheric pressure and cold cooling water should be used to quench the reaction
immediately. Employees should be trained in emergency and evacuation procedures.
The emission of volatile hydrocarbons should be minimized wherever possible. Methane
has the greenhouse gas potential of 21 times that of carbon monoxide while lighter hydro
carbons are known to produce smog (Agency, 2003). Additionally heavier hydrocarbons can
contaminate the soil and surrounding waterways if released into the environment (Agency,
2003). All plant vessels should have secondary containment in the event of a spill to prevent
endangering the environment.
33
This process also involves many chemicals that could be hazardous to employee’s health
and safety if not handled properly. A majority of the components within the system are
flammable and potentially explosive at certain conditions. Every precaution should be taken to
ensure that potential ignition sources are kept away from hazardous areas where explosions
could occur. Additionally, operations and employee offices should be located in safe locations
that would not be endangered in the case of emergency whenever possible. Listed below are
chemicals used in the system and the possible hazards they might present.
Methane
Methane is defined as an asphyxiant and personal protective equipment should be used in
case of a leak. In the event of a leak the area should be secured and immediately contact
emergency personnel. The flow of methane should be stopped as soon as possible to contain the
leak if it can be done safely. Methane forms explosive or flammable mixtures with oxygen and
other oxidizers. In case of fire water, carbon dioxide, or dry chemical should be used to
extinguish the blaze. (Material Safety Data Sheet: Methane)
Oxygen
Oxygen is an asphyxiant and an oxidizer. In the event of a release, it is important to
secure the area, use personal protective equipment as soon as possible and contact emergency
personnel. (Material Safety Data Sheet: Oxygen, 2011)
Carbon Dioxide
Carbon dioxide is an asphyxiant and personal protective equipment should be used in
case of a leak. In the event of a leak, the area should be secured and immediately contact
emergency personal. If possible the leak should be immediately contained if it is safe to do so.
(Material and Safety Data Sheet: Carbon Dioxide, 2005)
Carbon Monoxide
Carbon monoxide is extremely toxic and could be fatal if inhaled. Personal protection
equipment should be worn around areas containing carbon monoxide. Carbon monoxide is
extremely flammable and could cause flash fires. In the event of a leak the area should be
34
secured and immediately contact emergency personnel. In case of fire, water spray, foam or dry
chemical should be used to extinguish the blaze. (Material Safety Data Sheet: Carbon Monoxide,
2010)
Hydrogen
Hydrogen is an asphyxiant and personal protective equipment should be worn in the
event of a leak. Hydrogen is flammable over a wide range of concentrations in air and burns
with an almost invisible flame. In the event of fire use water, dry chemical, or carbon monoxide
to extinguish the blaze. (Material Safety Data Sheet: Hydrogen, 1996)
Ethane
Ethane is extremely flammable and personal protection equipment should be worn in
areas where exposure to ethane is possible. In the event of a leak the area should be secured and
immediately contact emergency personnel. In case of fire, water spray, foam or dry chemical
should be used to extinguish the blaze. (Material Safety Data Sheet: Ethane, 2010)
Propane
Propane is an asphyxiant and is extremely flammable. Personal protection equipment
should be worn in areas where exposure is possible. In the event of a leak the area should be
secured and immediately contact emergency personnel. In case of fire, water spray, foam or dry
chemical should be used to extinguish the blaze. (Material Safety Data Sheet: Propane, 2010)
Butane
Propane is an asphyxiant and is extremely flammable. Personal protection equipment
should be worn when handling butane. In the event of a leak the area should be secured and
immediately contact emergency personnel. In case of fire, water spray, foam or dry chemical
should be used to extinguish the blaze. Use proper disposal procedures in the event of a spill.
(Material Safety Data Sheet: N-Butane, 2010)
35
Naphtha
Naphtha is a known irritant and is extremely flammable. Personal protection equipment
should be worn when handling naphtha. In the event of a spill the area should be secured and
emergency personnel should be contacted. Avoid spreading the spill and use an inert absorbent
to absorb the spill. Use proper disposal procedures to contain the spill. In the event of a five,
use water, foam or dry chemical to extinguish the blaze (Material Safety Data Sheet: Naphtha,
2010).
Diesel
Diesel is a sever skin and eye irritant and is an asphyxiant. Personal protection
equipment should be worn when handling diesel. Diesel is also extremely flammable and should
be stored away from potential sources of ignition. In the event of a spill the area should be
secured and immediately contact emergency personnel. In case of fire, water spray, foam or dry
chemical should be used to extinguish the blaze. In the event of a spill, proper disposal and
cleanup procedures should be used (Material Safety Data Sheet: Diesel, 2002)
Petroleum Wax
Petroleum Wax can be a skin or eye irritant and can cause thermal burns if ignited. While not
considered “flammable” or “combustible” the material will burn if ignited. If a fire does occur,
use water spray, foam, carbon dioxide or dry chemical to extinguish the blasé. Use proper
disposal and cleanup procedures in the event of a spill. (Material Safety Data Sheet: Petroleum
Wax, 2007)
36
Unit Control and Instrumentation Description
Process control for the syngas unit is maintained by varying the flow rates of various
streams into the syngas unit to control the feed temperature, operating temperature, H2 to CO
ratio, conversion, and operating pressure. For these streams the fail safe for the control valves
would be fail close in order to cut off the fed into the syngas unit to stop the reaction.
Additional, block valves are placed before and after the syngas unit to close in the event of an
emergency. If pressure is building within the reactor a pressure relief valve is used to alleviate
the pressure back to the desired pressure.
The Fischer-Tropsch reactor is controlled primarily by using the shell side pressure to
change the boiling temperature of water within the shell. This method is used to control the
reactor temperature and conversion. For the inlet streams for the reactor have control valves
with a fail close failsafe. Block valves are also used to block off a reactor in case of emergency.
The control valve for the reactor pressure is designed to fail open in order to drop the pressure
within the reactor shell. This will cause the boiling temperature of water to decrease, decreasing
the temperature of the water within the shell, stopping the reaction. Additionally, the control
valve for the flow rate of water into the jacket has a fail open valve to ensure that in the event of
an emergency the reactor will still be contained in cooling water to prevent a runaway reaction.
The type of control primarily used in this system is PI Control. PI Control offers the
most responsive control type for these types of systems because it has no offset and low
oscillations. Pressure relief valves are also placed on a majority of the process equipment to
ensure that excessive pressure will not cause the unit to rupture and potentially harm employees.
A summary of the unit control and instrumentation can be viewed in Tables 35-38 in Appendix
A.3.
37
Conclusions
The proposed gas to liquid plant has the potential to economically bring to market remote
natural gas deposits by producing liquid fuels from natural gas. This plant design has the
potential of producing 14.6 million barrels of naphtha, 10.5 million barrels of diesel, and 4.2
million pounds of LPG per year with an annual revenue of $2 billion with an additional $879
million from credits. The product streams produce 100mol% petroleum wax to be converted
into desirable products in the hydro-isomerization unit and 100mol% diesel along with 98.6%
pure naphtha following the separation.
The plant’s efficiency was determined to 91.3% with a majority of the carbon lost from
the system being burned in the fired heater as carbon monoxide, methane or LPG. LPG is
burned with methane and any unreacted carbon monoxide and hydrogen in a fired heater to
ensure that no volatile hydrocarbons are released into the surroundings. The separation process
recovers 99.5% of the diesel and 98.8% of the naphtha produced in the FTR and has a plant of
91%. Waste water produced by the system is stripped of any volatile hydrocarbons that might
vaporize in the plant’s cooling towers to ensure that the process does not endanger the
environment.
Safety is a major concern in the design of the plant and proper safety training in the use
of personal protection equipment is mandatory. Control valves and proper instrumentation is
used to control the process and prevent the potential for emergencies. In the event of an
emergency the temperature of the cooling water within the shell can be reduced by lowering the
pressure within the shell and stopping the reaction.
Though the preliminary design it was determined that the process has a DCRFF of 42.2%
and return on investment of 2.36 years. Recent trends indicate that liquid fuels prices including
naphtha and diesel will continue to increase in coming years while natural gas prices will
decline. The production of liquid fuels from natural gas could offer a chance to produce higher
value products at remote locations. This plant has the potential to safely and economically bring
remote natural gas deposits to market as valuable liquid fuels.
38
Recommendations
The use of pseduocomponents in the separation is a board assumption and does not
account for possible interactions between these compounds. The NRTL thermodynamic model
predicts that there are multiple azeotropes within the system that could cause a potential problem
in separating the products leaving the FTR. Prior to the construction of the plant a more rigorous
investigation into the separation could be performed to ensure the viability of the plant.
Increasing the number of reactors in the plant could reduce the service factor. Currently
there are 43 reactors in series and one month of downtime is needed to coincide with catalyst
replacement. If additional reactors were added to the system, a section of reactors could be taken
out of service while the process is still in operation. However, prior to adding this to the design
it should be determined if it conditions will be safe to perform maintenance within the reactors
while the rest of the plant is in operation.
A major obstacle for this process is the incomplete conversion of the carbon monoxide
and hydrogen in the reactor. If the reactor can be designed to achieve near complete conversion
the capital and utility costs associated with the costly separation of products from the vapors
phase could be drastically reduced. Additionally if the amount of carbon monoxide and
hydrogen were reduced the recovery of LPG could become more economically viable and add an
additional source of revenue for the plant as well as increase the efficiency of the plant.
To improve the plant efficiency there are a few flash drums that require a large volume
due to the small amount of liquid recovered in these flash drums. The potential removal of these
flash drums could reduce the capital investment of the project, however the effects of the
removal of these flash drums on the subsequent separation and the environment.
39
Project Premises and Assumptions
Synthesis Gas Production Unit Assumptions
Oxygen fed into the syngas unit is completely combusted
Equilibrium constants for the steam reforming reaction and water-gas shift reaction did not
depend on temperature
The syngas unit would operate most efficiently at lower pressures
The desired hydrogen to carbon monoxide ratio was 2
The syngas unit operated isothermally
Fischer-Tropsch Reactor Assumptions
The 30% diesel and 70% wax would be fed into the hydro-isomerization unit to optimize the
reactor costs
The price of carbon steel was the same as the maximum weight in the range to price steel
All tubes within the reactor behaved the same as the tube being modeled
The vapor phase in the reactor contained all of the hydrogen, carbon monoxide, water,
methane, nitrogen, and 70% of the other components
The total pressure drop is 1.5 the calculated pressure drop to account for the condensation of
components within the reactor
Shell side of the reactor operated isothermally
Ambient shell temperature is constant
Inlet temperature of the reactor determined the product distribution
The stoicometric coefficient for hydrogen in the Fischer-Tropsch reaction was 2
HI Unit Assumptions
The less than 700°F boiling point material is not affected in the HI Unit
Separation Assumptions
Naphtha and diesel behave as pseduocomponent within the separation
Assumed a heat transfer coefficient for heat exchangers to be:
o Gas-Gas: 5.2833 Btu/ft2-hr-°F
o Gas-Water: 39.625 Btu/ft2-hr-°F
40
o Water-Water: 396.25 Btu/ft2-hr-°F
Reboilers and Condensers can sized and priced as heat exchangers
41
References
(n.d.). Retrieved 1 20, 2011, from Indiana University-Purdue University Indianapolis:
http://www.engr.iupui.edu/me/courses/shellandtube
Material Safety Data Sheet: Hydrogen. (1996, 6 7). Retrieved 2 11, 2011, from BOC Gases:
http://www.fergusongases.com/MSDS_files/HYDROGEN.pdf
Material and Safety Data Sheet: Carbon Dioxide. (2005, 4 11). Retrieved 2 11, 2011, from
Airgas: http://www.airgas.com/documents/pdf/1013.pdf
Wiley Critical Content - Petroleum Technology, Vol 1-2. (2007). John Wiley & Sons.
Cheresources. (2008). Retrieved 1 30, 2011, from
http://www.cheresources.com/questions/equipment_design-98.html
Material Safety Data Sheet: Carbon Monoxide. (2010, 9 2). Retrieved 2 11, 2011, from Airgas:
http://www.airgas.com/documents/pdf/001014.pdf
Material Safety Data Sheet: Ethane. (2010, 4 26). Retrieved 2 11, 2011, from Airgas:
http://www.airgas.com/documents/pdf/001024.pdf
Material Safety Data Sheet: Propane. (2010, 11 10). Retrieved 2 11, 2011, from Airgas:
http://www.airgas.com/documents/pdf/001045.pdf
Material Safety Data Sheet: Oxygen. (2011, 2 2). Retrieved 2 11, 2011, from Airgas:
http://www.airgas.com/documents/pdf/001043.pdf
Agency, U. S. (2003). Tool for the Reduction and Assessment of Chemical and Other
Environmental Impacts (TRACI). Washington D.C.: United States Environmental
Protection Agency.
Aspen Technology. (Vers. 2006). Aspen Plus.
Choi, G. N., Kramer, S. J., & Tam, S. S. (1996). Design/Economics of a Once-Trhough Natural
Gas Fischer-Tropsch Plant Weith Power Co-Production. Department of Energy.
Elbashir, N. O., & Roberts, C. B. (2004). Selective Control of Hydrocarbon Product Distribtuion
in Supercritical Phase Fischer-Tropsch Synthesis. Alabama: Auburn Univeristy
Department of Chemical Engineering.
Elliot, J. R., & Lira, C. T. (2009). Introductory Chemical Engineering Thermodynamics (Second
Edition). Upper Saddle River, NJ: Prentice Hall PTR, 2009.
Felder, R. M., & Rousseau, R. W. (2005). Elementary Principles of Chemical Processes (Third
Edition). Hoboken, NJ: John Wiley & Sons, Inc.
42
Fogler, S. H. (2008). Elements of Chemical Reaction Engineering (Fourth Edition). Westford
Massachusetts: Prentice Hall.
Gian Paolo, C., & Peter, M. M. (2006). Metal-Catalyst in Industrial Organic Processes. Royal
Society of Chmistry.
Haynes, W. M., & Lide, D. R. (2011). CNC Handbook of Chemistry and Physics (91st Edition).
Yalor and Francis Group.
Higman, C., & Maarten, v. d. (2008). Gasification (2nd Edition. Elsevier.
Letcher, T. M. (2008). Future Energy - Improved, Sustainable and Clean Options for Our
Planet. Elsevier.
Material Safety Data Sheet: Methane. (n.d.). Retrieved 2 11, 2010, from iSOC Technology:
http://www.isocinfo.com/DocumentRoot/13/Methane.pdf
McCabe, W. L., Smith, J. C., & Harriott, P. (2005). Unit Operations of Chemical Engineering
(Seventh Edition). New York, NY: McGraw-Hill.
National Institute of Standards and Technology. (2010, Febuary 1). NIST Chemistry WebBook.
Retrieved November 2010, from NIST Standard Reference Databse:
http://webbook.nist.gov/chemistry
Perry, R. H., Green, D. W., & Maloney, J. O. (1997). Perry's Chemical Engineers Handbook
(Seventh Edition). New York, NY: McGraw-Hill.
Peters, M. S., Timmerhaus, K. D., & West, E. R. (2003). Plant Design and Economics for
Chemical Engineers (Fifth Edition). New York, NY: McGraw-Hill.
Rackley, S. A. (2010). Carbon Capture and Storage. Elsevier.
Seader, J. D., & Henley, E. J. (2006). Separation Process Principles (Second Edition). Noboken,
NJ: John Wiley & Sons, Inc.
Singleton, A. H. (n.d.). Technology Development for Cobalt Fischer-Tropsch Catalyst. Pittsburn
PA: Energy international Corporation.
Storsaeter, S., Aaserud, C., Milmen, A.-.. M., Lindvag, O. A., Bergenes, E., Schankes, D., et al.
(2002). Effect of Water on the Fischer-Tropsch Synthesis on Supported Cobalt Catalyst.
Trondheim Norway: Norwegian University of Science and Technology Department of
Chemical Engineering.
Wang, X., & Michael, E. (2009). Advanced Natural Gas Engineering. Gulf Publishing
Company.
43
Yates, I., & Satterfield, C. (1991). Intrinsic Kinetics of the Fischer-Tropsch Synthesis on a
Cobalt Catalyst. Energy and Fuels Vol. 5, 168-173.
