This design final report summarizes a gas-to-liquids synthesis process found to be profitable. Raw materials including methane, steam, oxygen and carbon dioxide were converted via syngas, Fischer-Tropsch, separation and hydrocracking units to produce liquid hydrocarbons. The 15-year, 300-day/year plant was determined to have an NPV of $6 billion, IRR of 466%, and payback period of 23 days. Sensitivity analysis found sales most impacted NPV. Heat integration between units and annual credits improved economics. The process was recommended but maintaining costs of sales accuracy with feed input alterations.
Detailed account of Metaphysical Operations required for the The New Age, for which we had to depend on outside help.
The Time Line was altered,and the 2012 event which would have seen Earth bereft of humanity was changed completely.
These details form the the main message, the rest is about Spiritual Philosophy and the Nature and Destiny of Man.
Detailed account of Metaphysical Operations required for the The New Age, for which we had to depend on outside help.
The Time Line was altered,and the 2012 event which would have seen Earth bereft of humanity was changed completely.
These details form the the main message, the rest is about Spiritual Philosophy and the Nature and Destiny of Man.
Energy Systems Optimization of a Shopping Mall: The present study focuses on the development of software (general mathematical optimization model) which has the following characteristics:
• It will be able to find the optimal combination of installed equipment (power & heat generation etc) in a Shopping Mall (micro-grid)
• With multi-objective to maximize the cost at the same time as minimizing the environmental impacts (i.e. CO2 emissions).
• To date, this tool is scarce to the industry (similar to DER-CAM, Homer).
Energy Systems Optimization of a Shopping Mall: The present study focuses on the development of software (general mathematical optimization model) which has the following characteristics:
• It will be able to find the optimal combination of installed equipment (power & heat generation etc) in a Shopping Mall (micro-grid)
• With multi-objective to maximize the cost at the same time as minimizing the environmental impacts (i.e. CO2 emissions).
• To date, this tool is scarce to the industry (similar to DER-CAM, Homer).
Pressure Vessel Selection Sizing and Troubleshooting Karl Kolmetz
Vessels are a vital part of the operational units in the process industries. A vessel is
a container in which materials are processed, treated, or stored. Without this type of
equipment, the process industries would be unable to create and store large
amounts of Product. Pressure vessels used in industry are leak-tight pressure
containers, usually cylindrical or spherical in shape, with different head
configurations.
The process engineer should have some knowledge of the mechanical design of
vessels. For example, the process engineer may have to make a preliminary design
of vessels for a cost estimate. A vessel consists of a cylindrical shell and end caps,
called heads. For safety, vessel design is governed by codes.
This report consists of the analysis of results of evaluation of a 3inch diameter silicon wafer that was fabricated in the clean room at USC under Professor. Kaviani. The wafer consists of resistors, capacitors, MOSFETs and diodes. The device was tested and the results are used to characterize the device. The whole process was done in 100 class clean room, the Powell Foundation Instructional Laboratory. This report will show the calculations performed to do an analysis of the results and will aim to offer an insight into the theory behind the operation of these devices.
Internship Report: Interaction of two particles in a pipe flowPau Molas Roca
The present document sums up the development and results of the research internship carried out at LEGI Laboratory. The study aimed to understand the hydrodynamic forces involvement in the interaction between two red blood cells located in a capillary (pipe flow). The problem regarding Red Blood Cells (RBCs) moving through a capillary has been tackled from a two-dimensional point of view and has been both analytically and numerically outlined. Finite elements have been used to discretize the geometries considered. Several boundary conditions and geometries were simulated and deeply examined aiming to understand the mechanism governing hydrodynamic attraction and repulsion between red blood cells. The consequent results are analyzed in this report.
Internship Report: Interaction of two particles in a pipe flow
Design Final Report
1.