44
Synthesis Gas Unit
Fischer-Tropsch Reactor
Flash 1
Syngas Preheater
Vacuum Distillation 1
Vacuum Distillation 2
Absorber Stripper 1Stripper 2
Oxygen Preheater
HP Steam
MP Steam
Flash 6
Flash 7
Wax
Diesel
Compressor 1
Flash 2
Flash 3
Flash 4
Flash 5
Flash 8
Naphtha
Compressor 2
Refridgration
Fuel Gas
Blower
Waste Water
Fuel Air
DecanterMP Steam
Air
Hydro-Isomerization Unit
LPG
Naphtha
Diesel
LP Steam
Pump
MP Steam
Syngas Unit
Fischer-Tropsch Reactor Unit
Diesel and Wax Separation Unit
Naphtha Separation Unit
Appendix
Appendix A.1 – Process Flow Diagrams
Figure 1. Plant Process Flow Diagram
45
Appendix A.1.1 – Synthesis Gas Unit
High Pressure Steam Boiler
Medium Pressure Steam Boiler
Oxygen Preheater
Feed Preheater
Methane
Carbon Dioxide
Steam
Oxygen
Steam Condensate
Steam Condensate
Syngas Feed54,899.03 lbmol/hr Methane
27,449.52 lbmol/hr Water
17,784.69 lbmol/hr Carbon
Dioxide
Syngas Effluent277.52 lbmol/hr Methane
12,769.58 lbmol/hr Water
0 lbmol/hr Oxygen
10,44.72 lbmol/hr Carbon Dioxide
61,961.48 lbmol/hr Carbon Monoxide
123,922.95 lbmol/hr Hydrogen
164.65 lbmol/hr Nitrogen
Oxygen Feed16,3000.80 lbmol/hr Oxygen
164.65 lbmol/hr Nitrogen
Synthesis Gas Unit
Operating ConditionsTemperature: 1950
oF
Pressure: 300 psig
Steam to Methane Ratio: 0.5
Feed to Fischer-Tropsch Reactor277.52 lbmol/hr Methane
12,769.58 lbmol/hr Water
0 lbmol/hr Oxygen
10,44.72 lbmol/hr Carbon Dioxide
61,961.48 lbmol/hr Carbon Monoxide
123,922.95 lbmol/hr Hydrogen
164.65 lbmol/hr Nitrogen
Temperature: 425oF
Pressure: 300 psig
High Pressure Steam to
Other Processes1,878.99 klbm/hr
Temperature: 490oF
Pressure: 600 psig
Medium Pressure Steam
to Other Processes133.49 klbm/hr
Temperature: 353oF
Pressure: 125 psig
Figure 2. Synthesis Gas Unit PFD
46
Figure 3. ASPEN Diagram for the Syngas Unit Heat Exchangers
47
Table 2. Syngas Effluent and Feed Streams 1-DIOXIDE
1 2 3 4 5 DIOXIDE
Methane 277.52 277.52 277.52 - 54,899.03 -
Water 12,769.58 12,769.58 12,769.58 - 27,449.51 -
Oxygen - - - 16,300.80 - -
Carbon Dioxide 10,444.72 10,444.72 10,444.72 - 17,784.69 17,784.69
Carbon Monoxide 61,961.48 61,961.48 61,961.48 - - -
Hydrogen 123,923.00 123,923.00 123,923.00 - - -
Nitrogen 164.65 164.65 164.65 164.65 - -
Ethane - - - - - -
Propane - - - - - -
Butane - - - - - -
Naptha - - - - - -
Diesel - - - - - -
C21-C25 - - - - - -
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
Methane 0.0013 0.0013 0.0013 - 0.5483 -
Water 0.0609 0.0609 0.0609 - 0.2741 -
Oxygen - - - 0.9900 - -
Carbon Dioxide 0.0498 0.0498 0.0498 - 0.1776 1.0000
Carbon Monoxide 0.2957 0.2957 0.2957 - - -
Hydrogen 0.5914 0.5914 0.5914 - - -
Nitrogen 0.0008 0.0008 0.0008 0.0100 - -
Ethane - - - - - -
Propane - - - - - -
Butane - - - - - -
Naptha - - - - - -
Diesel - - - - - -
C21-C25 - - - - - -
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
209,541.00 209,541.00 209,541.00 16,465.45 100,133.00 17,784.69
2,684,160.00 2,684,160.00 2,684,160.00 526,219.00 2,157,940.00 782,701.00
16,745,100.00 12,988,700.00 6,860,740.00 501,111.00 1,527,870.00 207,532.00
1,883.77 1,358.07 500.47 1,000.00 336.32 100.00
314.70 314.70 314.70 514.70 514.70 514.70
Phase
Vapor Fraction 1.00 1.00 1.00 1.00 0.92 1.00
Liquid Fraction - - - - 0.08 -
-15,036 -19,284 -25,805 6,961 -75,036 -168,970
-1,174 -1,505 -2,014 218 -3,482 -3,839
-3.1506.E+09 -4.0408.E+09 -5.4071.E+09 1.1462.E+08 -7.5136.E+09 -3.0051.E+09
12.66 10.61 5.77 0.50 -16.45 -6.00
0.99 0.83 0.45 0.02 -0.76 -0.14
0.01 0.02 0.03 0.03 0.07 0.09
0.16 0.21 0.39 1.05 1.41 3.77
12.81 12.81 12.81 31.96 21.55 44.01
172,505.00 172,505.00 172,505.00 14,125.94 70,292.92 15,257.73
Entropy (Btu/lbm-R)
Density (lbmol/ft3)
Density (lbm/ft3)
Average Molecular Weight
Liquid Volume, 60°F (lbm/ft3)
Temperature (°F)
Pressure (psia)
Enthalpy (Btu/lbmol)
Enthalpy (Btu/lbm)
Enthalpy (Btu/hr)
Entropy (Btu/lbmol-R)
Stream Number
Molar Flow Rate (lbmol/hr)
Mole Fraction
Total Molar Flow Rate (lbmol/hr)
Total Mass Flow Rate (lbm/hr)
Total Volumetric Flow Rate (ft3/hr)
48
Table 3. Syngas Effluent and Feed Streams FTR-MP2OUT
FTR HP1IN HP1OUT METHANE MP2IN MP2OUT
Methane 277.52 - - 54,899.03 - -
Water 12,769.58 104,300.00 104,300.00 - 7,410.00 7,410.00
Oxygen - - - - - -
Carbon Dioxide 10,444.72 - - - - -
Carbon Monoxide 61,961.48 - - - - -
Hydrogen 123,923.00 - - - - -
Nitrogen 164.65 - - - - -
Ethane - - - - - -
Propane - - - - - -
Butane - - - - - -
Naptha - - - - - -
Diesel - - - - - -
C21-C25 - - - - - -
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
Methane 0.0013 - - 1.0000 - -
Water 0.0609 1.0000 1.0000 - 1.0000 1.0000
Oxygen - - - - - -
Carbon Dioxide 0.0498 - - - - -
Carbon Monoxide 0.2957 - - - - -
Hydrogen 0.5914 - - - - -
Nitrogen 0.0008 - - - - -
Ethane - - - - - -
Propane - - - - - -
Butane - - - - - -
Naptha - - - - - -
Diesel - - - - - -
C21-C25 - - - - - -
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
209,541.00 104,300.00 104,300.00 54,899.03 7,410.00 7,410.00
2,684,160.00 1,878,990.00 1,878,990.00 880,732.00 133,493.00 133,493.00
6,321,430.00 41,344.02 1,729,230.00 640,623.00 2,577.40 461,149.00
425.00 488.96 490.00 100.00 353.04 353.03
314.70 614.70 614.70 514.70 139.70 139.70
Phase
Vapor Fraction 1.00 - 1.00 1.00 - 1.00
Liquid Fraction - 1.00 - - 1.00 0.00
-26,357 -113,670 -100,570 -31,840 -117,390 -101,760
-2,058 -6,309 -5,582 -1,985 -6,516 -5,649
-5.5229.E+09 -1.1860.E+10 -1.0490.E+10 -1.7480.E+09 -8.6986.E+08 -7.5408.E+08
5.17 -27.14 -13.33 -25.95 -30.98 -11.75
0.40 -1.51 -0.74 -1.62 -1.72 -0.65
0.03 2.52 0.06 0.09 2.87 0.02
0.42 45.45 1.09 1.37 51.79 0.29
12.81 18.02 18.02 16.04 18.02 18.02
172,505.00 30,156.60 30,156.60 47,098.63 2,142.48 2,142.48
Entropy (Btu/lbm-R)
Density (lbmol/ft3)
Density (lbm/ft3)
Average Molecular Weight
Liquid Volume, 60°F (lbm/ft3)
Temperature (°F)
Pressure (psia)
Enthalpy (Btu/lbmol)
Enthalpy (Btu/lbm)
Enthalpy (Btu/hr)
Entropy (Btu/lbmol-R)
Stream Number
Molar Flow Rate (lbmol/hr)
Mole Fraction
Total Molar Flow Rate (lbmol/hr)
Total Mass Flow Rate (lbm/hr)
Total Volumetric Flow Rate (ft3/hr)
49
Table 4. Syngas Effluent and Feed Streams O2TOSYN-TOSYN
O2TOSYN OXYGEN STEAM SYNFEED SYNGAS TOSYN
Methane - - - 54,899.03 277.52 54,899.03
Water - - 27,449.51 27,449.51 12,769.58 27,449.51
Oxygen 16,300.80 16,300.80 - - - -
Carbon Dioxide - - - 17,784.69 10,444.72 17,784.69
Carbon Monoxide - - - - 61,961.48 -
Hydrogen - - - - 123,923.00 -
Nitrogen 164.65 164.65 - - 164.65 -
Ethane - - - - - -
Propane - - - - - -
Butane - - - - - -
Naptha - - - - - -
Diesel - - - - - -
C21-C25 - - - - - -
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
Methane - - - 0.5483 0.0013 0.5483
Water - - 1.0000 0.2741 0.0609 0.2741
Oxygen 0.9900 0.9900 - - - -
Carbon Dioxide - - - 0.1776 0.0498 0.1776
Carbon Monoxide - - - - 0.2957 -
Hydrogen - - - - 0.5914 -
Nitrogen 0.0100 0.0100 - - 0.0008 -
Ethane - - - - - -
Propane - - - - - -
Butane - - - - - -
Naptha - - - - - -
Diesel - - - - - -
C21-C25 - - - - - -
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
16,465.45 16,465.45 27,449.51 100,133.00 209,541.00 100,133.00
526,219.00 526,219.00 494,511.00 2,157,940.00 2,684,160.00 2,157,940.00
819,585.00 183,555.00 455,097.00 2,456,650.00 17,218,400.00 4,984,230.00
1,000.00 75.00 490.00 306.08 1,950.00 1,000.00
314.70 514.70 614.70 314.70 314.70 314.70
Phase
Vapor Fraction 1.00 1.00 1.00 0.94 1.00 1.00
Liquid Fraction - - - 0.06 - -
6,961 -14 -100,570 -75,036 -14,488 -66,146
218 0 -5,582 -3,482 -1,131 -3,069
1.1462.E+08 -2.3079.E+05 -2.7605.E+09 -7.5136.E+09 -3.0358.E+09 -6.6234.E+09
1.48 -6.98 -13.33 -15.54 12.89 -7.04
0.05 -0.22 -0.74 -0.72 1.01 -0.33
0.02 0.09 0.06 0.04 0.01 0.02
0.64 2.87 1.09 0.88 0.16 0.43
31.96 31.96 18.02 21.55 12.81 21.55
14,125.94 14,125.94 7,936.57 70,292.92 172,505.00 70,292.92
Entropy (Btu/lbm-R)
Density (lbmol/ft3)
Density (lbm/ft3)
Average Molecular Weight
Liquid Volume, 60°F (lbm/ft3)
Temperature (°F)
Pressure (psia)
Enthalpy (Btu/lbmol)
Enthalpy (Btu/lbm)
Enthalpy (Btu/hr)
Entropy (Btu/lbmol-R)
Stream Number
Molar Flow Rate (lbmol/hr)
Mole Fraction
Total Molar Flow Rate (lbmol/hr)
Total Mass Flow Rate (lbm/hr)
Total Volumetric Flow Rate (ft3/hr)
50
Appendix A.1.2 – Fischer-Tropsch Reactor Unit
Fischer-Tropsch Reactor Effluent4918.7 lbmol/hr Methane
69138.23 lbmol/hr Water
0 lbmol/hr Oxygen
10444.72 lbmol/hr Carbon Dioxide
5592.83 lbmol/hr Carbon Monoxide
11185.65 lbmol/hr Hydrogen
164.65 lbmol/hr Nitrogen
98.37 lbmol/hr Ethane
65.58 lbmol/hr Propane
49.19 lbmol/hr Butane
3184.93 lbmol/hr Naphtha
1245.26 lbmol/hr Diesel
208.58 lbmol/hr C21-C25
91.22 lbmol/hr C26-C29
73.83 lbmol/hr C30-C35
53.43 lbmol/hr C36-C47
12.79 lbmol/hr C48+
Temperature: 408.5oF
Pressure: 265.41 psia
Pump
Fischer-Tropsch Reactor
Operating ConditionsTemperature: 425
oF
Pressure: 300 psig
Shell Side Cooling Water11,570.3 klbm/hr
Temperature: 405.1oF
Pressure: 246.71 psig
Steam to Other Processes11,570.3 klbm/hr
Temperature: 405.1oF
Pressure: 246.71 psig
277.52 lbmol/hr Methane
12,769.58 lbmol/hr Water
0 lbmol/hr Oxygen
10,44.72 lbmol/hr Carbon Dioxide
61,961.48 lbmol/hr Carbon Monoxide
123,922.95 lbmol/hr Hydrogen
164.65 lbmol/hr Nitrogen
Figure 4. Fischer-Tropsch Reactor Unit PFD
51
Flash 1
Heat Exchanger 1
Fischer-Tropsch Reactor Effluent4918.7 lbmol/hr Methane
69138.23 lbmol/hr Water
10444.72 lbmol/hr Carbon Dioxide
5592.83 lbmol/hr Carbon Monoxide
11185.65 lbmol/hr Hydrogen
164.65 lbmol/hr Nitrogen
98.37 lbmol/hr Ethane
65.58 lbmol/hr Propane
49.19 lbmol/hr Butane
3184.93 lbmol/hr Naphtha
1245.26 lbmol/hr Diesel
208.58 lbmol/hr C21-C25
91.22 lbmol/hr C26-C29
73.83 lbmol/hr C30-C35
53.43 lbmol/hr C36-C47
12.79 lbmol/hr C48+
Temperature: 408.5oF
Pressure: 265.41 psia
Compressor 1
Heat Exchanger 16
To Flash 2
Flash 6
Heat Exchanger 10
Flash 7
Hydro-Isomerization Unit
LPG
Naphtha
Diesel
Flash 8
Vacuum Distillation 1
Vacuum Distillation 2
Vacuum Distillation 1 Feed1.71 lbmol/hr Methane
313.46 lbmol/hr Water
6.42 lbmol/hr Carbon Dioxide
1.15 lbmol/hr Carbon Monoxide
0.39 lbmol/hr Hydrogen
0.03 lbmol/hr Nitrogen
0.11lbmol/hr Ethane
0.2 lbmol/hr Propane
0.35 lbmol/hr Butane
327.97 lbmol/hr Naptha
1,239.78 lbmol/hr Diesel
208.48 lbmol/hr C21-C25
91.22 lbmol/hr C26-C29
73.83 lbmol/hr C30-C35
53.43 lbmol/hr C36-C47
12.79 lbmol/hr C48+
Liquid from Flash 60 lbmol/hr Methane
0.46 lbmol/hr Water
0.02 lbmol/hr Carbon Dioxide
0 lbmol/hr Carbon Monoxide
0 lbmol/hr Hydrogen
0 lbmol/hr Nitrogen
0 lbmol/hr Ethane
0 lbmol/hr Propane
0.1 lbmol/hr Butane
136.03 lbmol/hr Naptha
332.21 lbmol/hr Diesel
0.13 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Vapor from Flash 63.45 lbmol/hr Methane
69.15 lbmol/hr Water
17.39 lbmol/hr Carbon Dioxide
2.11 lbmol/hr Carbon Monoxide
0.54 lbmol/hr Hydrogen
0.06 lbmol/hr Nitrogen
0.28 lbmol/hr Ethane
0.44 lbmol/hr Propane
0.59 lbmol/hr Butane
586.7 lbmol/hr Naptha
13.11 lbmol/hr Diesel
0 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Liquid from Flash 70 lbmol/hr Methane
0.09 lbmol/hr Water
0.01 lbmol/hr Carbon Dioxide
0 lbmol/hr Carbon Monoxide
0 lbmol/hr Hydrogen
0 lbmol/hr Nitrogen
0 lbmol/hr Ethane
0 lbmol/hr Propane
0 lbmol/hr Butane
44.26 lbmol/hr Naptha
12.58 lbmol/hr Diesel
0 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Vapor from Flash 73.45 lbmol/hr Methane
69.05 lbmol/hr Water
17.38 lbmol/hr Carbon Dioxide
2.11 lbmol/hr Carbon Monoxide
0.54 lbmol/hr Hydrogen
0.06 lbmol/hr Nitrogen
0.28 lbmol/hr Ethane
0.44 lbmol/hr Propane
0.58 lbmol/hr Butane
542.45 lbmol/hr Naptha
0.53 lbmol/hr Diesel
0 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Liquid from Flash 11.03 lbmol/hr Methane
123.16 lbmol/hr Water
3.7 lbmol/hr Carbon Dioxide
0.71 lbmol/hr Carbon Monoxide
0.24 lbmol/hr Hydrogen
0.02 lbmol/hr Nitrogen
0.07 lbmol/hr Ethane
0.11 lbmol/hr Propane
0.19 lbmol/hr Butane
112.15 lbmol/hr Naptha
688.39 lbmol/hr Diesel
202.69 lbmol/hr C21-C25
90.82 lbmol/hr C26-C29
73.76 lbmol/hr C30-C35
53.43 lbmol/hr C36-C47
12.79 lbmol/hr C48+
Vapor from Flash 14,917.67 lbmol/hr Methane
69,015.07 lbmol/hr Water
10,441.02 lbmol/hr Carbon Dioxide
5,592.12 lbmol/hr Carbon Monoxide
11,185.41 lbmol/hr Hydrogen
164.63 lbmol/hr Nitrogen
98.3 lbmol/hr Ethane
65.47 lbmol/hr Propane
49.0 lbmol/hr Butane
3,072.78 lbmol/hr Naptha
556.87 lbmol/hr Diesel
5.89 lbmol/hr C21-C25
0.4 lbmol/hr C26-C29
0.07 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Liquid from Flash 40.01 lbmol/hr Methane
0.23 lbmol/hr Water
0.06 lbmol/hr Carbon Dioxide
0.01 lbmol/hr Carbon Monoxide
0 lbmol/hr Hydrogen
0 lbmol/hr Nitrogen
0 lbmol/hr Ethane
0 lbmol/hr Propane
0.1 lbmol/hr Butane
2.57 lbmol/hr Naptha
1.06 lbmol/hr Diesel
0 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Liquid from Flash 20.66 lbmol/hr Methane
189.5 lbmol/hr Water
2.63 lbmol/hr Carbon Dioxide
0.44 lbmol/hr Carbon Monoxide
0.14 lbmol/hr Hydrogen
0.01 lbmol/hr Nitrogen
0.05 lbmol/hr Ethane
0.08 lbmol/hr Propane
0.14 lbmol/hr Butane
32.97 lbmol/hr Naptha
205.55 lbmol/hr Diesel
5.67 lbmol/hr C21-C25
0.4 lbmol/hr C26-C29
0.07 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Liquid from Flash 33.45 lbmol/hr Methane
69.6 lbmol/hr Water
17.41 lbmol/hr Carbon Dioxide
2.11 lbmol/hr Carbon Monoxide
0.54 lbmol/hr Hydrogen
0.06 lbmol/hr Nitrogen
0.28 lbmol/hr Ethane
0.45 lbmol/hr Propane
0.6 lbmol/hr Butane
722.74 lbmol/hr Naptha
345.32 lbmol/hr Diesel
0.13 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Waxes to HI0 lbmol/hr Naptha
0.17 lbmol/hr Diesel
208.48 lbmol/hr C21-C25
91.22 lbmol/hr C26-C29
73.83 lbmol/hr C30-C35
53.43 lbmol/hr C36-C47
12.79 lbmol/hr C48+
Diesel0 lbmol/hr Naptha
0.17 lbmol/hr Diesel
208.48 lbmol/hr C21-C25
91.22 lbmol/hr C26-C29
73.83 lbmol/hr C30-C35
53.43 lbmol/hr C36-C47
12.79 lbmol/hr C48+
Naphtha0.44 lbmol/hr Methane
9.38 lbmol/hr Water
18.01 lbmol/hr Carbon Dioxide
0.1 lbmol/hr Carbon Monoxide
0 lbmol/hr Hydrogen
0 lbmol/hr Nitrogen
0.39 lbmol/hr Ethane
2.38 lbmol/hr Propane
6.62 lbmol/hr Butane
3,146.7 lbmol/hr Naptha
6.16 lbmol/hr Diesel
0 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Vacuum Distillation 1 Distillate1.71 lbmol/hr Methane
313.46 lbmol/hr Water
6.42 lbmol/hr Carbon Dioxide
1.15 lbmol/hr Carbon Monoxide
0.39 lbmol/hr Hydrogen
0.03 lbmol/hr Nitrogen
0.11lbmol/hr Ethane
0.2 lbmol/hr Propane
0.35 lbmol/hr Butane
327.97 lbmol/hr Naptha
1,239.78 lbmol/hr Diesel
0.01 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Vacuum Distillation 2 Distillate1.71 lbmol/hr Methane
313.46 lbmol/hr Water
6.42 lbmol/hr Carbon Dioxide
1.15 lbmol/hr Carbon Monoxide
0.39 lbmol/hr Hydrogen
0.03 lbmol/hr Nitrogen
0.11lbmol/hr Ethane
0.2 lbmol/hr Propane
0.35 lbmol/hr Butane
327.97 lbmol/hr Naptha
0.74 lbmol/hr Diesel
0 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
From Decanter27.37 lbmol/hr Methane
20.47 lbmol/hr Water
292.12 lbmol/hr Carbon Dioxide
12.99 lbmol/hr Carbon Monoxide
1.76 lbmol/hr Hydrogen
0.36 lbmol/hr Nitrogen
3.99 lbmol/hr Ethane
8.11 lbmol/hr Propane
11.33 lbmol/hr Butane
2,342.61 lbmol/hr Naptha
4.89 lbmol/hr Diesel
0 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+To Waste Water0.04 lbmol/hr Methane
330.37 lbmol/hr Water
1.81 lbmol/hr Carbon Dioxide
0.01 lbmol/hr Carbon Monoxide
0 lbmol/hr Hydrogen
0 lbmol/hr Nitrogen
0.04 lbmol/hr Ethane
0.24 lbmol/hr Propane
0.66 lbmol/hr Butane
0 lbmol/hr Naptha
To Absorber32.04 lbmol/hr Methane
63.24 lbmol/hr Water
296.1 lbmol/hr Carbon Dioxide
16.14 lbmol/hr Carbon Monoxide
2.68 lbmol/hr Hydrogen
0.45 lbmol/hr Nitrogen
3.95 lbmol/hr Ethane
6.13 lbmol/hr Propane
4.97 lbmol/hr Butane
66.32 lbmol/hr Naptha
0 lbmol/hr Diesel
Cooling Water 1 from
Heat Exchanger 3Temperature: 119.3
oF
2,042.96 kgal/hr
Exiting Cooling Water 1Temperature: 120
oF
2,042.96 kgal/hr
Heat Exchanger 14
Cooling Water 3Temperature: 85
oF
68.49 kgal/hr
Exiting Cooling Water 3Temperature: 120
oF
68.49 kgal/hr
Appendix A.1.3 – Diesel and Wax Separation Unit
Figure 5. Diesel and Wax Separation Unit PFD
52
Absorber
Stripper 1
Stripper 2
Heat Exchanger 9
Flash 2
Flash 3
Heat Exchanger 8
Flash 4Heat Exchanger 6
Flash 5
Compressor 2
Blower
Decanter
Steam Condensate
Steam Condensate
From Flsh 832.04 lbmol/hr Methane
63.24 lbmol/hr Water
296.1 lbmol/hr Carbon Dioxide
16.14 lbmol/hr Carbon Monoxide
2.68 lbmol/hr Hydrogen
0.45 lbmol/hr Nitrogen
3.95 lbmol/hr Ethane
6.13 lbmol/hr Propane
4.97 lbmol/hr Butane
66.32 lbmol/hr Naptha
0 lbmol/hr Diesel
To Flash 827.37 lbmol/hr Methane
20.47 lbmol/hr Water
292.12 lbmol/hr Carbon Dioxide
12.99 lbmol/hr Carbon Monoxide
1.76 lbmol/hr Hydrogen
0.36 lbmol/hr Nitrogen
3.99 lbmol/hr Ethane
8.11 lbmol/hr Propane
11.33 lbmol/hr Butane
2,342.61 lbmol/hr Naptha
4.89 lbmol/hr Diesel
0 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
To Vacuum Distillation 10.66 lbmol/hr Methane
189.5 lbmol/hr Water
2.63 lbmol/hr Carbon Dioxide
0.44 lbmol/hr Carbon Monoxide
0.14 lbmol/hr Hydrogen
0.01 lbmol/hr Nitrogen
0.05 lbmol/hr Ethane
0.08 lbmol/hr Propane
0.14 lbmol/hr Butane
32.97 lbmol/hr Naptha
205.55 lbmol/hr Diesel
5.67 lbmol/hr C21-C25
0.4 lbmol/hr C26-C29
0.07 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
To Flash 63.45 lbmol/hr Methane
69.6 lbmol/hr Water
17.41 lbmol/hr Carbon Dioxide
2.11 lbmol/hr Carbon Monoxide
0.54 lbmol/hr Hydrogen
0.06 lbmol/hr Nitrogen
0.28 lbmol/hr Ethane
0.45 lbmol/hr Propane
0.6 lbmol/hr Butane
722.74 lbmol/hr Naptha
345.32 lbmol/hr Diesel
0.13 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Heat Exchanger 7
Vapor from Flash 14,917.67 lbmol/hr Methane
69,015.07 lbmol/hr Water
10,441.02 lbmol/hr Carbon Dioxide
5,592.12 lbmol/hr Carbon Monoxide
11,185.41 lbmol/hr Hydrogen
164.63 lbmol/hr Nitrogen
98.3 lbmol/hr Ethane
65.47 lbmol/hr Propane
49.0 lbmol/hr Butane
3,072.78 lbmol/hr Naptha
556.87 lbmol/hr Diesel
5.89 lbmol/hr C21-C25
0.4 lbmol/hr C26-C29
0.07 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Vapor From Flash 24,917.0 lbmol/hr Methane
68,825.57 lbmol/hr Water
10,438.39 lbmol/hr Carbon Dioxide
5,591.68 lbmol/hr Carbon Monoxide
11,185.27 lbmol/hr Hydrogen
164.62 lbmol/hr Nitrogen
98.26 lbmol/hr Ethane
65.39 lbmol/hr Propane
48.86 lbmol/hr Butane
3,039.82 lbmol/hr Naptha
351.33 lbmol/hr Diesel
0.23 lbmol/hr C21-C25
0.0 lbmol/hr C26-C29
0.0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Medium Pressure Steam
to Other Processes35.13 klbm/hr
Temperature: 353oF
Pressure: 125 psig
Low Pressure Steam
to Other Processes900.76 klbm/hr
Temperature: 260oF
Pressure: 20 psig
Exiting Cooling Water 22,042.96 kgal/hr
Temperature: 120oF
Cooling Water 2 from
Heat Exchanger 62,042.96 kgal/hr
Temperature: 98.37oF
P-44
To Vacuum Distillation 10.01 lbmol/hr Methane
0.23 lbmol/hr Water
0.06 lbmol/hr Carbon Dioxide
0.01 lbmol/hr Carbon Monoxide
0 lbmol/hr Hydrogen
0 lbmol/hr Nitrogen
0 lbmol/hr Ethane
0 lbmol/hr Propane
0.1 lbmol/hr Butane
2.57 lbmol/hr Naptha
1.06 lbmol/hr Diesel
0 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Water Phase0.98 lbmol/hr Methane
88.07 lbmol/hr Water
10.48 lbmol/hr Carbon Dioxide
0.47 lbmol/hr Carbon Monoxide
0.06 lbmol/hr Hydrogen
0.01 lbmol/hr Nitrogen
0.14 lbmol/hr Ethane
0.29 lbmol/hr Propane
0.41 lbmol/hr Butane
0.2 lbmol/hr Naptha
0 lbmol/hr Diesel
Flash 3 Water Phase148.22 lbmol/hr Methane
63,080.66 lbmol/hr Water
748.25 lbmol/hr Carbon Dioxide
90.51 lbmol/hr Carbon Monoxide
23.07 lbmol/hr Hydrogen
2.57 lbmol/hr Nitrogen
11.89 lbmol/hr Ethane
19.19 lbmol/hr Propane
25.66 lbmol/hr Butane
0.53 lbmol/hr Naptha
0.05 lbmol/hr Diesel
0.1 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Flash 3 Vapor3.45 lbmol/hr Methane
69.6 lbmol/hr Water
17.41 lbmol/hr Carbon Dioxide
2.11 lbmol/hr Carbon Monoxide
0.54 lbmol/hr Hydrogen
0.06 lbmol/hr Nitrogen
0.28 lbmol/hr Ethane
0.45 lbmol/hr Propane
0.6 lbmol/hr Butane
722.74 lbmol/hr Naptha
345.32 lbmol/hr Diesel
0.13 lbmol/hr C21-C25
0 lbmol/hr C26-C29
0 lbmol/hr C30-C35
0 lbmol/hr C36-C47
0 lbmol/hr C48+
Cooling Water 2 to
Heat Exchanger 72,042.96 kgal/hr
Temperature: 98.37oF
Cooling Water 2 from
Heat Exchanger 52,042.96 kgal/hr
Temperature: 98.23oF
Flash 4 Water Phase0.8 lbmol/hr Methane
334.44 lbmol/hr Water
4.09 lbmol/hr Carbon Dioxide
0.49 lbmol/hr Carbon Monoxide
0.12 lbmol/hr Hydrogen
0.01 lbmol/hr Nitrogen
0.06 lbmol/hr Ethane
0.11 lbmol/hr Propane
0.14 lbmol/hr Butane
0 lbmol/hr Naptha
Flash 4 Vapor Phase4,764.52 lbmol/hr Methane
5,340.65 lbmol/hr Water
9,668.58 lbmol/hr Carbon Dioxide
5,498.58 lbmol/hr Carbon Monoxide
11,161.53 lbmol/hr Hydrogen
161.97 lbmol/hr Nitrogen
86.03 lbmol/hr Ethane
45.64 lbmol/hr Propane
22.46 lbmol/hr Butane
2,313.98 lbmol/hr Naptha
4.9 lbmol/hr Diesel
0 lbmol/hr C21-C25Heat Exchanger 5
Cooling Water 2 to
Heat Exchanger 62,042.96 kgal/hr
Temperature: 98.23oF
Cooling Water 2 from
Heat Exchanger 42,042.96 kgal/hr
Temperature: 85.4oF
Flash 5 Water Phase39.09 lbmol/hr Methane
5,241.96 lbmol/hr Water
389.02 lbmol/hr Carbon Dioxide
18.89 lbmol/hr Carbon
Monoxide
2.66 lbmol/hr Hydrogen
0.52 lbmol/hr Nitrogen
5.26 lbmol/hr Ethane
9.15 lbmol/hr Propane
9.36 lbmol/hr Butane
2.48 lbmol/hr Naptha
0.01 lbmol/hr Diesel
0 lbmol/hr C21-C25
Flash 5 Vapor4,708.08 lbmol/hr Methane
87.32 lbmol/hr Water
9,106.89 lbmol/hr Carbon Dioxide
5,471.30 lbmol/hr Carbon Monoxide
11,157.69 lbmol/hr Hydrogen
161.22 lbmol/hr Nitrogen
78.44 lbmol/hr Ethane
32.43 lbmol/hr Propane
8.95 lbmol/hr Butane
102.51 lbmol/hr Naptha
0 lbmol/hr Diesel
Flash 5 Liquid17.35 lbmol/hr Methane
11.36 lbmol/hr Water
172.68 lbmol/hr Carbon Dioxide
8.39 lbmol/hr Carbon Monoxide
1.18 lbmol/hr Hydrogen
0.23 lbmol/hr Nitrogen
2.33 lbmol/hr Ethane
4.06 lbmol/hr Propane
4.16 lbmol/hr Butane
2,209.0 lbmol/hr Naptha
4.89 lbmol/hr Diesel
Refridgeration
Exiting RefrigerantTemperature: 95
oF
Incoming RefrigerantTemperature: 5
oF
Absorber Gas Feed4,740.12 lbmol/hr Methane
150.56 lbmol/hr Water
9,402.98 lbmol/hr Carbon Dioxide
5,487.44 lbmol/hr Carbon Monoxide
11,160.37 lbmol/hr Hydrogen
161.67 lbmol/hr Nitrogen
82.39 lbmol/hr Ethane
38.56 lbmol/hr Propane
13.91 lbmol/hr Butane
168.83 lbmol/hr Naptha
0 lbmol/hr Diesel
Stripper 1 Feed Gas4,729.12 lbmol/hr Methane
108.54 lbmol/hr Water
9,273.06 lbmol/hr Carbon Dioxide
5,482.37 lbmol/hr Carbon
Monoxide
11,159.73 lbmol/hr Hydrogen
161.53 lbmol/hr Nitrogen
80.59 lbmol/hr Ethane
34.22 lbmol/hr Propane
6.34 lbmol/hr Butane
35.02 lbmol/hr Naptha
0 lbmol/hr Diesel
Absorber Liquid Effluent28.35 lbmol/hr Methane
108.54 lbmol/hr Water
302.6 lbmol/hr Carbon Dioxide
13.45 lbmol/hr Carbon Monoxide
1.82 lbmol/hr Hydrogen
0.37 lbmol/hr Nitrogen
4.13 lbmol/hr Ethane
8.4 lbmol/hr Propane
11.73 lbmol/hr Butane
2,342.81 lbmol/hr Naptha
4.89 lbmol/hr Diesel
Heat Exchanger 2
Heat Exchanger 3
P-54
Cooling Water 1 to
Heat Exchanger 10752.21 kgal/hr
Temperature: 119.3oF
Cooling Water 1 752.21 kgal/hr
Temperature: 85oF
Heat Exchanger 4
Cooling Water 2 2,042.96 kgal/hr
Temperature: 85oF
Cooling Water 2 to
Heat Exchanger 52,042.96 kgal/hr
Temperature: 85.4oF
P-60
Fuel Gas4,888.33 lbmol/hr Methane
1,505.6 lbmol/hr Water
10,074.08 lbmol/hr Carbon
Dioxide
5,579.05 lbmol/hr Carbon
Monoxide
11,183.96 lbmol/hr Hydrogen
164.27 lbmol/hr Nitrogen
93.32 lbmol/hr Ethane
54.46 lbmol/hr Propane
35.8 lbmol/hr Butane
38.22 lbmol/hr Naptha
0.07 lbmol/hr Diesel
Fuel Air29.93 lbmol/hr Methane
1,108.78 lbmol/hr Water
10,480.22 lbmol/hr Oxygen
352.62 lbmol/hr Carbon Dioxide
13.68 lbmol/hr Carbon Monoxide
1.69 lbmol/hr Hydrogen
39,441.98 lbmol/hr Nitrogen
4.67 lbmol/hr Ethane
8.74 lbmol/hr Propane
6.74 lbmol/hr Butane
0.01 lbmol/hr Naptha
0 lbmol/hr Diesel
Stripper 2 Feed Liquid29.93 lbmol/hr Methane
67,623.29 lbmol/hr Water
352.62 lbmol/hr Carbon
Dioxide
13.68 lbmol/hr Carbon
Monoxide
1.69 lbmol/hr Hydrogen
0.37 lbmol/hr Nitrogen
4.67 lbmol/hr Ethane
8.74 lbmol/hr Propane
6.77 lbmol/hr Butane
0.01 lbmol/hr Naptha
0 lbmol/hr Diesel
0.1 lbmol/hr C21-C25
Waste Water0 lbmol/hr Methane
6,514.5 lbmol/hr Water
19.78 lbmol/hr Oxygen
0 lbmol/hr Carbon Dioxide
0 lbmol/hr Carbon Monoxide
0 lbmol/hr Hydrogen
48.39 lbmol/hr Nitrogen
0 lbmol/hr Ethane
0 lbmol/hr Propane
0.04 lbmol/hr Butane
0 lbmol/hr Naptha
0 lbmol/hr Diesel
0.1 lbmol/hr C21-C25
Air10,480.22 lbmol/hr Oxygen
39,441.98 lbmol/hr Nitrogen
Appendix A.1.4 – Naphtha Separation Unit
Figure 6. Naphtha Separation Unit PFD
53
Figure 7. ASPEN Diagram for Diesel and Wax Separation Unit
54
Figure 8. ASPEN Diagram for Naphtha Separation Unit
55
Table 5. Stream Table for Separation Streams 1-6
1 2 3 4 5 6
Methane 4,917.67 1.03 4,917.67 4,917.67 4,917.00 0.66
Water 69,015.07 123.16 69,015.07 69,015.07 68,825.57 189.50
Oxygen - - - - - -
Carbon Dioxide 10,441.02 3.70 10,441.02 10,441.02 10,438.39 2.63
Carbon Monoxide 5,592.12 0.71 5,592.12 5,592.12 5,591.68 0.44
Hydrogen 11,185.41 0.24 11,185.41 11,185.41 11,185.27 0.14
Nitrogen 164.63 0.02 164.63 164.63 164.62 0.01
Ethane 98.30 0.07 98.30 98.30 98.26 0.05
Propane 65.47 0.11 65.47 65.47 65.39 0.08
Butane 49.00 0.19 49.00 49.00 48.86 0.14
Naptha 3,072.78 112.15 3,072.78 3,072.78 3,039.82 32.97
Diesel 556.87 688.39 556.87 556.87 351.33 205.55
C21-C25 5.89 202.69 5.89 5.89 0.23 5.67
C26-C29 0.40 90.82 0.40 0.40 0.00 0.40
C30-C35 0.07 73.76 0.07 0.07 0.00 0.07
C36-C47 0.00 53.43 0.00 0.00 0.00 0.00
C48+ 0.00 12.79 0.00 0.00 0.00 0.00
Methane 0.0468 0.0008 0.0468 0.0468 0.0470 0.0015
Water 0.6563 0.0903 0.6563 0.6563 0.6572 0.4324
Oxygen - - - - - -
Carbon Dioxide 0.0993 0.0027 0.0993 0.0993 0.0997 0.0060
Carbon Monoxide 0.0532 0.0005 0.0532 0.0532 0.0534 0.0010
Hydrogen 0.1064 0.0002 0.1064 0.1064 0.1068 0.0003
Nitrogen 0.0016 0.0000 0.0016 0.0016 0.0016 0.0000
Ethane 0.0009 0.0000 0.0009 0.0009 0.0009 0.0001
Propane 0.0006 0.0001 0.0006 0.0006 0.0006 0.0002
Butane 0.0005 0.0001 0.0005 0.0005 0.0005 0.0003
Naptha 0.0292 0.0823 0.0292 0.0292 0.0290 0.0752
Diesel 0.0053 0.5050 0.0053 0.0053 0.0034 0.4690
C21-C25 0.0001 0.1487 0.0001 0.0001 0.0000 0.0129
C26-C29 0.0000 0.0666 0.0000 0.0000 0.0000 0.0009
C30-C35 0.0000 0.0541 0.0000 0.0000 0.0000 0.0002
C36-C47 0.0000 0.0392 0.0000 0.0000 0.0000 0.0000
C48+ 0.0000 0.0094 0.0000 0.0000 0.0000 0.0000
105,165.00 1,363.26 105,165.00 105,165.00 104,726.00 438.30
2,427,950.00 341,148.00 2,427,950.00 2,427,950.00 2,373,730.00 54,217.66
3,691,580.00 8,495.44 3,641,110.00 3,493,420.00 3,492,020.00 1,395.42
408.49 408.49 396.69 365.00 365.00 365.00
265.41 265.41 265.41 265.41 265.41 265.41
Phase
Vapor Fraction 1.00 - 1.00 1.00 1.00 -
Liquid Fraction - 1.00 0.00 0.00 - 1.00
-89,392 -184,570 -89,518 -89,919 -89,734 -134,250
-3,872 -738 -3,877 -3,895 -3,959 -1,085
-9.4009.E+09 -2.5162.E+08 -9.4142.E+09 -9.4563.E+09 -9.3975.E+09 -5.8841.E+07
-12.20 -368.91 -12.35 -12.83 -12.11 -183.68
-0.53 -1.47 -0.53 -0.56 -0.53 -1.48
0.03 0.16 0.03 0.03 0.03 0.31
0.66 40.16 0.67 0.70 0.68 38.85
23.09 250.24 23.09 23.09 22.67 123.70
58,269.07 6,951.80 58,269.07 58,269.07 57,143.72 1,125.35
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
56
Table 6. Stream Table for Separation Streams 7-12
7 8 9 10 11 12
Methane 4,917.00 4,765.34 3.45 148.22 4,765.34 4,764.52
Water 68,825.57 5,675.31 69.60 63,080.66 5,675.31 5,340.65
Oxygen - - - - - -
Carbon Dioxide 10,438.39 9,672.73 17.41 748.25 9,672.73 9,668.58
Carbon Monoxide 5,591.68 5,499.07 2.11 90.51 5,499.07 5,498.58
Hydrogen 11,185.27 11,161.66 0.54 23.07 11,161.66 11,161.53
Nitrogen 164.62 161.98 0.06 2.57 161.98 161.97
Ethane 98.26 86.09 0.28 11.89 86.09 86.03
Propane 65.39 45.75 0.45 19.19 45.75 45.64
Butane 48.86 22.61 0.60 25.66 22.61 22.46
Naptha 3,039.82 2,316.55 722.74 0.53 2,316.55 2,313.98
Diesel 351.33 5.96 345.32 0.05 5.96 4.90
C21-C25 0.23 0.00 0.13 0.10 0.00 0.00
C26-C29 0.00 0.00 0.00 0.00 0.00 0.00
C30-C35 0.00 0.00 0.00 0.00 0.00 -
C36-C47 0.00 - 0.00 0.00 - -
C48+ 0.00 - - - - -
Methane 0.0470 0.1209 0.0030 0.0023 0.1209 0.1220
Water 0.6572 0.1440 0.0599 0.9833 0.1440 0.1367
Oxygen - - - - - -
Carbon Dioxide 0.0997 0.2454 0.0150 0.0117 0.2454 0.2475
Carbon Monoxide 0.0534 0.1395 0.0018 0.0014 0.1395 0.1407
Hydrogen 0.1068 0.2832 0.0005 0.0004 0.2832 0.2857
Nitrogen 0.0016 0.0041 0.0001 0.0000 0.0041 0.0041
Ethane 0.0009 0.0022 0.0002 0.0002 0.0022 0.0022
Propane 0.0006 0.0012 0.0004 0.0003 0.0012 0.0012
Butane 0.0005 0.0006 0.0005 0.0004 0.0006 0.0006
Naptha 0.0290 0.0588 0.6216 0.0000 0.0588 0.0592
Diesel 0.0034 0.0002 0.2970 0.0000 0.0002 0.0001
C21-C25 0.0000 0.0000 0.0001 0.0000 0.0000 0.0000
C26-C29 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000
C30-C35 0.0000 0.0000 0.0000 0.0000 0.0000 -
C36-C47 0.0000 - 0.0000 0.0000 - -
C48+ 0.0000 - - - - -
104,726.00 39,413.05 1,162.67 64,150.69 39,413.05 39,068.85
2,373,730.00 1,041,050.00 155,503.00 1,177,170.00 1,041,050.00 1,034,290.00