Design Final Report
Senior Design: CHEN 4120
10 December 2013
Michael Oseth, Jack Stringer, & Sarah Waldner
UNIVERSITY OF COLORADO AT BOULDER
2. 1
Contents
Table of Figures ..................................................................................................................................1
Table of Tables ...................................................................................................................................2
Executive Summary ............................................................................................................................4
Introduction .......................................................................................................................................5
Background ........................................................................................................................................6
Safety, Environmental, & Health Concerns ......................................................................................... 10
Project Premises............................................................................................................................... 14
Approach and Process Description..................................................................................................... 15
Syngas Unit Design........................................................................................................................ 15
Fischer-‐Tropsch Reactor Design and Optimization........................................................................... 17
Separation Unit Implementation.................................................................................................... 28
Hydro-‐isomerization Unit .............................................................................................................. 34
Heat Integration ........................................................................................................................... 37
Equipment Specifications .................................................................................................................. 43
Utility Summary................................................................................................................................ 47
Economic Analysis ............................................................................................................................ 51
Equipment Cost ............................................................................................................................ 51
Profitability Analysis ......................................................................................................................... 54
Profitability .................................................................................................................................. 54
Estimation of the Total Capital Investment.................................................................................. 54
Estimation of Operating Costs and Sales ..................................................................................... 55
Calculation of NPV, ROI, and PBP................................................................................................ 56
Sensitivity Analysis ........................................................................................................................ 57
References ....................................................................................................................................... 59
Appendices ...................................................................................................................................... 60
Appendix A: Full Process Flow Diagram .......................................................................................... 60
Appendix B: Separation Unit Stream Tables .................................................................................... 61
Appendix C: Economics Excel Spreadsheets .................................................................................... 64
Table of Figures
Figure 1: A PFD of the Syngas Unit ..................................................................................................... 16
Figure 2: Tornado plot for the sensitivity analysis of overall reactor cost.............................................. 25
3. 2
Figure 3: Tornado plot for the sensitivity analysis of the conversion of the reaction ............................. 25
Figure 4: Tornado plot for the sensitivity analysis of the average carbon chain length of the products... 26
Figure 5: A PFD of the FTR unit .......................................................................................................... 27
Figure 6: PFD for the Separation unit ................................................................................................. 32
Figure 7: Hydroisomerization unit PFD ............................................................................................... 36
Figure 8: Diagram of the Temperature Interval Method ...................................................................... 39
Figure 9: Final Heat integration Diagram ............................................................................................ 41
Figure 10: Process flow diagram of the integrated heat streams.......................................................... 43
Figure 11: Sensitivity Analysis on the NPV .......................................................................................... 58
Figure 12: PFD of Entire Process ........................................................................................................ 60
Table of Tables
Table 1: Personal Equipment............................................................................................................. 10
Table 2: Composition of Waste Streams............................................................................................. 10
Table 3: Material and Safety Data for all Components ........................................................................ 12
Table 4: The stream tables and mass balance of the Syngas unit ......................................................... 16
Table 5: First set of FTR Optimization Parameters............................................................................... 23
Table 6: Table of FTR Optimized Parameters ...................................................................................... 24
Table 7: Stream Tables for the FTR Unit ............................................................................................. 27
Table 8: Inlet and Outlet Stream Tables for the Separation Unit .......................................................... 33
Table 9: Mass Balance for the Separation Unit. .................................................................................. 33
Table 10: Mass Flow rates of Hydrocarbons from the Hydroisomerization Unit .................................... 34
Table 11: A Summary of the Separation in Distillation Column 1.......................................................... 35
Table 12: A Summary of the Separation in Distillation Column 2.......................................................... 35
Table 13: Stream tables and Mass Balance for the Hydroisomerization Unit......................................... 36
Table 14: Streams Considered for Heat Integration............................................................................. 38
Table 15: Calculated Enthalpies for the Temperature Interval Method ................................................ 40
Table 16: Flash Column Equipment Specifications .............................................................................. 44
Table 17: Distillation Column: Condenser/Top Stage Performance ...................................................... 44
Table 18: Distillation Column: Reboiler/ Bottom Stage Performance.................................................... 44
Table 19: Distillation Column Equipment Specifications ...................................................................... 44