1,186,210.00 1,155,810.00 4,035.61 21,386.59 1,130,230.00 1,140,840.00
260.00 265.61 265.61 265.61 250.00 262.53
265.41 265.41 265.41 265.41 265.41 265.41
Phase
Vapor Fraction 0.38 1.00 - - 1.00 1.00
Liquid Fraction 0.62 - 1.00 1.00 0.00 -
-101,090 -70,611 -119,900 -119,480 -70,789 -70,360
-4,460 -2,673 -896 -6,511 -2,680 -2,658
-1.0590.E+10 -2.7830.E+09 -1.3940.E+08 -7.6645.E+09 -2.7900.E+09 -2.7489.E+09
-26.32 -10.17 -211.30 -32.86 -10.42 -10.20
-1.16 -0.39 -1.58 -1.79 -0.39 -0.39
0.09 0.03 0.29 3.00 0.03 0.03
2.00 0.90 38.53 55.04 0.92 0.91
22.67 26.41 133.75 18.35 26.41 26.47
57,143.72 34,509.96 3,440.21 19,193.55 34,509.96 34,396.64
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
57
Table 7. Stream Table for Separation Streams 13-18
13 14 15 16 17 18
Methane 0.01 0.80 4,764.52 4,708.08 17.35 39.09
Water 0.23 334.44 5,340.65 87.32 11.36 5,241.96
Oxygen - - - - - -
Carbon Dioxide 0.06 4.09 9,668.58 9,106.89 172.68 389.02
Carbon Monoxide 0.01 0.49 5,498.58 5,471.30 8.39 18.89
Hydrogen 0.00 0.12 11,161.53 11,157.69 1.18 2.66
Nitrogen 0.00 0.01 161.97 161.22 0.23 0.52
Ethane 0.00 0.06 86.03 78.44 2.33 5.26
Propane 0.00 0.11 45.64 32.43 4.06 9.15
Butane 0.00 0.14 22.46 8.95 4.16 9.36
Naptha 2.57 0.00 2,313.98 102.51 2,209.00 2.48
Diesel 1.06 0.00 4.90 0.00 4.89 0.01
C21-C25 0.00 0.00 0.00 - 0.00 0.00
C26-C29 0.00 0.00 0.00 - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
Methane 0.0030 0.0024 0.1220 0.1523 0.0071 0.0068
Water 0.0574 0.9829 0.1367 0.0028 0.0047 0.9167
Oxygen - - - - - -
Carbon Dioxide 0.0153 0.0120 0.2475 0.2946 0.0709 0.0680
Carbon Monoxide 0.0018 0.0014 0.1407 0.1770 0.0034 0.0033
Hydrogen 0.0005 0.0004 0.2857 0.3609 0.0005 0.0005
Nitrogen 0.0001 0.0000 0.0041 0.0052 0.0001 0.0001
Ethane 0.0002 0.0002 0.0022 0.0025 0.0010 0.0009
Propane 0.0004 0.0003 0.0012 0.0010 0.0017 0.0016
Butane 0.0005 0.0004 0.0006 0.0003 0.0017 0.0016
Naptha 0.6526 0.0000 0.0592 0.0033 0.9070 0.0004
Diesel 0.2682 0.0000 0.0001 0.0000 0.0020 0.0000
C21-C25 0.0000 0.0000 0.0000 - 0.0000 0.0000
C26-C29 0.0000 0.0000 0.0000 - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
3.93 340.27 39,068.85 30,914.81 2,435.64 5,718.40
514.13 6,247.24 1,034,290.00 673,457.00 246,729.00 114,106.00
13.38 113.25 677,484.00 693,332.00 5,907.39 1,842.33
262.53 262.53 95.00 95.00 95.00 95.00
265.41 265.41 265.41 265.41 265.41 265.41
Phase
Vapor Fraction - - 0.76 1.00 - -
Liquid Fraction 1.00 1.00 0.24 - 1.00 1.00
-117,900 -119,550 -75,500 -63,734 -109,940 -124,450
-902 -6,511 -2,852 -2,926 -1,085 -6,237
-4.6384.E+05 -4.0678.E+07 -2.9497.E+09 -1.9703.E+09 -2.6776.E+08 -7.1164.E+08
-207.02 -32.93 -18.10 -2.43 -174.53 -35.58
-1.58 -1.79 -0.68 -0.11 -1.72 -1.78
0.29 3.00 0.06 0.04 0.41 3.10
38.42 55.16 1.53 0.97 41.77 61.94
130.68 18.36 26.47 21.78 101.30 19.95
11.42 101.90 34,396.64 26,704.74 5,748.94 1,942.96
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
58
Table 8. Stream Table for Separation Streams 19-24
19 20 21 22 23 24
Methane 4,740.12 17.35 4,729.12 28.35 4,729.12 4,729.12
Water 150.56 11.36 53.39 108.54 53.39 53.39
Oxygen - - - - - -
Carbon Dioxide 9,402.98 172.68 9,273.06 302.60 9,273.06 9,273.06
Carbon Monoxide 5,487.44 8.39 5,482.37 13.45 5,482.37 5,482.37
Hydrogen 11,160.37 1.18 11,159.73 1.82 11,159.73 11,159.73
Nitrogen 161.67 0.23 161.53 0.37 161.53 161.53
Ethane 82.39 2.33 80.59 4.13 80.59 80.59
Propane 38.56 4.06 34.22 8.40 34.22 34.22
Butane 13.91 4.16 6.34 11.73 6.34 6.34
Naptha 168.83 2,209.00 35.02 2,342.81 35.02 35.02
Diesel 0.00 4.89 0.00 4.89 0.00 0.00
C21-C25 - 0.00 0.00 0.00 0.00 0.00
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
Methane 0.1509 0.0071 0.1525 0.0100 0.1525 0.1525
Water 0.0048 0.0047 0.0017 0.0384 0.0017 0.0017
Oxygen - - - - - -
Carbon Dioxide 0.2994 0.0709 0.2990 0.1070 0.2990 0.2990
Carbon Monoxide 0.1747 0.0034 0.1768 0.0048 0.1768 0.1768
Hydrogen 0.3553 0.0005 0.3598 0.0006 0.3598 0.3598
Nitrogen 0.0051 0.0001 0.0052 0.0001 0.0052 0.0052
Ethane 0.0026 0.0010 0.0026 0.0015 0.0026 0.0026
Propane 0.0012 0.0017 0.0011 0.0030 0.0011 0.0011
Butane 0.0004 0.0017 0.0002 0.0041 0.0002 0.0002
Naptha 0.0054 0.9070 0.0011 0.8287 0.0011 0.0011
Diesel 0.0000 0.0020 0.0000 0.0017 0.0000 0.0000
C21-C25 - 0.0000 0.0000 0.0000 0.0000 0.0000
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
31,406.83 2,435.64 31,015.35 2,827.11 31,015.35 31,015.35
696,400.00 246,729.00 673,575.00 269,553.00 673,575.00 673,575.00
696,239.00 5,578.30 651,622.00 6,386.33 11,768,300.00 12,449,100.00
90.00 10.00 59.94 87.46 59.94 90.00
265.41 265.41 265.41 265.41 14.70 14.70
Phase
Vapor Fraction 1.00 - 1.00 - 1.00 1.00
Liquid Fraction 0.00 1.00 - 1.00 - -
-64,889 -114,340 -64,436 -112,460 -64,436 -64,201
-2,926 -1,129 -2,967 -1,180 -2,967 -2,956
-2.0379.E+09 -2.7849.E+08 -1.9985.E+09 -3.1795.E+08 -1.9985.E+09 -1.9912.E+09
-2.96 -182.78 -2.60 -161.75 3.15 3.59
-0.13 -1.80 -0.12 -1.70 0.14 0.17
0.05 0.44 0.05 0.44 0.00 0.00
1.00 44.23 1.03 42.21 0.06 0.05
22.17 101.30 21.72 95.35 21.72 21.72
27,209.14 5,748.94 26,699.16 6,258.91 26,699.16 26,699.16
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
59
Table 9. Stream Table for Separation Streams 25-30
25 26 27 28 29 30
Methane 189.14 189.14 189.14 27.37 0.98 3.45
Water 69,075.50 69,075.50 69,075.50 20.47 88.07 69.15
Oxygen - - - - - -
Carbon Dioxide 1,153.64 1,153.64 1,153.64 292.12 10.48 17.39
Carbon Monoxide 110.36 110.36 110.36 12.99 0.47 2.11
Hydrogen 25.92 25.92 25.92 1.76 0.06 0.54
Nitrogen 3.12 3.12 3.12 0.36 0.01 0.06
Ethane 17.39 17.39 17.39 3.99 0.14 0.28
Propane 28.98 28.98 28.98 8.11 0.29 0.44
Butane 36.23 36.23 36.23 11.33 0.41 0.59
Naptha 3.21 3.21 3.21 2,342.61 0.20 586.70
Diesel 0.07 0.07 0.07 4.89 0.00 13.11
C21-C25 0.10 0.10 0.10 0.00 0.00 0.00
C26-C29 0.00 0.00 0.00 - - 0.00
C30-C35 0.00 0.00 0.00 - - 0.00
C36-C47 0.00 0.00 0.00 - - 0.00
C48+ - - - - - -
Methane 0.0027 0.0027 0.0027 0.0100 0.0097 0.0050
Water 0.9778 0.9778 0.9778 0.0075 0.8710 0.0997
Oxygen - - - - - -
Carbon Dioxide 0.0163 0.0163 0.0163 0.1072 0.1036 0.0251
Carbon Monoxide 0.0016 0.0016 0.0016 0.0048 0.0046 0.0030
Hydrogen 0.0004 0.0004 0.0004 0.0006 0.0006 0.0008
Nitrogen 0.0000 0.0000 0.0000 0.0001 0.0001 0.0001
Ethane 0.0002 0.0002 0.0002 0.0015 0.0014 0.0004
Propane 0.0004 0.0004 0.0004 0.0030 0.0029 0.0006
Butane 0.0005 0.0005 0.0005 0.0042 0.0040 0.0008
Naptha 0.0000 0.0000 0.0000 0.8594 0.0020 0.8456
Diesel 0.0000 0.0000 0.0000 0.0018 0.0000 0.0189
C21-C25 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000
C26-C29 0.0000 0.0000 0.0000 - - 0.0000
C30-C35 0.0000 0.0000 0.0000 - - 0.0000
C36-C47 0.0000 0.0000 0.0000 - - 0.0000
C48+ - - - - - -
70,643.65 70,643.65 70,643.65 2,725.98 101.11 693.80
1,305,750.00 1,305,750.00 1,305,750.00 267,413.00 2,139.23 67,970.56
2,782,410.00 2,782,410.00 390,711.00 6,354.75 34.51 389,942.00
197.42 197.42 90.00 86.00 86.00 310.00
14.70 14.70 14.70 265.41 265.41 14.70
Phase
Vapor Fraction 0.08 0.08 0.01 - - 1.00
Liquid Fraction 0.92 0.92 0.99 1.00 1.00 -
-119,900 -119,900 -122,860 -112,260 -125,690 -85,514
-6,487 -6,487 -6,647 -1,144 -5,941 -873
-8.4702.E+09 -8.4702.E+09 -8.6791.E+09 -3.0602.E+08 -1.2708.E+07 -5.9330.E+07
-32.85 -32.85 -37.62 -166.92 -35.05 -134.88
-1.78 -1.78 -2.04 -1.70 -1.66 -1.38
0.03 0.03 0.18 0.43 2.93 0.00
0.47 0.47 3.34 42.08 61.99 0.17
18.48 18.48 18.48 98.10 21.16 97.97
21,374.49 21,374.49 21,374.49 6,221.39 37.51 1,572.75
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
60
Table 10. Stream Table for Separation Streams 31-36
31 32 33 34 35 36
Methane 0.00 3.45 3.45 0.00 1.71 1.71
Water 0.46 69.15 69.05 0.09 312.89 312.89
Oxygen - - - - - -
Carbon Dioxide 0.02 17.39 17.38 0.01 6.39 6.39
Carbon Monoxide 0.00 2.11 2.11 0.00 1.15 1.15
Hydrogen 0.00 0.54 0.54 0.00 0.39 0.39
Nitrogen 0.00 0.06 0.06 0.00 0.03 0.03
Ethane 0.00 0.28 0.28 0.00 0.11 0.11
Propane 0.00 0.44 0.44 0.00 0.19 0.19
Butane 0.01 0.59 0.58 0.00 0.33 0.33
Naptha 136.03 586.70 542.45 44.26 147.68 147.68
Diesel 332.21 13.11 0.53 12.58 894.99 894.99
C21-C25 0.13 0.00 0.00 0.00 208.35 208.35
C26-C29 0.00 0.00 0.00 0.00 91.22 91.22
C30-C35 0.00 0.00 0.00 0.00 73.83 73.83
C36-C47 0.00 0.00 - - 53.43 53.43
C48+ - - - - 12.79 12.79
Methane 0.0000 0.0050 0.0054 0.0000 0.0009 0.0009
Water 0.0010 0.0997 0.1084 0.0016 0.1733 0.1733
Oxygen - - - - - -
Carbon Dioxide 0.0000 0.0251 0.0273 0.0001 0.0035 0.0035
Carbon Monoxide 0.0000 0.0030 0.0033 0.0000 0.0006 0.0006
Hydrogen 0.0000 0.0008 0.0008 0.0000 0.0002 0.0002
Nitrogen 0.0000 0.0001 0.0001 0.0000 0.0000 0.0000
Ethane 0.0000 0.0004 0.0004 0.0000 0.0001 0.0001
Propane 0.0000 0.0006 0.0007 0.0000 0.0001 0.0001
Butane 0.0000 0.0008 0.0009 0.0000 0.0002 0.0002
Naptha 0.2901 0.8456 0.8517 0.7772 0.0818 0.0818
Diesel 0.7085 0.0189 0.0008 0.2210 0.4957 0.4957
C21-C25 0.0003 0.0000 0.0000 0.0000 0.1154 0.1154
C26-C29 0.0000 0.0000 0.0000 0.0000 0.0505 0.0505
C30-C35 0.0000 0.0000 0.0000 0.0000 0.0409 0.0409
C36-C47 0.0000 0.0000 - - 0.0296 0.0296
C48+ - - - - 0.0071 0.0071
468.87 693.80 636.86 56.94 1,805.49 1,805.49
87,531.95 67,970.56 60,462.37 7,508.19 395,880.00 395,880.00
2,198.35 323,253.00 323,059.00 193.68 10,270.48 290,232.00
310.00 235.00 235.00 235.00 399.98 385.63
14.70 14.70 14.70 14.70 265.41 14.70
Phase
Vapor Fraction - 0.92 1.00 - 0.01 0.25
Liquid Fraction 1.00 0.08 - 1.00 0.99 0.75
-145,410 -90,358 -88,136 -115,210 -172,210 -172,210
-779 -922 -928 -874 -785 -785
-6.8175.E+07 -6.2691.E+07 -5.6130.E+07 -6.5604.E+06 -3.1092.E+08 -3.1092.E+08
-288.54 -141.52 -134.96 -214.80 -323.46 -322.45
-1.55 -1.44 -1.42 -1.63 -1.48 -1.47
0.21 0.00 0.00 0.29 0.18 0.01
39.82 0.21 0.19 38.77 38.55 1.36
186.69 97.97 94.94 131.86 219.26 219.26
1,867.46 1,572.75 1,404.01 168.74 8,088.57 8,088.57
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
61
Table 11. Stream Table for Separation Streams 37-42
37 38 39 40 41 42
Methane 1.71 1.71 1.71 0.00 1.71 0.00
Water 313.46 313.46 313.46 0.00 313.46 0.00
Oxygen - - - - - -
Carbon Dioxide 6.42 6.42 6.42 0.00 6.42 0.00
Carbon Monoxide 1.15 1.15 1.15 0.00 1.15 0.00
Hydrogen 0.39 0.39 0.39 0.00 0.39 0.00
Nitrogen 0.03 0.03 0.03 0.00 0.03 0.00
Ethane 0.11 0.11 0.11 0.00 0.11 0.00
Propane 0.20 0.20 0.20 0.00 0.20 0.00
Butane 0.35 0.35 0.35 0.00 0.35 0.00
Naptha 327.97 327.97 327.97 0.00 327.97 0.00
Diesel 1,239.78 1,239.78 1,239.60 0.17 0.74 1,238.86
C21-C25 208.48 208.48 0.01 208.48 0.00 0.01
C26-C29 91.22 91.22 0.00 91.22 0.00 0.00
C30-C35 73.83 73.83 0.00 73.83 - -
C36-C47 53.43 53.43 0.00 53.43 - -
C48+ 12.79 12.79 0.00 12.79 - -
Methane 0.0007 0.0007 0.0009 0.0000 0.0026 0.0000
Water 0.1345 0.1345 0.1657 0.0000 0.4804 0.0000
Oxygen - - - - - -
Carbon Dioxide 0.0028 0.0028 0.0034 0.0000 0.0098 0.0000
Carbon Monoxide 0.0005 0.0005 0.0006 0.0000 0.0018 0.0000
Hydrogen 0.0002 0.0002 0.0002 0.0000 0.0006 0.0000
Nitrogen 0.0000 0.0000 0.0000 0.0000 0.0001 0.0000
Ethane 0.0000 0.0000 0.0001 0.0000 0.0002 0.0000
Propane 0.0001 0.0001 0.0001 0.0000 0.0003 0.0000
Butane 0.0001 0.0001 0.0002 0.0000 0.0005 0.0000
Naptha 0.1407 0.1407 0.1734 0.0000 0.5026 0.0000
Diesel 0.5318 0.5318 0.6554 0.0004 0.0011 1.0000
C21-C25 0.0894 0.0894 0.0000 0.4739 0.0000 0.0000
C26-C29 0.0391 0.0391 0.0000 0.2074 0.0000 0.0000
C30-C35 0.0317 0.0317 0.0000 0.1678 - -
C36-C47 0.0229 0.0229 0.0000 0.1215 - -
C48+ 0.0055 0.0055 0.0000 0.0291 - -
2,331.32 2,331.32 1,891.40 439.92 652.53 1,238.87
490,920.00 490,920.00 313,192.00 177,728.00 41,349.53 271,843.00
414,190.00 9,845,150.00 13,995,500.00 4,438.11 3,598,370.00 6,735.60
405.00 354.32 340.32 564.24 136.52 360.94
14.70 1.16 1.16 1.16 1.16 1.16
Phase
Vapor Fraction 0.27 0.56 1.00 - 1.00 -
Liquid Fraction 0.73 0.44 - 1.00 - 1.00
-159,720 -159,720 -121,660 -239,780 -96,466 -160,360
-758 -758 -735 -594 -1,522 -731
-3.7236.E+08 -3.7236.E+08 -2.3011.E+08 -1.0548.E+08 -6.2947.E+07 -1.9866.E+08
-305.91 -303.83 -222.65 -560.25 -81.60 -330.50
-1.45 -1.44 -1.34 -1.39 -1.29 -1.51
0.01 0.00 0.00 0.10 0.00 0.18
1.19 0.05 0.02 40.05 0.01 40.36
210.58 210.58 165.59 404.00 63.37 219.43
10,124.77 10,124.77 6,614.31 3,510.46 925.32 5,688.99
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
62
Table 12. Stream Table for Separation Streams 43-48
43 44 45 46 47 48
Methane 1.71 32.53 1.71 29.93 32.53 4,917.67
Water 313.46 402.98 313.46 67,623.29 402.98 69,015.07
Oxygen - - - - - -
Carbon Dioxide 6.42 315.92 6.42 352.62 315.92 10,441.02
Carbon Monoxide 1.15 16.25 1.15 13.68 16.25 5,592.12
Hydrogen 0.39 2.68 0.39 1.69 2.68 11,185.41
Nitrogen 0.03 0.45 0.03 0.37 0.45 164.63
Ethane 0.11 4.38 0.11 4.67 4.38 98.30
Propane 0.20 8.75 0.20 8.74 8.75 65.47
Butane 0.35 12.26 0.35 6.77 12.26 49.00
Naptha 327.97 3,213.03 327.97 0.01 3,213.03 3,072.78
Diesel 0.74 6.16 1,239.78 0.00 6.16 556.87
C21-C25 0.00 0.00 208.48 0.10 0.00 5.89
C26-C29 0.00 0.00 91.22 0.00 0.00 0.40
C30-C35 - 0.00 73.83 0.00 0.00 0.07
C36-C47 - - 53.43 0.00 - 0.00
C48+ - - 12.79 - - 0.00
Methane 0.0026 0.0081 0.0007 0.0004 0.0081 0.0468
Water 0.4804 0.1004 0.1345 0.9938 0.1004 0.6563
Oxygen - - - - - -
Carbon Dioxide 0.0098 0.0787 0.0028 0.0052 0.0787 0.0993
Carbon Monoxide 0.0018 0.0040 0.0005 0.0002 0.0040 0.0532
Hydrogen 0.0006 0.0007 0.0002 0.0000 0.0007 0.1064
Nitrogen 0.0001 0.0001 0.0000 0.0000 0.0001 0.0016
Ethane 0.0002 0.0011 0.0000 0.0001 0.0011 0.0009
Propane 0.0003 0.0022 0.0001 0.0001 0.0022 0.0006
Butane 0.0005 0.0031 0.0001 0.0001 0.0031 0.0005
Naptha 0.5026 0.8002 0.1407 0.0000 0.8002 0.0292
Diesel 0.0011 0.0015 0.5318 0.0000 0.0015 0.0053
C21-C25 0.0000 0.0000 0.0894 0.0000 0.0000 0.0001
C26-C29 0.0000 0.0000 0.0391 0.0000 0.0000 0.0000
C30-C35 - 0.0000 0.0317 0.0000 0.0000 0.0000
C36-C47 - - 0.0229 0.0000 - 0.0000
C48+ - - 0.0055 - - 0.0000
652.53 4,015.38 2,331.32 68,041.85 4,015.38 105,164.70
41,349.53 369,225.00 490,920.00 1,235,600.00 369,225.00 2,427,944.93
361,371.00 529,409.00 345,281.00 19,784.54 304,824.00 3,597,379.09
298.72 165.48 366.33 67.63 95.00 386.70
14.70 14.70 14.70 14.70 14.70 265.41
Phase
Vapor Fraction 1.00 0.28 0.24 - 0.18 1.00
Liquid Fraction - 0.72 0.76 1.00 0.82 0.00
-91,875 -105,120 -165,430 -123,150 -109,770 -89,631
-1,450 -1,143 -786 -6,782 -1,194 -3,882
-5.9952.E+07 -4.2210.E+08 -3.8566.E+08 -8.3795.E+09 -4.4077.E+08 -9.4260.E+09
-79.85 -146.59 -312.60 -38.93 -154.34 -12.48
-1.26 -1.59 -1.48 -2.14 -1.68 -0.54
0.00 0.01 0.01 3.44 0.01 0.03
0.11 0.70 1.42 62.45 1.21 0.67
63.37 91.95 210.58 18.16 91.95 23.09
925.32 8,550.72 10,124.77 19,923.82 8,550.72 58,269.07
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
63
Table 13. Stream Table for Separation Streams 49-AIR
49 50 51 52 53 AIR
Methane 0.04 32.04 4,740.12 32.04 3.45 -
Water 330.37 63.24 150.56 63.24 69.60 -
Oxygen - - - - - 10,500.00
Carbon Dioxide 1.81 296.10 9,402.98 296.10 17.41 -
Carbon Monoxide 0.01 16.14 5,487.44 16.14 2.11 -
Hydrogen 0.00 2.68 11,160.37 2.68 0.54 -
Nitrogen 0.00 0.45 161.67 0.45 0.06 39,500.00
Ethane 0.04 3.95 82.39 3.95 0.28 -
Propane 0.24 6.13 38.56 6.13 0.45 -
Butane 0.66 4.97 13.91 4.97 0.60 -
Naptha 0.00 66.32 168.83 66.32 722.74 -
Diesel 0.00 0.00 0.00 0.00 345.31 -
C21-C25 0.00 - - - 0.13 -
C26-C29 - - - - 0.00 -
C30-C35 - - - - 0.00 -
C36-C47 - - - - 0.00 -
C48+ - - - - - -
Methane 0.0001 0.0651 0.1509 0.0651 0.0030 -
Water 0.9916 0.1285 0.0048 0.1285 0.0599 -
Oxygen - - - - - 0.2100
Carbon Dioxide 0.0054 0.6018 0.2994 0.6018 0.0150 -
Carbon Monoxide 0.0000 0.0328 0.1747 0.0328 0.0018 -
Hydrogen 0.0000 0.0055 0.3553 0.0055 0.0005 -
Nitrogen 0.0000 0.0009 0.0051 0.0009 0.0001 0.7900
Ethane 0.0001 0.0080 0.0026 0.0080 0.0002 -
Propane 0.0007 0.0125 0.0012 0.0125 0.0004 -
Butane 0.0020 0.0101 0.0004 0.0101 0.0005 -
Naptha 0.0000 0.1348 0.0054 0.1348 0.6216 -
Diesel 0.0000 0.0000 0.0000 0.0000 0.2970 -
C21-C25 0.0000 - - - 0.0001 -
C26-C29 - - - - 0.0000 -
C30-C35 - - - - 0.0000 -
C36-C47 - - - - 0.0000 -
C48+ - - - - - -
333.18 492.02 31,406.83 492.02 1,162.67 50,000.00
6,082.76 22,942.82 696,400.00 22,942.82 155,502.07 1,442,519.86
100.85 209,789.00 718,734.00 17,747.36 5,030.01 20,434,359.10
124.24 124.24 106.77 432.43 380.99 100.00
14.70 14.70 265.41 265.41 265.41 14.70
Phase
Vapor Fraction - 1.00 1.00 1.00 0.02 1.00
Liquid Fraction 1.00 - 0.00 - 0.98 -
-122,070 -131,900 -64,727 -127,080 -109,666 160
-6,687 -2,829 -2,919 -2,725 -820 6
-4.0673.E+07 -6.4898.E+07 -2.0329.E+09 -6.2525.E+07 -1.2751.E+08 8.0154.E+06
-37.32 -22.66 -2.67 -21.83 -198.51 1.31
-2.04 -0.49 -0.12 -0.47 -1.48 0.05
3.30 0.00 0.04 0.03 0.23 0.00
60.31 0.11 0.97 1.29 30.91 0.07
18.26 46.63 22.17 46.63 133.75 28.85
98.57 504.40 27,209.14 504.40 3,440.20 42,895.68
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
64
Table 14. Stream Table for Separation Streams FuelAir-Waste H2O
DIESEL FTR FUELAIR FUELGAS NAPTHA WASTEH2O
Methane 0.00 4,918.70 29.93 4,888.33 0.44 0.00
Water 0.00 69,138.23 1,108.78 1,505.60 9.38 66,514.50
Oxygen - - 10,480.22 - - 19.78
Carbon Dioxide 0.00 10,444.72 352.62 10,074.08 18.01 0.00
Carbon Monoxide 0.00 5,592.83 13.68 5,579.05 0.10 0.00
Hydrogen 0.00 11,185.65 1.69 11,183.96 0.00 0.00
Nitrogen 0.00 164.65 39,441.98 164.27 0.00 58.39
Ethane 0.00 98.37 4.67 93.32 0.39 0.00
Propane 0.00 65.58 8.74 54.46 2.38 0.00
Butane 0.00 49.19 6.74 35.80 6.62 0.02
Naptha 0.00 3,184.93 0.01 38.22 3,146.70 0.00
Diesel 1,238.86 1,245.26 0.00 0.07 6.16 0.00
C21-C25 0.01 208.58 0.00 0.00 0.00 0.10
C26-C29 0.00 91.22 0.00 0.00 - 0.00
C30-C35 - 73.83 0.00 0.00 - 0.00
C36-C47 - 53.43 0.00 0.00 - 0.00
C48+ - 12.79 - - - -
Methane 0.0000 0.0462 0.0006 0.1454 0.0001 0.0000
Water 0.0000 0.6490 0.0216 0.0448 0.0029 0.9988
Oxygen - - 0.2037 - - 0.0003
Carbon Dioxide 0.0000 0.0980 0.0069 0.2997 0.0056 0.0000
Carbon Monoxide 0.0000 0.0525 0.0003 0.1660 0.0000 0.0000
Hydrogen 0.0000 0.1050 0.0000 0.3327 0.0000 0.0000
Nitrogen 0.0000 0.0015 0.7666 0.0049 0.0000 0.0009
Ethane 0.0000 0.0009 0.0001 0.0028 0.0001 0.0000
Propane 0.0000 0.0006 0.0002 0.0016 0.0007 0.0000
Butane 0.0000 0.0005 0.0001 0.0011 0.0021 0.0000
Naptha 0.0000 0.0299 0.0000 0.0011 0.9864 0.0000
Diesel 1.0000 0.0117 0.0000 0.0000 0.0019 0.0000
C21-C25 0.0000 0.0020 0.0000 0.0000 0.0000 0.0000
C26-C29 0.0000 0.0009 0.0000 0.0000 - 0.0000
C30-C35 - 0.0007 0.0000 0.0000 - 0.0000
C36-C47 - 0.0005 0.0000 0.0000 - 0.0000
C48+ - 0.0001 - - - -
1,238.87 106,528.00 51,449.06 33,617.15 3,190.18 66,592.79
271,843.00 2,769,090.00 1,477,540.45 743,723.00 340,199.00 1,200,578.74
6,735.60 3,700,070.00 19,751,003.80 13,450,400.00 8,322.56 19,164.66
360.94 408.49 66.05 88.25 124.24 58.18
1.16 265.41 14.70 14.70 14.70 14.70
Phase
Vapor Fraction - 0.99 1.00 1.00 - -
Liquid Fraction 1.00 0.01 - - 1.00 1.00
-160,360 -90,610 -3,526 -68,145 -105,070 -122,987
-731 -3,486 -123 -3,080 -985 -6,822
-1.9866.E+08 -9.6525.E+09 -1.8141.E+08 -2.2908.E+09 -3.3520.E+08 -8.1900.E+09
-330.50 -16.77 0.90 3.13 -186.67 -39.39
-1.51 -0.65 0.03 0.14 -1.75 -2.18
0.18 0.03 0.00 0.00 0.38 3.47
40.36 0.75 0.07 0.06 40.88 62.65
219.43 25.99 28.72 22.12 106.64 18.03
5,688.99 65,220.87 43,520.22 28,149.84 7,947.74 19,299.28
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
65
Table 15. Stream Table for Separation Streams FuelAir-Waste H2O
WAX
Methane 0.00
Water 0.00
Oxygen -
Carbon Dioxide 0.00
Carbon Monoxide 0.00
Hydrogen 0.00
Nitrogen 0.00
Ethane 0.00
Propane 0.00
Butane 0.00
Naptha 0.00
Diesel 0.17
C21-C25 208.48
C26-C29 91.22
C30-C35 73.83
C36-C47 53.43
C48+ 12.79
Methane 0.0000
Water 0.0000
Oxygen -
Carbon Dioxide 0.0000
Carbon Monoxide 0.0000
Hydrogen 0.0000
Nitrogen 0.0000
Ethane 0.0000
Propane 0.0000
Butane 0.0000
Naptha 0.0000
Diesel 0.0004
C21-C25 0.4739
C26-C29 0.2074
C30-C35 0.1678
C36-C47 0.1215
C48+ 0.0291
439.92
177,728.00
4,593.45
564.24
1.16
Phase
Vapor Fraction 0.00
Liquid Fraction 1.00
-239,780
-594
-1.0548.E+08
-560.25
-1.39
0.10
38.69
404.00
3,510.46
Stream Number
Mole Fraction
Pressure (psia)
Temperature (°F)
Total Volumetric Flow Rate (ft3/hr)
Total Mass Flow Rate (lbm/hr)
Total Molar Flow Rate (lbmol/hr)
Molar Flow Rate (lbmol/hr)
Enthalpy (Btu/hr)
Enthalpy (Btu/lbm)
Enthalpy (Btu/lbmol)
Liquid Volume, 60°F (lbm/ft3)
Average Molecular Weight
Density (lbm/ft3)
Density (lbmol/ft3)
Entropy (Btu/lbm-R)
Entropy (Btu/lbmol-R)
66
Table 16. Stream Table for Heat Integration Cool1-1 to Cool2-3
COOL1-1 COOL1IN COOL1OUT COOL2-1 COOL2-2 COOL2-3
Methane - - - - - -
Water 344,850.00 344,850.00 344,850.00 929,600.00 929,600.00 929,600.00
Oxygen - - - - - -
Carbon Dioxide - - - - - -
Carbon Monoxide - - - - - -
Hydrogen - - - - - -
Nitrogen - - - - - -
Ethane - - - - - -
Propane - - - - - -
Butane - - - - - -
Naptha - - - - - -
Diesel - - - - - -
C21-C25 - - - - - -
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
Methane - - - - - -
Water 1.0000 1.0000 1.0000 1.0000 1.0000 1.0000
Oxygen - - - - - -
Carbon Dioxide - - - - - -
Carbon Monoxide - - - - - -
Hydrogen - - - - - -
Nitrogen - - - - - -
Ethane - - - - - -
Propane - - - - - -
Butane - - - - - -
Naptha - - - - - -
Diesel - - - - - -
C21-C25 - - - - - -
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
344,850.00 344,850.00 344,850.00 929,600.00 929,600.00 929,600.00
6,212,569.31 6,212,569.31 6,212,569.31 16,747,004.30 16,747,004.30 16,747,004.30
102,491.35 100,555.61 102,524.06 271,121.98 273,037.01 273,104.55
119.34 85.00 119.90 85.39 98.23 98.67
14.70 14.70 14.70 14.70 14.70 14.70
Phase
Vapor Fraction 1.00 1.00 1.00 1.00 1.00 1.00
Liquid Fraction - - - - - -
-122,104 -122,688 -122,094 -122,682 -122,466 -122,458
-6,778 -6,810 -6,777 -6,810 -6,798 -6,797
-42,110,000,000 -42,310,000,000 -42,100,000,000 -114,000,000,000 -113,800,000,000 -113,800,000,000
-38 -39 -38 -39 -38 -38
-2 -2 -2 -2 -2 -2
3 3 3 3 3 3
61 62 61 62 61 61
18 18 18 18 18 18
99,708 99,708 99,708 268,778 268,778 268,778
Entropy (Btu/lbm-R)
Density (lbmol/ft3)
Density (lbm/ft3)
Average Molecular Weight
Liquid Volume, 60°F (lbm/ft3)
Temperature (°F)
Pressure (psia)
Enthalpy (Btu/lbmol)
Enthalpy (Btu/lbm)
Enthalpy (Btu/hr)
Entropy (Btu/lbmol-R)
Stream Number
Molar Flow Rate (lbmol/hr)
Mole Fraction
Total Molar Flow Rate (lbmol/hr)
Total Mass Flow Rate (lbm/hr)
Total Volumetric Flow Rate (ft3/hr)
67
Table 17. Stream Table for Heat Integration Cool2In to LP1Out
COOL2IN COOL2OUT COOL3IN COOL3OUT LP1IN LP1OUT
Methane - - - - - -
Water 929,600.00 929,600.00 31,400.00 31,400.00 50,000.00 50,000.00
Oxygen - - - - - -
Carbon Dioxide - - - - - -
Carbon Monoxide - - - - - -
Hydrogen - - - - - -
Nitrogen - - - - - -
Ethane - - - - - -
Propane - - - - - -
Butane - - - - - -
Naptha - - - - - -
Diesel - - - - - -
C21-C25 - - - - - -
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
Methane - - - - - -
Water 1.0000 1.0000 1.0000 1.0000 1.0000 1.0000
Oxygen - - - - - -
Carbon Dioxide - - - - - -
Carbon Monoxide - - - - - -
Hydrogen - - - - - -
Nitrogen - - - - - -
Ethane - - - - - -
Propane - - - - - -
Butane - - - - - -
Naptha - - - - - -
Diesel - - - - - -
C21-C25 - - - - - -
C26-C29 - - - - - -
C30-C35 - - - - - -
C36-C47 - - - - - -
C48+ - - - - - -
929,600.00 929,600.00 31,400.00 31,400.00 50,000.00 50,000.00
16,747,004.30 16,747,004.30 565,679.79 565,679.79 900,764.00 900,764.00
271,064.22 276,383.75 9,156.00 9,335.55 16,209.20 11,129,643.60
85.00 119.99 85.00 119.96 258.86 260.00
14.70 14.70 14.70 14.70 34.70 34.70
Phase
Vapor Fraction 1.00 1.00 1.00 1.00 1.00 -
Liquid Fraction - - - - - -
-122,688 -122,093 -122,688 -122,093 -119,478 -102,481
-6,810 -6,777 -6,810 -6,777 -6,632 -5,689
-114,100,000,000 -113,500,000,000 -3,852,000,000 -3,834,000,000 -5,974,000,000 -5,124,000,000
-39 -38 -39 -38 -34 -10
-2 -2 -2 -2 -2 -1
3 3 3 3 3 0
62 61 62 61 56 0
18 18 18 18 18 18
268,778 268,778 9,079 9,079 14,457 14,457
Entropy (Btu/lbm-R)
Density (lbmol/ft3)
Density (lbm/ft3)
Average Molecular Weight
Liquid Volume, 60°F (lbm/ft3)
Temperature (°F)
Pressure (psia)
Enthalpy (Btu/lbmol)
Enthalpy (Btu/lbm)
Enthalpy (Btu/hr)
Entropy (Btu/lbmol-R)
Stream Number
Molar Flow Rate (lbmol/hr)
Mole Fraction
Total Molar Flow Rate (lbmol/hr)
Total Mass Flow Rate (lbm/hr)
Total Volumetric Flow Rate (ft3/hr)
68
Table 18. Stream Table for Heat Integration MP1In and MP1Out
MP1IN MP1OUT
Methane - -
Water 1,950.00 1,950.00
Oxygen - -
Carbon Dioxide - -
Carbon Monoxide - -
Hydrogen - -
Nitrogen - -
Ethane - -
Propane - -
Butane - -
Naptha - -
Diesel - -
C21-C25 - -
C26-C29 - -
C30-C35 - -
C36-C47 - -
C48+ - -
Methane - -
Water 1.0000 1.0000
Oxygen - -
Carbon Dioxide - -
Carbon Monoxide - -
Hydrogen - -
Nitrogen - -
Ethane - -
Propane - -
Butane - -
Naptha - -
Diesel - -
C21-C25 - -
C26-C29 - -
C30-C35 - -
C36-C47 - -
C48+ - -
1,950.00 1,950.00
35,129.80 35,129.80
678.26 120,546.16
353.04 353.03
139.70 139.70
Phase
Vapor Fraction 1.00 0.01
Liquid Fraction - -
-117,390 -101,869
-6,516 -5,655
-228,910,359 -198,645,173
-31 -12
-2 -1
3 0
52 0
18 18
564 564
Entropy (Btu/lbm-R)
Density (lbmol/ft3)
Density (lbm/ft3)
Average Molecular Weight
Liquid Volume, 60°F (lbm/ft3)
Temperature (°F)
Pressure (psia)
Enthalpy (Btu/lbmol)
Enthalpy (Btu/lbm)
Enthalpy (Btu/hr)
Entropy (Btu/lbmol-R)
Stream Number
Molar Flow Rate (lbmol/hr)
Mole Fraction
Total Molar Flow Rate (lbmol/hr)
Total Mass Flow Rate (lbm/hr)
Total Volumetric Flow Rate (ft3/hr)
69
Appendix A.2 –Economics Summary
Table 19. Economic Summary Breakdown
Revenue
Unit Price ($) Units per Year Yearly Profit ($/year)
Products
LPG (C3, C4) ($0.30/lb.) 0.30$ 4,161,649.73 1,248,494.92$
Naphtha (C5 - C10) ($75.00/bbl.) 75.00$ 14,567,835.50 1,092,587,662.17$
Diesel (C10 - C20) ($90.00/bbl.) 90.00$ 10,524,316.93 947,188,523.57$
Total Product Revenue 29,253,802.16 2,041,024,680.65$
Credits
600 psia, 490°F HP Steam ($4.00/klbm) 4.00$ 15,088,319.41 60,353,277.64$
120 psia, 353°F MP Steam ($3.00/klbm) 3.00$ 94,263,783.12 282,791,349.37$
20 psia, 260°F LP Steam ($2.50/klbm) 2.50$ 7,233,134.92 18,082,837.30$
Electricity ($0.03/kW-h) 0.03$ - -$
Fuel Gas ($2.00/MBTU) 2.00$ 22,021,992.58 44,043,985.17$
Steam Condensate (99.9% v/v pure) ($2.00/klbm) 2.00$ 21,045,670.68 42,091,341.37$
Process/Cooling Tower Water (95% v/v Pure) ($0.35/kgal) 0.35$ 1,234,378,855.35 432,032,599.37$
Total Credits Revenue 1,394,031,756.08 879,395,390.22$
Total Revenue 2,920,420,070.88$
Variable Costs
Unit Price ($) Units per Year Yearly Cost ($/year)
Raw Materials Costs
Methane Feed ($2,000/MSCF) 2,000.00$ 167,291.67 334,583,333.33$
Hydrogen ($0.06/lb.) 0.06$ 7,984,590.67 479,075.44$
Carbon Dioxide (500psig, 100°F) ($400/MSCF) 400.00$ 54,194.59 21,677,835.89$
Oxygen ($100/short ton) 100.00$ 4,188,496.24 418,849,623.65$
Catalyst ($10/lb.) 10.00$ 3,450,922.88 34,509,228.77$
Total Raw Materials Costs 15,845,496.04 810,099,097.09$
Utilities
600 psia, 490°F HP Steam (klbm/hr) ($5/klbm) 5.00$ 24,966,493.34 124,832,466.70$
125psia, 353°F MP Steam (klbm/hr) ($4/klbm) 4.00$ 50,102.66 200,410.65$
20 psia, 260°F LP Steam (klbm/hr) ($3.5/klbm) 3.50$ - -$
Electricity (kw-hr/hr) ($0.04/kW-hr) 0.04$ 33,049,801.20 1,321,992.05$
Fuel Gas (MBTU/hr) ($3/MBTU) 3.00$ 737,297.85 2,211,893.54$
Steam Condensate (99.9% v/v pure) ($2.00/klbm) 2.00$ 111,386,648.86 222,773,297.72$
Process/Cooling Tower Water (95% v/v Pure) (kgal/hr) ($0.5/kgal) 0.50$ 1,231,732,522.81 615,866,261.40$
Waste Water Treatment (75% v/v Pure) (kgal/hr) ($6/kgal) 6.00$ - -$
Total Utility Costs 1,401,922,866.72 967,206,322.06$
Total Variable Costs 1,777,305,419.14$
Fixed Costs
Yearly Operating Expenses 30,614,848.25$
HI Unit Capital Recovery 100,000,000.00$
Syngas Unit Capital Recovery 400,000,000.00$
Depreciation 68,032,996.12$
Total Fixed Costs 598,647,844.38$
Total Manufacturing Costs 2,375,953,263.52$
15 Year Service Life, No Salvage
3% of total Capital Investment
70
Table 21. Products Breakdown
Unit Price
From HI Unit
(Units/Hour)
Syngas Unit
(Units/hr)
From Air
Separation Unit
(Units/Hour)
Waste Water
(Units/hr)
Miscellaneous
(Units/Hour)
Distillation
Columns
(Units/Hour)
Heat
Exchangers
(Units/Hour)
Fuel Gas
(Units/Hour)
Total
(Units/Hour)
Boiler Makeup
and Cooling
Water Makeup Total ($/hr)
Methane (MSCF/hr) ($2000/SMCF) 2,000.00$ - 20.83 - - - - - - 20.83 - 41,666.67$
Hydrogen (lbm/hr) ($0.06/lbm) 0.06$ 994.35 - - - - - - - 994.35 - 59.66$
Carbon Dioxide ($400/MSCF) 400.00$ - 6.75 - - - - - - 6.75 - 2,699.61$
Oxygen (Short Ton/hr) ($100/Short Ton) 100.00$ - 260.80 260.80 - - - - - 521.61 - 52,160.60$
Nitrogen (lbm/hr) ($0.05/lbm) -$ - - - - - - - - - - -$
-
-
600 psia, 490°F HP Steam (klbm/hr) ($5/klbm) 4.00$ - 494.51 2,608.03 - - 6.61 - - 3,109.15 - 12,436.61$
125psia, 353°F MP Steam (klbm/hr) ($4/klbm) 3.00$ 6.24 - - - - - - - 6.24 - 18.72$
20 psia, 260°F LP Steam (klbm/hr) ($3.5/klbm) 2.50$ - - - - - - - - - - -$