Table 20: Compressor Equipment Specifications................................................................................. 44
Table 21: Decanter Equipment Specifications ..................................................................................... 45
Table 22: Heat Exchanger Equipment Specifications ........................................................................... 45
Table 23: Valve Equipment Specifications .......................................................................................... 45
Table 24: Syngas Reactor Specifications ............................................................................................. 46
Table 25: HI Distillation Column: Condenser/Top Stage Performance .................................................. 46
Table 26: HI Distillation Column: Reboiler/ Bottom Stage Performance................................................ 46
Table 27: HI Distillation Column Specifications ................................................................................... 46
Table 28: HI Compressor Equipment Specifications ............................................................................ 46
Table 29: FTR Equipment Specifications ............................................................................................. 47
Table 30: Utility Summary: Heat integration....................................................................................... 47
4. 3
Table 31: Utility Summary for Syngas Unit.......................................................................................... 48
Table 32: Utility Summary for Separation Unit.................................................................................... 49
Table 33: Utility Summary for Air Separation Unit .............................................................................. 49
Table 34: Utility Summary for HI Unit................................................................................................. 50
Table 35: Utility Summary for the HI Unit........................................................................................... 50
Table 36: Equipment and Installation Costs for Flash Drums and Heaters............................................. 51
Table 37: Equipment and Installation Costs of Other Separation Equipment ........................................ 52
Table 38: Bare-‐Module Costs and Total Capital Investment ................................................................. 53
Table 39: Investment Summary ......................................................................................................... 54
Table 40: Associated Credits and Costs for Plant Operation ................................................................ 55
Table 41: Sales and Costs for the First Three Years of Plant Operation ................................................. 56
Table 42: NPV, ROI, PBP Values ......................................................................................................... 56
Table 43: Separation Unit Stream Table ............................................................................................. 61
Table 44: Bare Module Costs ............................................................................................................. 64
Table 45: Utilities ............................................................................................................................. 65
Table 46: Profitability Analysis........................................................................................................... 65
5. 4
Executive Summary
The gas to liquids synthesis process described in this report was found to be a profitable
venture. Syngas, FTR, separation, and hydrocracking units were used to convert raw materials, including
methane, to hydrocarbons as fuel. Implementing a grass-‐roots plant, with mixed indoor and outdoor
process facilities produced an overwhelming profit in the production of liquid hydrocarbons: LPG, diesel,
and naptha. Data indicated that the production process was favorable for investors even at high interest
rates, with a profit occurring after the first year of plant operation.
Raw materials of methane, steam, oxygen, and carbon dioxide obtained from locations ideally
located in close proximity to the plant were inputted as follows. Methane feed was fed at a rate of 500
MSCFD, steam at 12 klb/day, carbon dioxide at 34 MSCFD, and oxygen at 100 ton/day. Specifically, a 15-‐
year, 300 day per year plant life was interpreted to produce naptha product at a rate of 10 barrels/day,
diesel product at a rate of 35 barrels/day, and LPG product at a rate of 640 lb/day.
These production rates equate to an annual net earnings of $770,000,000 at full plant capacity,
indicating a net present value at the end of the 15 year period of $6 billion for an interest rate of 8.5%,
using straight line depreciation. The pay-‐back period at full plant capacity was found to be 23 days when
total depreciable capital was $50 million. Return on investment tremendously exceeded the target of
20%, and at full plant capacity, equated to a return of 840%, given a total capital investment of $92
million. The internal rate of return was determined to be 466%.
Sensitivity analysis indicated that sales chiefly determined the net present value. Heat
generated in the syngas unit was integrated into the separation unit as steam, with left-‐over quantities
generating a year to year credit. Bare-‐module costs for equipment used was determined from ASPEN
PLUS software. A recommendation for further calculations is to maintain the accuracy of the cost of
sales when altering raw feed inputs.
6. 5
Introduction
The plant designed and evaluated by this report is a gas to liquid fuels production facility that
employs a Fischer-‐Tropsch Reactor for the conversion of synthetic gas (syngas) to liquid hydrocarbons.
Syngas is produced via auto-‐thermal reforming technologies where steam and methane gas are
reformed to produce a reactant stream that is mainly hydrogen gas and carbon monoxide. This works by
combining steam methane reforming and partial oxidation technologies. This will be described more in-‐
depth in the background section. The Syngas Unit was designed and tested using ASPEN PLUS programs.
The reactant stream is then sent to the Fischer-‐Tropsch Reactor where hydrogen and carbon monoxide
are catalytically reacted to form fuel. The FT reactor was designed and optimized using MATLAB. This
optimization process is discussed in detail in the Approach and Process Description section. The product
stream from the FT unit is then sent to the Separation Unit. The product stream from the FT unit
contains hydrocarbons of carbon length one to sixty. They can be divided based on carbon chain into the
following categories: methane, ethane, propane, butane, naptha, diesel, and carbon chain length 21+.
The stream is separated into naptha, diesel, and wax. Wax is sent to the Hydroisomerization Unit and
to produce naptha, diesel, and LPG gas (methane butane) which are sold as product.
The Separation Unit was designed and optimized in ASPEN PLUS by an external group called the NERDs.
The Hydroisomerization Unit was also designed in ASPEN PLUS. Both of these units are discussed in
detail in the Approach and Process Description section. Safety, Environmental, and Health concerns are
discussed in the Safety, Environmental, & Health Concerns section. All equipment involved in the
process was designed (sized, unit operations, material of construction) using ASPEN PLUS. These details
are discussed in the Equipment Specifications section and the cost of individual equipment components
are located in the Equipment Cost Section. An Economic Analysis was performed and the total capital
investment, operating costs, and sales were estimated and the NPV, ROI, and PBP values were
calculated for the life of the plant; all values are located in the Profitability Analysis section. A sensitivity
7. 6
analysis was performed to explore how certain variables effect the final NPV of the plant. This analysis is
located in the Sensitivity Analysis section.