Electricity (kw-hr/hr) ($0.04/kW-hr) 0.04$ 1,559.86 - - - 2,555.93 - - - 4,115.79 - 164.63$
Fuel Gas (MBTU/hr) ($3/MBTU) 3.00$ 49.92 - - - - 41.90 - - 91.82 - 275.45$
Steam Condensate (99.9% v/v pure) (klbm/hr) ($2.00/1000 gal) 2.00$ - - - - 253.37 - 13,617.95 - 13,871.31 - 27,742.63$
Process/Cooling Tower Water (95% v/v Pure) (kgal/hr) ($0.5/kgal) 0.50$ 187.18 - 150,222.53 - - 117.97 2,863.66 - 153,391.35 - 76,695.67$
Waste Water Treatment (75% v/v Pure) (kgal/hr) ($6/kgal) 6.00$ - - - - - - - - - - -$
-
-
600 psia, 490°F HP Steam (klbm/hr) ($4.00/1000 lbm) 4.00$ - - - - - - 1,878.99 - 1,878.99 - 7,515.97$
120 psia, 353°F MP Steam (klbm/hr) ($3.00/1000 lbm) 3.00$ - - - - - - 11,738.95 - 11,738.95 - 35,216.86$
20 psia, 260°F LP Steam (klbm/hr) ($2.50/1000 lbm) 2.50$ - - - - - - 900.76 - 900.76 - 2,251.91$
Electricity (kw-hr/hr) ($0.03/kW-h) 0.03$ - - - - - - - - - - -$
Fuel Gas (MBTU/hr) ($2.00/MBTU) 2.00$ - - - - - - - 2,742.46 2,742.46 - 5,484.93$
Steam Condensate (99.9% v/v pure) (klbm/hr) ($2.00/1000 gal) 2.00$ 6.24 - 2,608.03 - - 6.61 - - 2,620.88 (78.63) 5,241.76$
Process/Cooling Tower Water (95% v/v Pure) (kgal/hr) ($0.35/1000gal) 0.35$ 187.18 - 150,222.53 144.68 253.37 117.97 2,795.17 - 153,720.90 (3,764.40) 53,802.32$
Waste Water Treatment (75% v/v Pure) (kgal/hr) ($6/kgal) 6.00$ - - - - - - - - - - -$
213,920.25$
Utilities and Credits Breakdown
Total
Raw Materials
Utilities
Credits
Unit Price Unit per Hour Total ($/hr)
LPG (C3, C4) ($0.4/lbm) 0.40$ 12,438.31 4,975.32$
Naptha (C5 - C10) ($75/bbl) 75.00$ 1,814.18 136,063.22$
Diesel (C11 - C20) ($90/bbl) 90.00$ 1,310.62 117,956.23$
258,994.77$
Products Breakdown
Products
Total
Table 20. Utilities and Credits Breakdown
71
Table 22. Inflation Effects on Profitability
Table 24. Economics Summary
year Revenue ($/year) Manufacturing Costs ($/year) Depreciation ($/year) Cash Flow ($/year)
0 2,920,420,070.88$ 2,307,920,267.40$ 68,032,996.12$ 432,825,757.05$
1 3,008,032,673.01$ 2,377,157,875.42$ 68,032,996.12$ 445,137,003.10$
2 3,098,273,653.20$ 2,448,472,611.68$ 68,032,996.12$ 457,817,586.53$
3 3,191,221,862.79$ 2,521,926,790.03$ 68,032,996.12$ 470,878,587.47$
4 3,286,958,518.68$ 2,597,584,593.74$ 68,032,996.12$ 484,331,418.43$
5 3,385,567,274.24$ 2,675,512,131.55$ 68,032,996.12$ 498,187,834.32$
6 3,487,134,292.46$ 2,755,777,495.49$ 68,032,996.12$ 512,459,942.69$
7 3,591,748,321.24$ 2,838,450,820.36$ 68,032,996.12$ 527,160,214.31$
8 3,699,500,770.87$ 2,923,604,344.97$ 68,032,996.12$ 542,301,494.08$
9 3,810,485,794.00$ 3,011,312,475.32$ 68,032,996.12$ 557,897,012.24$
10 3,924,800,367.82$ 3,101,651,849.58$ 68,032,996.12$ 573,960,395.94$
11 4,042,544,378.85$ 3,194,701,405.07$ 68,032,996.12$ 590,505,681.16$
12 4,163,820,710.22$ 3,290,542,447.22$ 68,032,996.12$ 607,547,324.93$
13 4,288,735,331.53$ 3,389,258,720.63$ 68,032,996.12$ 625,100,218.02$
14 4,417,397,391.47$ 3,490,936,482.25$ 68,032,996.12$ 643,179,697.90$
15 4,549,919,313.22$ 3,595,664,576.72$ 68,032,996.12$ 661,801,562.17$
Inflation Effects on Profitability
Down Time (Months) 1
Service Factor 0.92
Operating Life (Years) 15
Tax Rate (33%) 0.33
Total Equipment Costs ($2010) 212,603,112.88$
Lang Factor 4.8
Total Capital Investment ($2010) 1,020,494,941.82$
Depreciation 68,032,996.12$
Revenue ($/year) 2,920,420,070.88$
Manufacturing Costs ($/year) 2,375,953,263.52$
Gross Profit ($/year) 544,466,807.36$
Net Profit ($/year) 364,792,760.93$
Cash Flow ($/year) 432,825,757.05$
DCFRR (%) 42.20%
Cash Flow Payback Period (years) 2.36
Return on Investment (%/year) 35.75%
Interest Rate (%) 8%
Future Value of Earnings ($2010) 11,752,134,266.21$
Present Value of Earnings ($2010) 3,704,762,843.07$
Net Present Worth of Plant ($2010) 2,684,267,901.25$
Economics Summary
Table 23. Energy Efficiency
Metahne Carbon (lbm/hr) 659,375.78
Diesel Carbon (lbm/hr) 226,225.73
Naphtha Carbon (lbmC/hr) 230,294.46
HI Unit Carbon (lbmC/hr) 145,842.37
Carbon Efficency (%) 91%
Energy Efficiency
72
Appendix A.3 – Equipment Sizing Summary
Table 25. Equipotent Costs Summary
Fischer-Tropsch Reactors Separation
Reactors (43) 112,941,932.69$ Flash Drums
112,941,932.69$ Flash Drum 1 15,077,467.40$
Flash Drum 2 24,982,093.40$
Heat Integration and Steam Generation Flash Drum 3 1,145,707.06$
Heat Exchanger 1 619,713.84$ Flash Drum 4 16,105,093.61$
Heat Exchanger 2 34,315.66$ Flash Drum 5 1,236,657.18$
Heat Exchanger 3 4,256,875.63$ Flash Drum 6 110,971.38$
Heat Exchanger 4 305,247.34$ Flash Drum 7 116,036.98$
Heat Exchanger 5 2,133,654.58$ Flash Drum 8 47,271.07$
Heat Exchanger 6 25,602.86$ 58,821,298.07$
Heat Exchanger 7 1,082,934.26$
Heat Exchanger 8 5,939,158.30$
Heat Exchanger 9 995,320.55$ Decanter 1 38,950.49$
Heat Exchanger 10 12,987.46$ Total 38,950.49$
Heat Exchanger 11 377,255.89$
Heat Exchanger 12 638,453.56$
Heat Exchanger 13 4,003,341.68$ Distillation Column 1 5,162,589.48$
Heat Exchanger 14 44,495.26$ Distillation Column 2 2,331,833.54$
Heat Exchanger 15 719,040.64$ Absorber 617,227.81$
Heat Exchanger 16 385,149.88$ Stripper 1 900,839.18$
21,573,547.38$ Stripper 2 1,798,965.39$
Total 7,494,423.02$
Refrigeration
Compressor 80,000.00$
Condenser 331,260.89$ Compressor 1 231,595.27$
Evaporator 39,702.05$ Compressor 2 198,510.23$
450,962.94$ Fired Heater 10,587,212.46$
Blower 264,680.31$
Total 11,281,998.28$
Total Equipment Costs 212,603,112.88$
Distillation Columns
Decanter
Cost ($2010)
Miscellaneous Equipment
Total
Total
Total
Total
73
Table 26. Heat Exchanger Sizing Summary, Heat Exchangers 1-8
Table 27. Heat Exchanger Sizing Summary, Heat Exchangers 9-16
Heat Exchanger 1
(Vacuum)
Heat Exchanger 2
(Absorb)
Heat Exchanger 3
(Water)
Heat Exchanger 4
(Water)
Heat Exchanger 5
(Water)
Heat Exchanger 6
(Water)
Heat Exchanger 7
(Water)
Heat Exchanger 8
(LP Steam)
Pressure (psia) 265.41 14.70 14.70 265.41 265.41 265.41 265.41 265.41
Pressure (kPa) 1,829.93 101.33 101.33 1,829.93 1,829.93 1,829.93 1,829.93 1,829.93
Thin (⁰F) 408.49 197.42 196.38 107.61 262.52 265.61 322.89 365.00
Thout (⁰F) 396.69 196.38 90.00 90.00 95.00 250.00 260.00 322.89
Tcin (⁰F) 366.34 58.97 85.00 85.00 85.39 98.23 98.37 258.86
Tcout (⁰F) 405.00 90.00 119.34 85.39 98.23 98.37 119.99 260.00
Duty (Q) (Btu/hr) 13,298,509.80 7,508,704.81 201,448,612.00 6,087,116.51 200,820,741.00 7,048,922.85 339,500,782.00 849,852,921.00
Uo (Btu/hr-ft2-°F) 39.62 39.62 39.62 39.62 39.62 39.62 39.62 39.62
LMTD (⁰F) 12.42 121.80 26.34 11.54 54.49 159.38 181.49 82.84
Area (ft2) 27,015.78 1,555.79 192,997.01 13,306.94 93,014.44 1,116.13 47,209.39 258,911.40
Number of 6000ft3 Heat Exchangers 4.50 0.26 32.17 2.22 15.50 0.19 7.87 43.15
Cost of Heat Exchanger ($2010) 619,713.84$ 34,315.66$ 4,256,875.63$ 305,247.34$ 2,133,654.58$ 25,602.86$ 1,082,934.26$ 5,939,158.30$
Heat Exchanger 9
(MPSteam)
Heat Exchanger 10
(Water)
Heat Exchanger 11
(HP Steam)
Heat Exchanger 12
(MPSteam)
Heat Exchanger
13 (SynFeed)
Heat Exchanger 14
(Water)
Heat Exchanger 15
(SynFeed)
Heat Exchanger 16
(Naphtha Sep)
Pressure (psia) 265.41 265.41 614.70 314.70 314.70 14.70 314.70 265.41
Pressure (kPa) 1,829.93 1,829.93 4,238.18 2,169.75 2,169.75 101.33 2,169.75 1,829.93
Thin (⁰F) 396.69 310.00 1,358.07 500.47 1,883.77 165.28 1,950.00 396.69
Thout (⁰F) 365.00 235.00 500.47 425.00 1,358.07 95.00 1,883.77 386.70
Tcin (⁰F) 353.00 119.34 488.96 353.04 306.08 85.00 75.00 265.61
Tcout (⁰F) 353.00 119.90 490.00 353.03 1,000.00 119.97 1,000.00 380.99
Duty (Q) (Btu/hr) 42,164,981.60 3,360,858.23 115,783,640.00 116,053,926.00 890,187,130.00 18,682,218.10 114,847,111.00 11,897,607.00
Uo (Btu/hr-ft2-°F) 39.62 39.62 39.62 39.62 5.28 396.25 5.28 39.62
LMTD (⁰F) 24.52 149.81 198.17 105.23 965.44 23.37 1,333.62 51.59
Area (ft2) 43,389.96 566.17 14,744.76 27,832.72 174,521.50 2,017.31 16,299.83 5,820.55
Number of 6000ft3 Heat Exchangers 7.23 0.09 2.46 4.64 29.09 0.34 2.72 0.97
Cost of Heat Exchanger ($2010) 995,320.55$ 12,987.46$ 377,255.89$ 638,453.56$ 4,003,341.68$ 44,495.26$ 719,040.64$ 385,149.88$
74
Table 28. Flash Drum Sizing Summary
Table 29. Decanter Sizing Summary
Flash 1 Flash 2 Flash 3 Flash 4 Flash 5 Flash 6 Flash 7 Flash 8
Vessel Diameter (ft) 25.8989899 28.09742619 20.7219173 24.65322086 21.54593783 6.219369156 6.219369156 6.366149563
Vessel Height (ft) 77.70 84.29 62.17 73.96 64.64 18.66 19.10 30.79
Vessel Volume (ft3) 40931.68421 52264.97572 20965.29341 35304.66827 23567.16505 566.8264956 580.2039012 979.9377014
Pressure (psia) 265.409475 265.409475 265.409475 265.409475 265.409475 265.409475 265.409475 14.695949
Temperature (°F) 408.493586 408.493586 408.493586 408.493586 408.493586 408.493586 408.493586 95
Number of Flash Drums 6.00 12.00 1.00 10.00 1.00 1.00 1.00 1.00
Total Cost of Vessel ($2010) 15,077,467.40$ 24,982,093.40$ 1,145,707.06$ 16,105,093.61$ 1,236,657.18$ 110,971.38$ 116,036.98$ 47,271.07$
Flash Drum Summary
Flash 1
Vessel Diameter (ft) 3.577339635
Vessel Height (ft) 10.73
Vessel Volume (ft3) 107.8677597
Pressure (psia) 265.409475
Temperature (°F) 408.493586
Number of Flash Drums 1.00
Total Cost of Vessel ($2010) 38,950.49$
Decanter Summary
75
Table 30. Distillation Column Sizing Summary
Table 31. Compressor Sizing Summary
Table 32. Fired Heater Sizing Summary
Distillation 1 Distillation 2 Absorber 1 Stripper 1 Stripper 2
Column Diameter (ft) 25.121 25.8033 13.2 28.64176 25.94875
Column Height (ft) 28 16 14 16 26
Number of Stages 14 8 7 8 13
Column Pressure (psia) 1.1602033 1.1602033 265.409475 14.695949 14.695949
Tray Type Sieve Sieve Sieve Sieve Sieve
Condenser Temperature (°F) 340.316 230 230 230 230
Condenser Duty (Btu/hr) 5137357.3 36306024 0 0 0
Condenser Surface Area (ft2) 432.8270914 5731.23527 0 0 0
Reboiler Temperature (°F) 564 360 400 400 400
Reboiler Duty (Btu/hr) 698,373.88 130.00 90.00 90.00 90.00
Reboiler Surface Area (ft2) 0 3 6 9 12
Reboiler Steam Required (klbm/hr) 0 6.611063346 0 0 0
Reboiler Fuel Gas Required (MBTU/hr) 41.90 6,611.06 - - -
Condenser Cooling Water Required (kgal/hr) 0 117.9709322 0 0 0
Number of Columns 1 1 1 1 2
Total Column Cost ($2010) 5,162,589.48$ 2,331,833.54$ 617,227.81$ 900,839.18$ 1,798,965.39$
Distillation Column Summary
877.89
231,595.27$
695.48
198,510.23$
36,781,846.50
19,851.02$ Cost of Blower ($2010)
Compressor Summary
Compressor 1 Duty (kW)
Cost of Compressor 1 ($2010)
Compressor 2 Duty (kW)
Cost of Compressor 2 ($2010)
Volumetric Flow Rate (ft3/hr)
2,742.46
10,587,212.46$
Fuel Gas Duty (MBTU/hr)
Cost of Fired Heater ($2010)
Fired Heater Summary
76
Table 33. Refrigeration Summary
Table 34. Reactor Sizing Summary
982.56
264,680.31$
Evaporator
Thin (⁰F) 95.00
Thout (⁰F) 20.00
Tcin (⁰F) 10.00
Tcout (⁰F) 10.00
Duty (Q) (Btu/hr) 10,728,433.00
Area (ft2) 2,040.85
Cost of Evaporator 39,702.05$
Condenser
Thin (⁰F) 114.00
Thout (⁰F) 95.00
Tcin (⁰F) 85.00
Tcout (⁰F) 110.00
Duty (Q) (Btu/hr) 13,410,541.25
Area (ft2) 13,653.28
Cost of Condenser 331,260.89$
Cooling Water Necessary (kgal/hr) 253.37
Compressor Duty (kW)
Cost of Compressor ($2010)
Refrigeration Summary
Reactor Diamter (ft) 20
Reactor Lenggth (ft) 40.98
Reactor Pressure (psig) 314.6959488
Tube Outer Diameter (in) 1.25
Number of Tubes per Reactor 9552.367547
Shell Wall Thickness (in) 4.056950235
Weight of Steel Required (lbm) 791,106.70
Cost of Stainless Steel ($2010/lbm) 3.32$
Cost of Reactor ($2010) 2,626,556.57$
Number of Reactors In Parallel 43
Cost of Reactor System ($/2010) 112,941,932.69$
Volume of Catalyst (ft3) 276392.4086
Cost of Catlyst ($/lbm) 10
Catalyst every 4 years (lbm) 13803691.51
Cost of Catalyst every 4 years ($2010) 138036915.1
Cost of Catalyst per year ($/year) 34509228.77
Cooling Water Flow Rate (kgal/hr) 1619.605829
Reactor Summary
77
Table 35. Syngas Unit Process Control Summary Table
Table 36. Fischer-Tropsch Reactor Control Summary Table
Table 37. Distillation Control Summary Table
Expected Disturbance Controlled Variable Manipulated Variable Manipulated Variable Response Type of Valve Fail Safe Alarms
Low inlet temperature Temperature of Inlet Stream Fired Heater Fuel Gas Flow Turn on Fired Heater PI Control Fail Close
High Inlet Temperature Temperature of Inlet Stream Oxygen Flow Rate Decrease Oxygen Flow PI Control Fail Close
Low Operating Temperature Operating Temperature Oxygen Flow Rate Increase Oxygen Flow PI Control Fail Close
High Internal Temperature Operating Temperature Oxygen Flow Rate Decrease Oxygen Flow PI Control Fail Close
H2 to CO ratio below 2 H2 to CO Ratio Carbon Dioxide Flow Rate Decrease Carbon Dioxide Feed PI Control Fail Close
H2 to CO ratio above 2 H2 to CO Ratio Carbon Dioxide Flow Rate Increase Carbon Dioxide Feed PI Control Fail Close
Low Conversion Operating Temperature Oxygen Flow Rate Increase Oxygen Flow PI Control Fail Close
High Conversion Operating Temperature Oxygen Flow Rate Decrease Oxygen Flow PI Control Fail Close
High Pressure Operating Pressure Inlet Flow Rate Open Pressure Release On-Off Control Fail Open High Pressure Alarm
Synthesis Gas Unit Control
Expected Disturbance Controlled Variable Manipulated Variable Manipulated Variable Response Type of Valve Fail Safe Alarms
Low Reactor Temperature Reactor Temperature Shell Side Pressure Increase Shell Side Pressure PI Control Fail Open
High Reactor Temperature Reactor Temperature Shell Side Pressure Decrease Shell Side Pressure PI Control Fail Open
Low Cooling Water Temperature Reactor Temperature Shell Side Pressure Increase Shell Side Pressure PI Control Fail Open
High Cooling Water Temperature Reactor Temperature Shell Side Pressure Decrease Shell Side Pressure PI Control Fail Open
Low Conversion Reactor Temperature Shell Side Pressure Increase Shell Side Pressure PI Control Fail Open
High Conversion Reactor Temperature Shell Side Pressure Decrease Shell Side Pressure PI Control Fail Open
High Feed Temperature Feed Temperature Medium Steam Condensate Flow Rate Increase Steam Condensate Flow Rate PI Control Fail Open
Low Feed Temperature Feed Temperature Medium Steam Condensate Flow Rate Decrease Steam Condensate Flow Rate PI Control Fail Open
High Pressure Reactor Pressure Inlet Flow Rate Open Pressure Release On-Off Control Fail Open High Pressure Alarm
Fischer-Tropsch Reactor Control
Expected Disturbance Controlled Variable Manipulated Variable Manipulated Variable Response Type of Valve Fail Safe Alarms