Background
As crude oil reserves near exhaustion around the world, gas to liquid fuel production via
synthetic gas has increased in viability. The Fischer-‐Tropsch process, which has been around for nearly
ninety years, provides a feasible option for the production of these fuels.(1)
The FT process involves the
synthesis of synthetic gas from hydrocarbon feedstock (natural gas, coal, naptha, petroleum coke, and
biomass) and catalytic conversion of this gas into liquid fuel. Franz Fischer developed the process in
plants were built and operating in Germany, however, the plants were closed shortly thereafter as the
FT process was not economically practical at the time.(2)
The FT production of hydrocarbons is only
practical if the price per barrel of crude oil is low enough for FT production of fuel to be similarly
affordable. Interest in the FT process disappeared after World War II, but resurfaced during the oil crisis
(2)
A company called Sasol in South Africa constructed the two largest FT complexes ever
built and the company flourished
the next two decades the price of crude oil per/barrel continued to fluctuate and the number of FT
plants slowly increased across the world.(3)
Syngas is the mixture of hydrogen and carbon monoxide gas and can be generated by different
technologies. Generation of syngas usually encompasses 60% -‐ 70% of the total capital investment
required to implement a grass-‐roots FT plant.(4)
It is preferable to use methane gas when it is available
rather than coal as it is less expensive, more efficient, and leaves a smaller carbon footprint. The
technologies for syngas generation: catalytic steam methane reforming (SMR), heat exchange
reforming, partial oxidation (POX) and auto-‐thermal reforming (ATR) will be reviewed. The most
8. 7
extensively used in industry is steam methane reforming (SMR) where steam and methane gas are
catalytically converted to syngas. This process is advantageous as it requires no oxygen and has the
lowest process temperature. However, the H2/CO ratio is usually higher than the optimal ratio for fuel
production ( > 4)and it produces high air emissions. Heat exchange reforming uses heat recovered from
the product syngas as a portion of the heat required for the heat of reaction. It is more compact and
efficient than other technologies, reduces the plant footprint and capital cost. One disadvantage of this
technology is that it is not usually implemented as the sole syngas generator and must be used in
tandem with another process. Partial oxidation generates syngas via the highly exothermic, non-‐catalytic
reaction of methane and steam. While this technology does not require a catalyst, it does require an
oxygen feed. This process is generally not implemented alone as it produces a low H2/CO ratio (< 2) and
requires high operating temperatures. The auto-‐thermal reforming technology combines the SMR and
the POX process to produce a syngas with a favorable H2/CO ratio. This technology combines the partial
oxidation of reactants fueled by the internal combustion of some of the feedstock. While this
technology may seem like the most efficient option, it has experienced limited commercial use. This
technology has been shown to be the least expensive option that fulfills the syngas composition
required for FT processing and commercial use of the reforming technology is increasing.(4)(5)(6)
ATR is the
syngas generation technology discussed in this report.