Low Reboiler Temperature Reboiler Temperature Steam or Fuel Gas Flow Rate Increase Steam or Fuel Gas Flow Rate PI Control Fail Close
High Reboiler Temperature Reboiler Temperature Steam or Fuel Gas Flow Rate Decrease Steam or Fuel Gas Flow Rate PI Control Fail Close
High Cooling Water Temperature Condenser Temperature Cooling Water Flow Rate Increase Cooling Water Flow Rate PI Control Fail Open
Low Cooling Water Temperatuer Condenser Temperature Cooling Water Flow Rate Decrease Cooling Water Flow Rate PI Control Fail Open
High liquid Level Liquid Level Steam or Fuel Gas Flow Rate Increase Steam or Fuel Gas Flow Rate PI Control Fail Close Flood Alarm
Low Liquid Level Liquid Level Steam or Fuel Gas Flow Rate Decrease Steam or Fuel Gas Flow Rate On-Off Control Fail Close Flood Alarm
Distillation Column Control Control
78
Table 38. Flash Drum Control Summary Table
Figure 9. Sensitivity Analysis of the Plant
Expected Disturbance Controlled Variable Manipulated Variable Manipulated Variable Response Type of Valve Fail Safe Alarms
High Inlet Temperature Flash Drum Temperature Cooling Flow Rate Increase Cooling Flow Rate PI Control Fail Open
Low Inlet Temperature Flash Drum Temperature Cooling Flow Rate Decrease Cooling Flow Rate PI Control Fail Open
High Pressure Flash Drum Pressure Inlet Flow Rate Open Pressure Release On-Off Control Fail Open High Pressure Alarm
Flash Drum Control Control
0.00%
20.00%
40.00%
60.00%
80.00%
100.00%
120.00%
50% 60% 70% 80% 90% 100% 110% 120% 130% 140% 150%
DC
FRR
(%
)
Percent Change in Price (%)
Sensitivity Analysis
Raw Materials
Utilities
Fixed Capital
Products
79
Appendix A.4 – Computer Simulation Outputs
Synthesis Gas Unit MATLAB Design Function
function SynGasDesign
clc
function Tinput = T Tinput = 1950; %F end
function Tpre1 = Tpre Tpre1 = 1000; %F end
function Pinput1 = P Pinput1 = 314.6959488; %psia end
function WtoCH4ratio1 = WtoCH4ratio WtoCH4ratio1 = 0.5; end
%Calculation of Enthalpies from NIST function h = H(i,T) Th = ((T-32)*(5/9)+273.15)/1000;
% [CH4 (298-1300) H2O(500-1700) O2(700-2000) CO2(1200-6000) CO(298-
1300) H2(1000-2500) N2(500-2000)] A = [ -0.703029 30.092000 30.032350 24.997350
25.567590 33.0661780 19.505830 ]; B = [ 108.47730 6.8325140 8.7729720 55.186960
6.0961300 -11.363417 19.887050 ]; C = [ -42.52157 6.7934350 -3.988133 -33.69137
4.0546560 11.4328160 -8.598535 ]; D = [ 5.8627880 -2.534480 0.7883130 7.9483870 -
2.671301 -2.7728740 1.3697840 ]; E = [ 0.6785650 0.0821390 -0.741599 -0.136638
0.1310210 -0.1585580 0.5276010 ]; F = [ -76.84376 -250.8810 -11.32468 -403.6075 -
118.0089 -9.9807970 -4.935202 ]; G = [ 158.71630 223.39670 236.16630 228.24310
227.36650 172.707974 212.39000 ]; H = [ -74.87310 -241.8264 0.0000000 -393.5224 -
110.5271 0.00000000 0.0000000 ];
h = (A(i)*Th+(B(i)*Th^2)/2+(C(i)*Th^3)/3+(D(i)*Th^4)/4-E(i)/Th+F(i)-
H(i))*0.94805*453.59237; %BTU/lbmol end
%Heats of Formation from NIST function f = Hf(i)
80
% [ CH4 H2O O2 CO2 CO H2 N2 ] f1 = [ -74.87 -241.83 0 -393.5224 -110.5271 0 0 ]; f = f1(i); end
function c = Cp100(i) % [ CH4 H2O O2 CO2 CO H2
N2 ] c1 = [ 80.49370 44.41339 35.84960 56.79109 34.40867
31.22951 33.96532]; c = c1(i); end
function F = equationsolve(x)
% function x1 = x(i) % x2 = [43400.4031864778 11498.370384958 -6801.2208055209
24581.8491447255]; % x1 = x2(i); % end
%Intial Flow Rate of Methane R = 10.73159; %ft^3-psi/R-lbmol Tstand = 60+459.67; %R Pstand = 14.6959488; %psia CH4Feed = 500000000; %SCF/Day FCH4 = (Pstand*CH4Feed)/(R*Tstand); %lbmol/Day
%Reaction Kinetics from Reference 1 (Gas-To-Liquid) Ksyn = exp(30.53-4.85E4/T+2.42E6/T^2+2.49E9/T^3); Kwater = exp(-2.93+3.61E3/T+5.04E6/T^2+1.82E9/T^3); H2toCOratio = 2; %Desired Hydrogen to CO ratio
%Heat of Reaction Calculation Tref = 298.15; %K Tk = (T-32)*(5/9)+273.15; %K Hrxn1input = 226.1; %kJ/gmol from Reference 1 Hrxn2ref = (Hf(5)+2*Hf(2))-(Hf(1)+(3/2)*Hf(3)); %kJ/gmol Hrxn2input = Hrxn2ref+((Cp100(1)+(3/2)*Cp100(3))*(Tref-
Tk)+(Cp100(5)+2*Cp100(2))*(Tk-Tref))/1000; %kJ/gmol Hrxn3input = -41; %kJ/gmol from Reference 1
Hrxn1 = Hrxn1input*0.94805*453.59237; %Btu/lbmol Hrxn2 = Hrxn2input*0.94805*453.59237; %Btu/lbmol Hrxn3 = Hrxn3input*0.94805*453.59237; %Btu/lbmol
%Inlet Flow Rates FCH4in = FCH4/24 %lbmol/hr FH2Oin = WtoCH4ratio*FCH4in %lbmol/hr FO2in = (3/2)*x(2) %lbmol/hr FCO2in = x(4) %lbmol/hr FCOin = 0; %lbmol/hr FH2in = 0; %lbmol/hr FN2in = (FO2in/0.99)*0.01; %lbmol/hr Ftotalin = FCH4in+FH2Oin+FO2in+FCO2in+FCOin+FH2in+FN2in; %lbmol/hr
81
%Outlet Flow Rates FCH4out = FCH4in-x(1)-x(2) %lbmol/hr FH2Oout = FH2Oin+2*x(2)-x(1)-x(3) %lbmol/hr FO2out = 0; %lbmol/hr FCO2out = FCO2in+x(3) %lbmol/hr FCOout = x(1)+x(2)-x(3) %lbmol/hr FH2out = 3*x(1)+x(3) %lbmol/hr FN2out = FN2in %lbmol/hr Ftotalout = FCH4out+FH2Oout+FO2out+FCO2out+FCOout+FH2out+FN2out; %lbmol/hr
%Oulet mole fraction yCH4 = FCH4out/Ftotalout; yH2O = FH2Oout/Ftotalout; yO2 = FO2out/Ftotalout; yCO2 = FCO2out/Ftotalout; yCO = FCOout/Ftotalout; yH2 = FH2out/Ftotalout; yN2 = FN2out/Ftotalout; Sumy = yCH4+yH2O+yO2+yCO2+yCO+yH2+yN2;
%It is assumed that the second reaction goes to completion and there is no %oxygen remaining in the system
%Energy Balance Setup
Hin =
FCH4in*H(1,Tpre)+FH2Oin*H(2,Tpre)+FO2in*H(3,Tpre)+FCO2in*H(4,Tpre)+FCOin*H(5,
Tpre)+FH2in*H(6,Tpre)+FN2in*H(7,Tpre) %Btu/hr Hout =
FCH4out*H(1,T)+FH2Oout*H(2,T)+FO2out*H(3,T)+FCO2out*H(4,T)+FCOout*H(5,T)+FH2o
ut*H(6,T)+FN2out*H(7,T) %Btu/hr Hrxntotal = x(1)*Hrxn1+x(2)*Hrxn2+x(3)*Hrxn3 %Btu/hr
%Energy Balance eqn1 = Hin+Hrxntotal-Hout
%H2 to CO Ratio eqn2 = FH2out/FCOout-H2toCOratio
%Extent of Reaction 1 Trank = T+459.67; %R eqn3 = ((yCO*yH2^3)/(yCH4*yH2O))-Ksyn
%Extent of Reaction 3 eqn4 = (yH2*yCO2)/(yH2O*yCO)-Kwater
F = [eqn1; eqn2; eqn3; eqn4]
end
x0 = [43400.4; 11498.37; -6801.22; 24581.8];
options = optimset('Display','iter','TolX',1E-6,'MaxIter',4000);
82
[x,fval,exitflag] = fsolve(@equationsolve, x0, options)
%Oxygen Balance OxygenBalance = FH2Oin+FO2in*2+FCO2in*2+FCOin-
(FH2Oout+FO2out*2+FCO2out*2+FCOout)
%Carbon Balance CarbonBalance = FCH4in+FCO2in+FCOin-(FCH4out+FCO2out+FCOout)
end
83
Synthesis Gas Unit MATLAB Design Output
FCH4in =
54899.0304437681
FH2Oin =
27449.5152218841
FO2in =
16300.7992377292
FCO2in =
17784.6908452316
FCH4out =
277.523275520134
FH2Oout =
12769.5760933577
FCO2out =
10444.7212809684
FCOout =
61961.4767325112
FH2out =
123922.953465022
FN2out =
164.654537754841
Hin =
1109755700.2328
Hout =
3012849499.00076
Hrxntotal =
1903093795.463
eqn1 =
-3.30496740341187
eqn2 =
4.44089209850063e-016
eqn3 =
-1.9077166371062e-006
eqn4 =
4.15068710513111e-008
F =
-3.30496740341187
4.44089209850063e-016
84
-1.9077166371062e-006
4.15068710513111e-008
14 75 8.36554e-017 0.000567478 2.6e-008
3.81e+003
Optimization terminated: first-order optimality is less than options.TolFun.
x =
43754.3076764285
10867.1994918195
-7339.96956426318
17784.690580219
fval =
0
4.44089209850063e-016
9.14633346837945e-009
5.10702591327572e-015
exitflag =
1
OxygenBalance =
2.91038304567337e-011
CarbonBalance =
1.45519152283669e-011
85
Fischer-Tropsch Reactor MATLAB Design Function
function FTRDesignFinal
clc
function Tinput = T Tinput = 425; %F end
function Pinput2 = P0 Pinput = 314.6959488; %psia Pinput2 = Pinput*0.068045964; %atm end
function NTinput = NT NTinput = 410000; %Number of Tubes end
function Twinput = Twater Twinput = 405.1; %F end
[W Z] = ode15s(@eqn,[0 14],[0 1 T]);
%Calculation of Heat Capacities from NIST function c = Cp(i2,T) TCp = ((T-32)*(5/9)+273.15)/1000;
% [CH4(298-1300) H2O(500-1700) CO2(298-1200) CO(298-1300) H2(298-
1000) N2(100-500)] A = [ -0.703029 30.092000 24.997350 25.567590
33.0661780 28.986410 ]; B = [ 108.47730 6.8325140 55.186960 6.0961300 -
11.363417 1.8539780 ]; C = [ -42.52157 6.7934350 -33.69137 4.0546560
11.4328160 -9.647459 ]; D = [ 5.8627880 -2.534480 7.9483870 -2.671301 -
2.7728740 16.635370 ]; E = [ 0.6785650 0.0821390 -0.136638 0.1310210 -
0.1585580 0.0001170 ]; F = [ -76.84376 -250.8810 -403.6075 -118.0089 -
9.9807970 -8.671914 ]; G = [ 158.71630 223.39670 228.24310 227.36650
172.707974 226.41680 ]; H = [ -74.87310 -241.8264 -393.5224 -110.5271
0.00000000 0.0000000 ];
c =
(A(i2)+B(i2)*TCp+C(i2)*TCp^2+D(i2)*TCp^3+(E(i2)/(TCp^2)))*(453.59237*(9/5)/10
55.05585); %BTU/lbmol-F end
%Calculation of Hydrocarbon Heat Capacities (CpC) %BTU/lbmol-F function b = CpC(i3) %Note: These Cp values are for 500K
86
if i3 >= 11 b = (34.780167*i3+8.622333)*(453.59237*(9/5)/1055.05585); elseif i3 == 1 b = 46.63*(453.59237*(9/5)/1055.05585); %BTU/lbmol-F elseif i3 == 2 b = 77.94*(453.59237*(9/5)/1055.05585); %BTU/lbmol-F elseif i3 == 3 b = 112.59*(453.59237*(9/5)/1055.05585); %BTU/lbmol-F elseif i3 == 4 b = 148.66*(453.59237*(9/5)/1055.05585); %BTU/lbmol-F elseif i3 == 5 b = 182.39*(453.59237*(9/5)/1055.05585); %BTU/lbmol-F elseif i3 == 6 b = 217.28*(453.59237*(9/5)/1055.05585); %BTU/lbmol-F elseif i3 == 7 b = 252.09*(453.59237*(9/5)/1055.05585); %BTU/lbmol-F elseif i3 == 8 b = 286.81*(453.59237*(9/5)/1055.05585); %BTU/lbmol-F elseif i3 == 9 b = 321.54*(453.59237*(9/5)/1055.05585); %BTU/lbmol-F elseif i3 == 10 b = 356.43*(453.59237*(9/5)/1055.05585); %BTU/lbmol-F end end
function Scorr2 = Scorr(T) Tk = (T-32)*(5/9)+273.15; T4 = exp(250*(1/Tk-1/473)); T3 = exp(-10000*(1/Tk-1/473)); alpha = 0.93*T4; v(1) = 0.03*T3; v(2:4) = 0.04*0.03*T3; for i = 5:1:60 v(i) = (1-alpha)*alpha^(i-1); end Scorr2 = 1/sum(v); end
function S2 = S(n,T) Tk = (T-32)*(5/9)+273.15; T4 = exp(250*(1/Tk-1/473)); T3 = exp(-10000*(1/Tk-1/473)); alpha = 0.93*T4; if n >= 5 S2 = ((1-alpha)*alpha^(n-1))/n*Scorr(T); elseif n >= 2 S2 = (0.04*0.03*T3)/n*Scorr(T); elseif n == 1 S2 = (0.03*T3)/n*Scorr(T); end end
for i10 = 1:1:60 Scheck(i10) = (S(i10,T))*i10; end
87
Scheck = sum(Scheck)
function dz = eqn(w,z)
T1 = exp(-4492*(1/((z(3)-32)*5/9+273.15)-1/473)); T2 = exp(8237*(1/((z(3)-32)*5/9+273.15)-1/473)); T3 = exp(-10000*(1/((z(3)-32)*5/9+273.15)-1/473));
%Inital Flow Rates
FCH4 = 277.523275520134; %lbmol/hr FH2O = 12769.5760933577; %lbmol/hr FCO2 = 10444.7212809684; %lbmol/hr FCO = 61961.4767325112; %lbmol/hr FH2 = 123922.953465022; %lbmol/hr FN2 = 164.654537754841; %lbmol/hr Ftotalin = FCH4+FH2O+FCO+FH2+FCO2+FN2; %Total molar flow (lbmol/hr)
ThetaCH4 = FCH4/FCO; ThetaH2O = FH2O/FCO; ThetaCO = FCO/FCO; ThetaH2 = FH2/FCO; ThetaCO2 = FCO2/FCO; ThetaN2 = FN2/FCO;
%Densities from ASPEN rhoCH4 = 9.307358; %kg/m^3 rhoH2O = 10.45173; %kg/m^3 rhoCO2 = 25.5327; %kg/m^3 rhoCO = 16.25228; %kg/m^3 rhoH2 = 1.169532; %kg/m^3 rhoN2 = 16.25228; %kg/m^3 rhogas
=((FCH4*rhoCH4+FH2O*rhoH2O+FCO2*rhoCO2+FCO*rhoCO+FH2*rhoH2+FN2*rhoN2)/Ftotali
n)*(2.20462262/(3.2808399^3)); %lbm/ft^3
%Viscosities from ASPEN muCH4 = 0.0166811; %cP muH2O = 0.0169866; %cP muCO2 = 0.232632; %cP muCO = 0.025402; %cP muH2 = 0.0124972; %cP muN2 = 0.0254666; %cP mugas =
((FCH4*muCH4+FH2O*muH2O+FCO2*muCO2+FCO*muCO+FH2*muH2+FN2*muN2)/Ftotalin)*0.00
1*((3600*2.20462262)/3.2808399); %lbm/ft-hr
%Molecular Weights from NIST MWCH4 = 16.0425; %lbm/lbmol MWH2O = 18.0153; %lbm/lbmol MWCO2 = 44.0095; %lbm/lbmol MWCO = 28.0101; %lbm/lbmol MWH2 = 2.01588; %lbm/lbmol MWN2 = 28.0134; %lbm/lbmol
88
MWgas =
(FCH4*MWCH4+FH2O*MWH2O+FCO2*MWCO2+FCO*MWCO+FH2*MWH2+FN2*MWN2)/Ftotalin;
%lbm/lbmol
%Delta Calculation for h = 1:60 g(h) = (1+1/h-(2+1))*S(h,z(3))*h; end delta = sum(g); yCO = FCO/Ftotalin; epsilon = yCO*delta; PCO0 = yCO*P0; %atm
StoichH2 = 2; %Depending on Source, 2 or (11/5)
FCOTube = FCO/NT; %lbmol/hr
phi = 0.4; %Void Fraction rhobulk = 0.8*62.428; %Catalyst Bulk Density (lbm/ft^3) rhocat = rhobulk/(1-phi); %Catalyst Density (lbm/ft^3) Doinput = 1.25; %Outter Tube Diamter (inches) Do = Doinput/12; %ft Do2 = Doinput*2.54; %cm Diinput = 1.12; %Inner Tube Diamter (inches) Di = Diinput/12; %ft
Ac = (pi/4)*Di^2; %cross-sectional area of one tube (ft^2) G = (Ftotalin*MWgas)/(Ac*NT*phi); %superficial mass velocity of
thru one tube (lbm/ft^2-h) G2 = G*(453.59237/(30.4862^2)); %g/hr-cm^2
gc = 32.174*3600^2; %Gravitational Constant (lbm-ft/h^2-lbf) Dp = (1/16)/12; %Particle Diameter (ft) k = 0.0173*(453.59237/(453.59237*0.8)); %lbmolCO/hr-lbmcat-atm^2 k2 = 4.512; %atm^-1
Hrxn = 70200; %BTU/lbmolCO a = 4/Do; %Heat exchange per unit volume of reactor (ft^-1)
B0 = (G*(1-phi)/(rhogas*gc*Dp*phi^3))*(150*(1-phi)*mugas/Dp+1.75*G);
%lbf/ft^3 B02 = B0*(1/(144*14.7)); %atm/ft alpha = 2*B02/(Ac*rhocat*(1-phi)*P0); %1/lbm PCO = PCO0*((ThetaCO-z(1))/(1+epsilon*z(1)))*z(2); %atm PH2 = PCO0*(ThetaH2-StoichH2*z(1))/(1+epsilon*z(2)); %atm rCO = k*T1*PH2*PCO/(1+k2*T2*PCO)^2; %lbmolCO/hr-lbmcat
Uo = (0.385*G2^0.8)/Do2^0.2; %BTU/ft^2-hr-F
%ODE45 Differential Equations %z(1) = Conversion %z(2) = Percentage of Pressure Drop %z(3) = Temperature dz = zeros(3,1); dz(1) = (rCO*NT)/FCO; %1/lbmcat
89
dz(2) = -alpha/(2*z(2))*(1+epsilon*z(1))*z(3)/((T-
273.15)*9/5+32)*1.5; %1/lbmcat
SumFCp =
((FCOTube*((ThetaCH4+S(1,z(3))*z(1))*Cp(1,z(3))+(ThetaH2O+z(1))*Cp(2,z(3))+(1
-z(1))*Cp(4,z(3))+... (ThetaH2-
2*z(1))*Cp(5,z(3))+ThetaCO2*Cp(3,z(3))+ThetaN2*Cp(6,z(3))+...
S(2,z(3))*CpC(2)*z(1)+S(3,z(3))*CpC(3)*z(1)+S(4,z(3))*CpC(4)*z(1)+S(5,z(3))*C
pC(5)*z(1)+S(6,z(3))*CpC(6)*z(1)+...
S(7,z(3))*CpC(7)*z(1)+S(8,z(3))*CpC(8)*z(1)+S(9,z(3))*CpC(9)*z(1)+S(10,z(3))*
CpC(10)*z(1)+S(11,z(3))*CpC(1)*z(1)+...
S(12,z(3))*CpC(12)*z(1)+S(13,z(3))*CpC(13)*z(1)+S(14,z(3))*CpC(14)*z(1)+S(15,
z(3))*CpC(15)*z(1)+S(16,z(3))*CpC(16)*z(1)+...
S(17,z(3))*CpC(17)*z(1)+S(18,z(3))*CpC(18)*z(1)+S(19,z(3))*CpC(19)*z(1)+S(20,
z(3))*CpC(20)*z(1)+S(21,z(3))*CpC(21)*z(1)+...
S(22,z(3))*CpC(22)*z(1)+S(23,z(3))*CpC(23)*z(1)+S(24,z(3))*CpC(24)*z(1)+S(25,
z(3))*CpC(25)*z(1)+S(26,z(3))*CpC(26)*z(1)+...
S(27,z(3))*CpC(27)*z(1)+S(28,z(3))*CpC(28)*z(1)+S(29,z(3))*CpC(29)*z(1)+S(30,
z(3))*CpC(30)*z(1)+S(31,z(3))*CpC(31)*z(1)+...
S(32,z(3))*CpC(32)*z(1)+S(33,z(3))*CpC(33)*z(1)+S(34,z(3))*CpC(34)*z(1)+S(35,
z(3))*CpC(35)*z(1)+S(36,z(3))*CpC(36)*z(1)+...
S(37,z(3))*CpC(37)*z(1)+S(38,z(3))*CpC(38)*z(1)+S(39,z(3))*CpC(39)*z(1)+S(40,
z(3))*CpC(40)*z(1)+S(41,z(3))*CpC(41)*z(1)+...
S(42,z(3))*CpC(42)*z(1)+S(43,z(3))*CpC(43)*z(1)+S(44,z(3))*CpC(44)*z(1)+S(45,
z(3))*CpC(45)*z(1)+S(46,z(3))*CpC(46)*z(1)+...
S(47,z(3))*CpC(47)*z(1)+S(48,z(3))*CpC(48)*z(1)+S(49,z(3))*CpC(49)*z(1)+S(50,
z(3))*CpC(50)*z(1)+S(51,z(3))*CpC(51)*z(1)+...
S(52,z(3))*CpC(52)*z(1)+S(53,z(3))*CpC(53)*z(1)+S(54,z(3))*CpC(54)*z(1)+S(55,
z(3))*CpC(55)*z(1)+S(56,z(3))*CpC(56)*z(1)+...