The FT reactor catalytically converts the syngas into a stream of hydrocarbons. This reaction is
highly exothermic and requires cooling water to control the temperature rise. FT reactors can be divided
into two general categories: high temperature FT and low temperature FT. High temperature FT (HTFT)
reactors (300 350 ) are mainly used to produce olefins and gasoline. Low temperature FT (LTFT)
reactors (200 240 are used for diesel and linear wax production. FT reactors may also produce a
small amount of alcohols, aldehydes, carboxylic acids, and other oxygenated products, and in addition
those operating at high temperatures may produce minute amounts of ketones and aromatic
9. 8
compounds. Wax products are sent to a hydro-‐
they are reduced to naptha and diesel products.(7)(8)
Catalysts that have been considered for FT reactors
include: Ni, Co, Fe, and Ru. Ni easily hydrogenates which produces high methane concentrations in the
product stream. Use of Ni would also require high operating temperatures to avoid the formation of
nickel carbonyls which dissipates the catalyst in the reactor during operation. For both of these reasons
Ni catalysts are not used commercially. Ru catalysts are rare and extraordinarily expensive which makes
it a non-‐viable option for commercial facilities. Co and Fe catalysts are both readily available. On
average, Co is about 200 times the price of Fe, but exhibits an FT activity per metal site 3 times greater
than Fe. This means that one particle of Co can catalyzes the conversion of three times as many
methane molecules as Fe particles.(9)
Despite the high cost, Co catalysts are still widely used. It has been
shown that the partial pressure of the water content of the product stream from a reactor that uses Fe
catalyst has a debilitating effect on the reaction kinetics. Co catalysts experience little to no effect. FT
reactors commonly use iron catalysts and those operating at low temperatures use iron or cobalt
catalysts if the product stream has high water content.(10)
ame a reality within the last ten years as
companies are extracting oil at a faster pace today than ever before. Other options have had to be
considered so that the transition from crude oil dependence to alternative forms of energy is as smooth
as possible. Over the last decade FT technology has experienced increased attention. The technology
remains more costly than processing and using crude oil. In order for FT technology to become
competitive with to crude oil while it is still available, it needs to be made more cost effective. Studies
have estimated that the FT process became a viable option when the price of crude oil per barrel
surpassed $16/barrel. Today the cost of crude oil is approximately $87/barrel. (11)
Even though this technology has been around for almost a century, it is still being developed
and can be improved in a number of ways. One syngas generation method that is currently being
10. 9
explored is the combination of heat exchange reforming with auto-‐thermal reforming. By using heat
recovered from the exit gas this eliminates the need for a gas fired heater to supply energy to the auto-‐
thermal reformer. Potential benefits include a decrease in oxygen requirements and reduced plant
footprints.(11)
Co and Fe catalysts used in Ft reactors can also be improved. It is desirable to design a
catalyst with increased selectivity to hydrocarbons of shorter chain lengths. This would reduce the utility
of the hydro-‐isomerization unit. Coated catalysts are currently being researched as a method of
increasing the selectivity of a catalyst. Recent research has been done using alkanethiol coated catalysts.
The catalysts are prepared by saturating the catalyst in the liquid alkanethiol and allowing the sulfur
atom of the molecule to bind to the catalyst surface. These molecules form a self-‐assembled monolayer
on the surface of the catalyst. The coating changes the surface chemistry of the catalyst through
electronic and steric effects which influences the selectivity of the reaction. It has been shown that by
coating catalysts with alkanethiols and then using these catalysts in hydrogenation reactions, the
selectivity to the desired product increased substantially. However, life of these catalysts is short and
the coatings are have been shown to deteriorate quickly, reducing the heightened selectivity.(12)
Research is still in its preliminary stages, the potential to design catalysts to be selective to particular
products poses an interesting development for FT technology.
Today several large companies have accepted the challenge of enabling economic exploitation
of the coal and natural gas reserves around the world via the Fischer-‐Tropsch Process. The largest
development of FT plants are located in South Africa and operated by Sasol. These plants have proved
profitable are there is little oil in South Africa, but an abundance of coal reserves. Sasol has been slowly
building a FT presence in the United States and are currently developing in Louisiana. PetroSA is another
South A
South Africa. Shell owns a massive FT complex in Bintulu, Malaysia and is currently using natural gas
there to produce diesel fuels and low-‐grade wax. (4)(6)(9)
11. 10
Safety, Environmental, & Health Concerns
Employee Safety Precautions
All employees on shift at the plant must be fully equipped the required personal safety equipment. The
employer is responsible for the adequacy of such equipment. The employer shall ensure that each
employee wears the equipment listed in Table (1) when working in the concerned areas as defined by
OSHA:
Table 1: Personal Equipment
Personal Equipment Area of Concern
Protective Helmet Potential for head injury due to falling objects
Safety Glasses and/or detachable side
protectors
Potential for hazardous flying particles such as
molten metal, liquid chemicals, acids, caustic
liquids, chemical gases or vapors, and injurious
light radiation
Protective footwear Potential for foot injury via falling or rolling
objects, objects piercing the sole, or electrical
hazards
Protective Handwear Potential skin adsorption of harmful substances,
severe cuts, lacerations, abrasions, punctures,
chemical burns and thermal burns
Respirator Potential of inhalation of harmful dusts, fogs,
fumes, mists, gases, smokes, sprays, or vapors
Protective Clothing Worn at all times. Must be non-‐flammable,
protect against electric shock, and non-‐reactive.
Waste Stream Considerations
The process encompasses three main waste streams. The flue gas from strippers 1 and 2 and the waste
water from stripper 2. The compositions of each of these streams are located in Table (2) below.
Table 2: Composition of Waste Streams
Component Mole Flow
(lbmol/day)
Flue Gas 1 Flue Gas 2 Waste H2O
CO 4.66 0.00280 0
CO2 11.5 0.0426 0
H2 5.96 0.000366 0
Water 30.6 26.6 2.80
N2 0.549 2.10 0.00104
CH4 1.48 0.00183 0
12. 11
Ethane 0.0417 0.000256 0
Propane 0.0411 0.000644 0
Butane 0.040 0.000948 0
Naptha 0.130 0.00000106 0
Diesel 0.0082 0 0
C21-‐C25 0 0 0
C26-‐C29 0 0 0
C30-‐C35 0 0 0
C36-‐C47 0 0 0
C48+ 0 0 0
Oxygen 0 7.91 0.00447
Handling Instructions:
Flue Gas 1: Waste gas must be collected, condensed, and properly stored for waste disposal. CO,
CO2, methane, ethane, propane, butane, naptha, and diesel must be removed before waste
water can be safely disposed off. Naptha and diesel can be separated out using adsorption
techniques. Methane, ethane, and propane can be removed using the difference in boiling
points from water. CO2 and CO can be removed using a series of strippers.
Flue Gas 2: Will be treated the same as Flue Gas 2
Waste H2O: This waste water stream contains only nitrogen gas, which is an inert, and can be
disposed of in the sanitary sewer system of the plant
State & Federal Permits
Permits that must be obtained for operation of this plant include the following (Note: Permits may vary
depending on the state that the plant is in):
Building permits for land
Oil refinery permits
Hazardous Waste Permit for disposal in sanitary sewer systems, air pollution, and incineration
permits if incineration is the preferred waste disposal method. Depends on State.
Personal work permits will be required for those doing hot work (welding, cutting, grinding, and
spark producing work), those working with line break (Liquid and gaseous chemicals, and sewer
13. 12
and process water), those needing confined space entry, and those performing heavy lift
equipment operation. Depends on State.
Material and Safety Data for all Process Components
Table 3: Material and Safety Data for all Components
Component Physical
State
MW
(g/mol)
Health Flammability Reactivity Special
Methane Gas 16.04 1 4 0 Simple
Asphyxiant
Water Gas 18.02 0 0 0 None
Oxygen Gas 32.00 0 0 0 Oxidizer
Nitrogen Gas 28.02 0 0 0 Simple
Asphyxiant
Carbon
Dioxide
Gas 44.01 1 0 0 None
Carbon
Monoxide
Gas 28.01 3 4 0 None
Hydrogen Gas 2.02 0 4 0 None
Ethane Gas 30.08 1 4 0 None
Propane Gas 44.11 1 4 2 None
Butane Gas 58.14 1 4 0 None
Pentane Liquid 72.15 1 4 0 None
Hexane Liquid 86.18 1 3 0 None
Heptane Liquid 100.21 1 3 0 None
Octane Liquid 114.23 2 3 0 None
Nonane Liquid 128.26 2 3 0 None
Decane Liquid 142.28 0 2 0 None
Undecane Liquid 156.31 0 2 0 None
Dodecane Liquid 170.34 1 2 0 None
Tridecane Liquid 184.37 1 2 0 None
Tetradecane Liquid 198.4 2 1 0 None
Pentadecane Liquid 212.42 1 1 0 None
Hexadecane Liquid 226.00 0 1 0 None
Heptadecane Liquid 240.48 2 1 0 None
Octadecane Liquid 254.5 0 1 0 None
Nonadecane Liquid 268.52 2 1 0 None
Eicosane Liquid 282.56 1 1 0 None
*Hydrocarbons of length 21 60 carbons can adhere to the same safety data as Eicosane as the
information does not vary significantly.