S(57,z(3))*CpC(57)*z(1)+S(58,z(3))*CpC(58)*z(1)+S(59,z(3))*CpC(59)*z(1)+S(60,
z(3))*CpC(60)*z(1))));
dz(3) = ((Uo*a)/rhobulk*(Twater-z(3))+rCO*Hrxn)/SumFCp; %1/lbmcat end
figure(1) plot(W,Z(:,1),'b') xlabel('Weight of Catalyst (lbm)') ylabel('Percent of Conversion (%)') title('Conversion vs. Catalyst Weight') legend('Conversion (%)')
figure(2)
90
plot(W,Z(:,2),'r') xlabel('Weight of Catalyst (lbm)') ylabel('Percent of Pressure (%)') title('Pressure Drop vs. Catalyst Weight') legend('Pressure Drop (%)')
figure(3) plot(W,Z(:,3),'g') xlabel('Weight of Catalyst (lbm)') ylabel('Temperature (°F)') title('Temperature vs. Catalyst Weight') legend('Temperature (°F)')
weight = W(:); conv = Z(:,1); FCOexit = FCO*(1-conv); %Flow Rate of CO (lbmol/hr) FH2exit = FH2-FCO*StoichH2*conv; %Flow Rate of H2 (lbmol/hr) FH2Oexit = FH2O+FCO*conv; %Flow Rate of H2O (lbmol/hr)
figure(4) plot(weight,FCOexit,'b',weight,FH2exit,'r',weight,FH2Oexit,'g') xlabel('Weight of Catalyst (lbm)') ylabel('Molar Flow Rate (lbmol/hr)') title('Molar Flow Rate vs. Catalyst Weight') legend('Carbon Monoxide','Hydrogen','Water')
%Exiting Conditions CatalystperTube = W(end) %lbm Conversion = Z(end,1) PDrop = (1-Z(end,2))*P0*14.695949%psia TotalCatalystWeight = CatalystperTube*NT; %lbm Temperature = Z(end,3) %F Pressure = Z(end,2)*P0*14.695949 %psia;
%Reactor Design CTP = 0.93; %For a One Tube Pass CL = 0.87; %For a 30-60 Equilateral Tri Pitch PT = Do+(1/12); %3inches inbetween tubes with an inch spacing (ft) Ds = 20; %Shell Diamter (ft) A1 = CL*PT^2; %ft^2 TubesperReactor = CTP*(pi*Ds^2)/(4*A1) RequiredLength = CatalystperTube/rhobulk/Ac %ft ReactorsParallel = NT/TubesperReactor ReactorsSeries = RequiredLength/60
HeatTransArea = NT*RequiredLength*2*pi*(Do/2) %ft^2 LMTD = (T-Twater) %F Q = Uo*HeatTransArea*LMTD %Btu/hr HvapH2O = 821.453; %BTU/lbm CoolingWater = Q/HvapH2O %lbm/hr
%Exiting Flow Rates Methaneout = FCO*S(1,T)*Z(end,1) %lbmol/hr H2Oout = FCO*(ThetaH2O+Z(end,1)) %lbmol/hr
91
CO2out = FCO*(ThetaCO2) %lbmol/hr COout = FCO*(1-Z(end,1)) %lbmol/hr H2out = FCO*(ThetaH2-2*Z(end,1)) %lbmol/hr N2out = FCO*(ThetaN2) %lbmol/hr Ethaneout = FCO*S(2,T)*Z(end,1) %lbmol/hr Propaneout = FCO*S(3,T)*Z(end,1) %lbmol/hr Butaneout = FCO*S(4,T)*Z(end,1) %lbmol/hr Naphthaout = FCO*(S(5,T)+S(6,T)+S(7,T)+S(8,T)+S(9,T)+S(10,T))*Z(end,1);
%lbmol/hr Naphthabreak = [FCO*S(5,T)*Z(end,1) FCO*S(6,T)*Z(end,1) FCO*S(7,T)*Z(end,1)
FCO*S(8,T)*Z(end,1) FCO*S(9,T)*Z(end,1) FCO*S(10,T)*Z(end,1)] %lbmol/hr Dieselout =
FCO*(S(11,T)+S(12,T)+S(13,T)+S(14,T)+S(15,T)+S(16,T)+S(17,T)+S(18,T)+S(19,T)+
S(20,T))*Z(end,1); %lbmol/hr Dieselbreak = [FCO*S(11,T)*Z(end,1) FCO*S(12,T)*Z(end,1) FCO*S(13,T)*Z(end,1)
FCO*S(14,T)*Z(end,1) FCO*S(15,T)*Z(end,1) FCO*S(16,T)*Z(end,1)... FCO*S(17,T)*Z(end,1) FCO*S(18,T)*Z(end,1) FCO*S(19,T)*Z(end,1)
FCO*S(20,T)*Z(end,1)] %lbmol/hr Waxout =
FCO*(S(21,T)+S(22,T)+S(23,T)+S(24,T)+S(25,T)+S(26,T)+S(27,T)+S(28,T)+S(29,T)+
S(30,T)+S(31,T)+S(32,T)+...
S(33,T)+S(34,T)+S(35,T)+S(36,T)+S(37,T)+S(38,T)+S(39,T)+S(40,T)+S(41,T)+S(42,
T)+S(43,T)+S(44,T)+S(45,T)+S(46,T)+...
S(47,T)+S(48,T)+S(49,T)+S(50,T)+S(51,T)+S(52,T)+S(53,T)+S(54,T)+S(55,T)+S(56,
T)+S(57,T)+S(58,T)+S(59,T)+S(60,T))*Z(end,1); %lbmol/hr C21_25 = FCO*(S(21,T)+S(22,T)+S(23,T)+S(24,T)+S(25,T)) %lbmol/hr C26_29 = FCO*(S(26,T)+S(27,T)+S(28,T)+S(29,T)) %lbmol/hr C30_35 = FCO*(S(30,T)+S(31,T)+S(32,T)+S(33,T)+S(34,T)+S(35,T)) %lbmol/hr C36_47 =
FCO*(S(36,T)+S(37,T)+S(38,T)+S(39,T)+S(40,T)+S(41,T)+S(42,T)+S(43,T)+S(44,T)+
S(45,T)+S(46,T)+S(47,T)) %lbmol/hr C47plus =
FCO*(S(48,T)+S(49,T)+S(50,T)+S(51,T)+S(52,T)+S(53,T)+S(54,T)+S(55,T)+S(56,T)+
S(57,T)+S(58,T)+S(59,T)+S(60,T))*Z(end,1) %lbmol/hr
%Carbon Balance for i1 = 1:60 CarbonOut(i1) = FCO*S(i1,T)*Z(end,1)*i1; end
CarbonBalance = FCO-(sum(CarbonOut)+COout); CarbonBalancePercent = CarbonBalance/FCO*100
end
92
Fischer-Tropsch Reactor MATLAB Design Function
Scheck =
0.999999999999999
CatalystperTube =
14
Conversion =
0.909737058557914
PDrop =
49.2864790016159
Temperature =
408.493586202037
Pressure =
265.409475022285
TubesperReactor =
9552.3675474669
RequiredLength =
40.972757411597
ReactorsParallel =
42.921296522842
ReactorsSeries =
0.682879290193283
HeatTransArea =
5497404.43848456
LMTD =
19.9
Q =
9504481354.07364
CoolingWater =
11570328.8612661
Methaneout =
4918.69638501423
H2Oout =
69138.2276798971
CO2out =
10444.7212809684
COout =
5592.82514597183
H2out =
11185.6502919433
N2out =
164.654537754841
Ethaneout =
98.3739277002845
Propaneout =
65.5826184668564
Butaneout =
49.1869638501423
Naphthabreak =
Columns 1 through 5
93
903.312630623071 686.289450785425
536.3039801511 427.828349886977 346.710922525941
Column 6
284.485751419763
Dieselbreak =
Columns 1 through 5
235.786166056751 197.05172739996
165.832109974795 140.389436105665 119.459775080001
Columns 6 through 10
102.104164419589 87.6122842094676
75.4383022632208 65.1570325012516 56.4332924050259
C21_25 =
208.583960967908
C26_29 =
91.2247745637791
C30_35 =
73.8309007924141
C36_47 =
53.4343525987551
C47plus =
12.793340144485
CarbonBalancePercent =
1.17427117587818e-014
94
Appendix A.5 – Supporting Graphs
Figure 10. Syngas Temperature Optimization
Figure 11. Syngas Water to Methane Ratio Optimization
$38,000.00
$38,200.00
$38,400.00
$38,600.00
$38,800.00
$39,000.00
$39,200.00
$39,400.00
$39,600.00
$39,800.00
$40,000.00
$40,200.00
1780 1800 1820 1840 1860 1880 1900 1920 1940 1960
Op
era
tin
g C
ost
s p
er
year
($
/ye
ar)
Syngas Operating Temperature (F)
Syngas Operating Costs vs. Temperature
Total ($/hr)
$39,000.00
$40,000.00
$41,000.00
$42,000.00
$43,000.00
$44,000.00
$45,000.00
$46,000.00
$47,000.00
0.4 0.5 0.6 0.7 0.8 0.9 1 1.1
Op
era
tin
g C
ost
s p
er
year
($
/ye
ar)
Water to Methane Ratio
Syngas Operating Costs vs. Water to Methane Ratio
Total ($/hr)
95
Figure 12. Reactor Temperature Optimization
Figure 13. Reactor Cost Optimization
$1,740.00
$1,760.00
$1,780.00
$1,800.00
$1,820.00
$1,840.00
$1,860.00
$1,880.00
$1,900.00
$1,920.00
380 390 400 410 420 430 440 450 460
Re
ven
ue
pe
r ye
ar (
$/y
ear
)
Mill
ion
s
Reactor Operating Temperatuer (F)
Revenue per Year vs. Reactor Temperature
Revenue per Year
$-
$20.00
$40.00
$60.00
$80.00
$100.00
$120.00
$140.00
$160.00
380 390 400 410 420 430 440 450 460
Re
acto
r C
ost
s p
er
year
($
/ye
ar) M
illio
ns
Reactor Temperature (F)
Cost of Reactor ($/year)
Cost of Reactor ($/year)
96
Table 39. Syngas Flow Rates for Temperature Optimization
Table 40. Syngas Cost Optimization for Temperature Optimization
Table 41. Syngas Flow Rates for Water to Methane Ratio Optimization
1800 54,899.03 27,449.52 16,664.50 24,705.53 1,402.95 14,004.82 17,983.18 60,218.43 120,436.86
1850 54,899.03 27,449.52 16,652.00 21,773.50 805.23 13,589.00 14,843.00 61,023.70 122,047.00
1900 54,899.03 27,449.52 16,513.60 19,526.00 468.24 13,180.50 12,392.00 61,565.30 123,130.60
1950 54,899.03 27,449.52 16,300.80 17,784.70 277.52 12,769.60 10,444.72 61,961.50 123,922.95
Syngas Temperature Optimization Flow Rates
FCOout
(lbmol/hr)
FH2out
(lbmol/hr)
FH2Oout
(lbmol/hr)
FCO2out
(lbmol/hr)
FCH4out
(lbmol/hr)T (°F)
FCH4in
(lbmol/hr)
FO2in
(lbmol/hr)
FH2Oin
(lbmol/hr)
FCO2in
(lbmol/hr)
1800 2,472.56$ 33,701.02$ 3,750.15$ 39,923.72$
1850 2,472.56$ 33,675.74$ 3,305.08$ 39,453.38$
1900 2,472.56$ 33,395.85$ 2,963.93$ 38,832.33$
1950 2,472.56$ 32,965.50$ 2,699.61$ 38,137.67$
Syngas Temperature Optimization
FH2Oin ($/hr) FO2in ($/hr) FCO2in ($/hr) Total ($/hr)T (°F)
0.5 54,899.03 27,449.52 16,664.50 24,705.53 1,402.95 14,004.82 17,983.18 60,218.43 120,436.86
0.6 54,899.03 32,939.40 16,357.00 31,211.78 978.70 18,802.00 24,143.00 60,989.00 121,978.00
0.7 54,899.03 38,429.00 15,977.80 37,733.00 727.20 23,618.00 30,327.00 61,577.00 123,154.00
0.8 54,899.03 43,919.00 15,561.00 44,262.00 563.00 28,445.00 36,525.00 62,072.00 124,144.00
0.9 54,899.03 49,409.00 15,125.00 50,796.00 450.97 33,277.00 42,730.00 62,514.00 125,028.00
1 54,899.03 54,899.00 14,675.00 57,332.00 369.00 38,112.00 48,939.00 62,922.00 125,845.00
1.1 54,899.03 60,388.90 14,216.00 63,871.00 308.00 42,950.00 55,151.00 63,309.00 126,619.00
1.2 54,899.03 65,878.80 13,752.70 70,410.00 261.00 47,790.50 61,366.00 63,682.00 127,364.00
1.3 54,899.03 71,368.00 13,284.00 76,950.00 224.00 52,631.00 67,582.00 64,043.00 128,087.00
1.4 54,899.03 76,858.00 12,812.00 83,492.00 194.00 57,472.00 73,799.00 64,397.00 128,794.00
1.5 54,899.03 82,348.00 12,339.00 90,033.00 170.00 62,314.00 80,016.00 64,745.00 129,490.00
FH2out
(lbmol/hr)
FCOout
(lbmol/hr)
FCO2out
(lbmol/hr)
FCH4in
(lbmol/hr)
FO2in
(lbmol/hr)
FCO2in
(lbmol/hr)
Syngas Water to Methane ratio Optimization Flow Rates
WtoCH4
FH2Oin
(lbmol/hr)
FCH4out
(lbmol/hr)
FH2Oout
(lbmol/hr)
97
Table 42. Syngas Costs for Water to Methane Ratio Optimization
0.5 2,472.56$ 33,701.02$ 3,750.15$ 39,923.72$
0.6 2,967.07$ 33,079.16$ 4,737.76$ 40,783.98$
0.7 3,461.55$ 32,312.29$ 5,727.64$ 41,501.48$
0.8 3,956.07$ 31,469.39$ 6,718.70$ 42,144.15$
0.9 4,450.59$ 30,587.65$ 7,710.52$ 42,748.76$
1 4,945.11$ 29,677.61$ 8,702.64$ 43,325.36$
1.1 5,439.62$ 28,749.36$ 9,695.22$ 43,884.20$
1.2 5,934.13$ 27,812.42$ 10,687.80$ 44,434.35$
1.3 6,428.58$ 26,864.55$ 11,680.53$ 44,973.67$
1.4 6,923.10$ 25,910.02$ 12,673.57$ 45,506.69$
Syngas Water to Methane ratio Optimization
WtoCH4 FH2Oin ($/hr) FO2in ($/hr) FCO2in ($/hr) Total ($/hr)
98
Table 43. Material Balance for Syngas and FTR
Species
To Preheater
(lbmol/hr)
To Syngas Unit
(lbmol/hr)
To FTR
(lbmol/hr)
To Separations
(lbmol/hr)
Methane (CH4) 54,899.03 54,899.03 277.52 4,918.70
Water 27,449.52 27,449.52 12,769.58 69,138.23
Oxygen 16,300.80 16,300.80 - -
Carbon Dioxide 17,784.69 17,784.69 10,444.72 10,444.72
Carbon Monoxide - - 61,961.48 5,592.83
Hydrogen - - 123,922.95 11,185.65
Nitrogen 164.65 164.65 164.65 164.65
Ethane (C2H6) - - - 105.11
Propane (C3H8) - - - 70.07
n-Butane (C4H10) - - - 52.56
n-Pentane (C5H12) - - - 193.04
n-Hexane (C6H14) - - - 733.29
n-Heptane (C7H16) - - - 573.04
n-Octane (C8H18) - - - 457.13
n-Nonane (C9H20) - - - 370.46
n-Decane (C10H22) - - - 303.97
n-Undecane (C11H24) - - - 251.94
n-Dodecane (C12H26) - - - 210.55
n-Tridecane (C13H28) - - - 177.19
n-Tetradecane (C14H30) - - - 150.00
n-Pentadecane (C15H32) - - - 127.64
n-Hexadecane (C16H34) - - - 109.10
n-Heptadecane (C17H36) - - - 93.61
n-Octadecane (C18H38) - - - 80.61
n-Nonadecane (C19H40) - - - 69.62
n-Icosane (C20H42) - - - 60.30
C21 - - - 52.36
C22 - - - 45.56
C23 - - - 39.73
C24 - - - 34.72
C25 - - - 30.38
C26 - - - 26.64
C27 - - - 23.38
C28 - - - 20.56
C29 - - - 18.10
C30 - - - 15.95
99
Table 44. Material Balance for Syngas and FTR 2
Species
To Preheater
(lbmol/hr)
To Syngas Unit
(lbmol/hr)
To FTR
(lbmol/hr)
To Separations
(lbmol/hr)
C31 - - - 14.07
C32 - - - 12.43
C33 - - - 10.99
C34 - - - 9.72
C35 - - - 8.61
C36 - - - 7.63
C37 - - - 6.77
C38 - - - 6.01
C39 - - - 5.34
C40 - - - 4.75
C41 - - - 4.22
C42 - - - 3.76
C43 - - - 3.35
C44 - - - 2.98
C45 - - - 2.66
C46 - - - 2.37
C47 - - - 2.11
C48 - - - 1.89
C49 - - - 1.69
C50 - - - 1.51
C51 - - - 1.35
C52 - - - 1.20
C53 - - - 1.08
C54 - - - 0.96
C55 - - - 0.86
C56 - - - 0.77
C57 - - - 0.69
C58 - - - 0.62
C59 - - - 0.56
C60 - - - 0.50
Total (lbmol/hr) 116,598.69 116,598.69 209,540.91 106,062.80
Carbon Balance 72,683.72 72,683.72 72,683.72 72,069.32
Oxygen Balance 95,620.49 95,620.49 95,620.50 95,620.50
100
Figure 14. Graph of Conversion vs. Catalyst Weight in a Single Tube
Figure 15. Graph of Percent of Pressure Drop vs. Weight of Catalyst
0 2 4 6 8 10 12 140
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1
Weight of Catalyst (lbm)
Perc
ent
of
Convers
ion (
%)
Conversion vs. Catalyst Weight
Conversion (%)
0 2 4 6 8 10 12 140.84
0.86
0.88
0.9
0.92
0.94
0.96
0.98
1
1.02
Weight of Catalyst (lbm)
Perc
ent
of
Pre
ssure
(%
)
Pressure Drop vs. Catalyst Weight
Pressure Drop (%)
101
Figure 16. Graph of Temperature vs. Catalyst Weight of a Single Tube
Figure 17. Graph of Molar Flow Rates of Carbon Monoxide, Hydrogen, and Water vs. Catalyst Weight in a
Single Tube
0 2 4 6 8 10 12 14405
410
415
420
425
430
435
440
Weight of Catalyst (lbm)
Tem
pera
ture
(°F
)Temperature vs. Catalyst Weight
Temperature (°F)
0 2 4 6 8 10 12 140
2
4
6
8
10
12
14x 10
4
Weight of Catalyst (lbm)
Mola
r F
low
Rate
(lb
mol/hr)
Molar Flow Rate vs. Catalyst Weight
Carbon Monoxide
Hydrogen
Water
102
Appendix A.6 – Sample Calculations
Equation 1. Steam Reforming Equilibrium Kinetics
[
]
Equation 2. Water-Gas Shift Reaction Equilibrium Kinetics
[
]
Where,
T = Temperature in degrees Fahrenheit
Equation 3. Anderson-Shulz-Flory (ASF) equation
[
( )
]
Or
( )
Where,
α = ASF Chain Growth Parameter
Wn = Relative Weight Fraction of Carbon Number n
Mn = Relative mol fraction of carbon number n
( )
[ (
)]
Where,
T = Reactor Temperature, in Kelvin
Equation 4. Selectivity of Methane
( )
Equation 5. Selectivity of Ethane, Propane, and Butane
( )
( )
[ (
)]
Where,
T = Reactor Temperature, in Kelvin
103
Equation 6. Design Equation for a Packed-Bed Reactor
Where, FCO = Carbon monoxide Molar Flow Rate
Equation 7. Fischer-Tropsch Rate Equation
( )
( )
( )
( )
( )
Where,
k = 0.0173 gmolCO/hr-cm3catalyst-atm
2
k2 = 4.512 atm-1
X = Conversion
PH2 = Partial Pressure of Hydrogen
PCO = Partial Pressure of Carbon Monoxide
PCO0 = Intial Pressure of Carbon Monoxide
y = Percentage of Pressure Drop
yCO0 = Initial Partial Pressure of Carbon Monoxide
δ = Total Change in Moles
Equation 8. Ergun Equation to Model Pressure Drop in Packed-Bed Reactors
( )
( )
( )
[ ( )
]
Where,
T = Temperature
T0 = Initial Temperature
Φ = Porosity
gc = 32.174 lbm-ft/s2-lbf
G = Superficial Mass Velocity
Ac = Cross-Sectional Area
104
Dp = Diameter of Particle in Bed
ρc = Density of Catalyst
μ = Viscosity of Gas
Equation 9. Temperature Gradient Calculation
( )
∑
Where,
Uo = Overall Temperature Coefficient
ρb = Density of bed
Ta = Ambient Temperature
D = Tube Diameter
G = Inlet Gas Mass Velocity
Equation 10. Thickness for Cylindrical Shells
Equation 11. Thickness for Torispherical Heads
Equation 12. Volume of a Torispherical Head
(
)
Where,
P = Pressure of the Vessel
S = Maximum Allowable Working Stress
ri = Inside Radius of Shell Before Corrosion Allowance
Cc = Corrosion Allowance
OD = Outside Diameter of Shell
a = 2 for thicknesses less than 0.0254m and 3 for thicknesses greater than 0.0254m
La = Inside Radius of Head
Equation 13. Cost of Steel for a Pressure Vessel
( ) ( )
Where,
Wv = Weight of Steel
105
Equation 14. Maximum Vapor Velocity for a Flash Drum
[
]
Where,
ρL = Liquid Density
ρV = Vapor Density
k = 0.35
Equation 15. Number of Tubes in a Shell
Where,
Nt = Number of Tubes
CTP = 0.93 for One Tube Pass
CL = 0.87 for 30&60 Equilateral Triangular Pitch
PT = Tube Pitch
Ds = Diameter of Shell
Depreciation
Future Value
( )∑( )
Present Value
( )
Net Present Value
( )
106
Revenue
( )( )
Gross Profit
Net Profit
( )
Cash Flow
Borrowed Working Capital
( )( )
∑( )
Return on Investment
Cash Flow Payback Period
Inflation
[ ( ) ( ) ]( )
Where,
dj = Depreciation
φ = Tax Rate
s = Revenue
c = Manufacturing Costs
j = Year Number
i = Inflation Rate