14. 13
Safe plant operating procedures
For general safety of all employed by the plant the following protection/accident prevention
equipment must be installed, maintained, and supervised at all times:
Temperature/Pressure indicator instruments
Sound regulation system and emergency shut off
System leak alarm system
For general safety of all employed by the plant the following must be adhered to at all times:
Employees will possess appropriate personal protection equipment
Plant must remain well lit to ensure proper visibility
All safety instrumentation and alarm systems must be monitored
Employees must not work longer in a high stress environment than time defined by OSHA
standards for the nature of that environment
All equipment must be regularly cleaned and maintained
Waste must be disposed of properly adhering to OSHA standards
All working personnel must be properly trained in the following operating procedures and it is the
responsibility of the employer to keep employees up to date with the following operating procedures
and changes in such:
1. Normal Operating Procedures
2. Abnormal Operating Procedures
3. Operator Response Procedures
4. Emergency Operating Procedures
15. 14
Project Premises
Design Premises:
-‐
Location: In the vicinity of a large body of water to naturally supply cooling water and a
Hydraulic Fracturing Rig
Material Source:
o Cooling Water: Obtained from nearby reservoir
o Oxygen: Air Separation Plant
o Methane: Nearby Hydraulic Fracturing Rig
o Carbon Dioxide: External import
o Electricity: Xcel; some heat obtained from recycling energy within the process
Plant Capacity:
o Naptha production: 10 barrels/day
o Diesel Production: 35 barrels/day
o LPG production: 640 lbs/day
Recycle Steams:
Economic Premises:
Cost of Raw materials:
o Methane: $2/MSCFD
o Steam: $5/klb
o Carbon Dioxide: $0.40/MSCFD
o Oxygen: $100/ton
Sales price of products:
o Naptha: $75/barrel
o Diesel: $90/barrel
o LPG: $0.30/lb
Project Life: 15 years
Depreciation Method: Straight line, 15 yr
Total Capital Investment: $ 91,500,000
Tax Rates: 37%
Targeted ROI: 20%
Operation: 330 days/yr
16. 15
Approach and Process Description
Syngas Unit Design
The first part of the overall design of the Fischer-‐Tropsch Reaction unit was to design the syngas
unit. Carbon dioxide, methane, steam, and oxygen are combusted in an equilibrium reactor to produce
Syngas. The three primary reactions that occur simultaneously in the reactor are as follows.
Steam reforming:
Partial oxidation of methane:
Shift reaction:
Two process units are involved in the Syngas unit. Four streams (carbon dioxide, oxygen (99%
oxygen and 1% nitrogen), methane and steam) are fed into a mixer at 100 . The methane is fed in with
steam in order to prevent coking of the process equipment. The material is then fed directly into the
equilibrium reactor. The only component flow-‐rate that was definitively defined was methane. The
component flow rates of carbon dioxide and air were given arbitrary values because these flow-‐rates
were optimized using the optimization package in Aspen to optimize the production of hydrogen.
There are three constraints for the system. First, the ratio of hydrogen production to carbon
monoxide production must be 2:1. Second, the equilibrium reactor is adiabatic. Third, the molar ratios
of steam to methane flow rates is greater than or equal to 1:2. The parameters being varied using the
optimization package are the temperature (1600 1950 ) and pressure (300 500 psig) of the reactor,
the flow rates of air (10 1,000 MSCFD) , carbon dioxide (10 -‐1,000 MSCFD), and steam (250 -‐500
MSCFD). The product of the Syngas unit was then feed into a Fischer Tropsch Reactor. Figure (1) is the
process flow diagram of the Syngas unit and Table (4) provides the stream tables and mass balance of
the unit.
17. 16
MIXER 1
SYNGAS REACTOR
CO2
AIR
STEAM
C1
FEED SYNGAS
Figure 1: A PFD of the Syngas Unit
Table 4: The stream tables and mass balance of the Syngas unit
Units AIR CO2 FEED METHANE STEAM SYNGAS
From MIXER SYNUNIT
To MIXER MIXER SYNUNIT MIXER MIXER
Substream: MIXED
Phase: Vapor Vapor Mixed Vapor Liquid Vapor
Component Mole
Flow
METHA-‐01 MSCFD 0.00 0.00 500.00 500.00 0.00 4.48
WATER MSCFD 0.00 0.00 250.00 0.00 250.00 394.01
NITRO-‐01 MSCFD 3.59 0.00 3.59 0.00 0.00 3.59
OXYGE-‐01 MSCFD 355.77 0.00 355.77 0.00 0.00 0.00
CO MSCFD 0.00 0.00 0.00 0.00 0.00 423.52
CO2 MSCFD 0.00 34.38 34.38 0.00 0.00 106.38
HYDRO-‐01 MSCFD 0.00 0.00 0.00 0.00 0.00 847.04
Mole Flow MSCFD 359.36 34.38 1143.74 500.00 250.00 1779.02
Mass Flow LB/DAY 30264.26 3987.18 67257.43 21137.67 11868.31 67257.43
Temperature F 100.00 100.00 96.42 100.00 100.00 1929.46
Pressure PSIG 500.00 500.00 500.00 500.00 500.00 300.00
Total Inlet Mass
(LB/DAY)
67257.43
Total Outlet Mass
(LB/DAY)
67257.43
18. 17
Fischer-‐Tropsch Reactor Design and Optimization
A Fischer Tropsch reaction unit (FTR) was designed to meet certain specifications. During this
reaction, a feed stream of CO and H2 react to produce alkanes ranging from C1 to C60 and H2O. Inert
side components were evident in the feed stream as well; N2 and CO2. Using component outputs
generated from ASPEN PLUS simulation software in the previous milestone, a MATLAB code was created
to observe changes in temperature, pressure, and conversion by assessing various constraints such as
the number of reactor tubes, inlet temperature and pressure, tube diameter, and cooling water
temperature.
Equation Development
For the FTR reaction, where syngas is converted to hydrocarbons and water, a variety of
equations were used to implement the parameters set forth by the problem. A detailed outline of the
equations used and their explanations will be presented. The reaction used to declare rate equations
(particularly the stoichiometry) for and was by Equation (1):
(1)
For , the following reaction, Equation (2), was used to declare the same:
(2)
Other components in the system, such as N2 and CO2, are considered inert and therefore do not affect
the system. The overall rate equation for the consumption of is in Langmuir-‐Hinshelwood form is
given by Equation (3) with the following parameters:
(3)
Where the variables involved are defined by Equations (4), (5), and (6):
(4)
19. 18
(5)
(6)
and were determined from the total system pressure and the mole fractions of and as
shown by Equations (7) and (8):
(7)
(8)
The selectivity of the produced alkanes ( for n values of 1 to 60) are as follows, given by
Equation (9). For methane (n=1):
(9)
Where is denoted by Equation (10):
(10)
For C2 to C4 alkanes (n=2, 3, 4) the selectivity is given by Equation (11):
(11)
Where is given by Equation (12):
(12)
To obtain flow rates for C1 through C4 alkanes in the product stream, simple rearrangement of the
selectivity for methane produces , given by Equation (13):
(13)
Thus, for alkanes C2 through C4,Equations (14) through (16) give:
(14)
(15)
(16)
20. 19
For C5+ alkanes (n=5 to 60), the distribution of alkane products is defined by Mn, the relative mole
fraction of Cn , given by Equation (17):
(17)
given by Equation (18):
(18)
To obtain selectivity, Mn must be divided by the sum of all Mn
in question. Then, the mole fraction must be multiplied by n to obtain a per carbon basis. Since the
selectivity is simply (1-‐St), the sum of the selectivities of the first four alkanes, the selectivity for C5+
alkanes is defined by Equation (19):
(19)
Where
(20)
The rate of production of C5+ alkanes is then given by Equation (21):
(21)
For the pressure drop consideration of the FTR, the Ergun equation was used to obtain an expression for
the change in pressure per catalyst weight, given by Equation (22):
(22)
Where is the cross-‐sectional area of the pipes, is the given void fraction, is the density of the
solid catalyst particles, is the fraction of feed pressure over total pressure, is the fraction of total
21. 20
temperature over initial temperature, and is the total flow rate over the initial flow rate into the
reactor. is defined by the following, given by Equation (23)::
(23)
Where is the inlet gas mass velocity, is the density of the gas, is the gravitational constant, is
the diameter of the catalyst particle, and is the viscosity of the gas, which was assumed to be the
average of the viscosities of each component in the feed.
A similar expression was obtained from literature for the change in temperature in the FTR per catalyst
weight. This expression is as follows, given by Equation (24):
(24)
Where is the given heat of reaction, is the inside surface area per volume, is the temperature
of the cooling water, and is the bulk density of the catalyst ( ). Equations for and
can be seen below. refers to the diameter of a single tube, given by Equation (25):
(25)
For methane, H2, O2, H2O, N2, and CO2, average heat capacities have been obtained from ASPEN HYSYS.
22. 21
For C2 though C10 alkanes, heat capacity values were obtained from literature and needed no equation
to compute. However, for C2+ alkanes, a general formula has been obtained to calculate heat capacities,
given by Equation (26):
(26)
For economic considerations, a number of equations were used to determine the cost of the reactor, .
This is obtained by estimating the required weight of stainless steel for the shell and tubes of the
reactor, as seen below, given by Equation (27):
(27)
Where, given by Equation (28):
(28)
Where is the inner diameter of the FTR, L is the reactor length, and is the given density of
stainless steel. , shell thickness, is given by Equation (29):
(29)
Where S is the maximum allowable stress, in psi, and E is the weld efficiency. , the design pressure, is
given by Equation (30):
(30)
where is the nominal pressure in psig, assumed to be the shell side pressure.
FTR Optimization
The optimization process for the Fischer Tropsch Reactor began with a fractional factorial
design. A factorial design is a method that aids in the optimization of systems that involves multiple
23. 22
independent variables. The parameters that were optimized were: the cost of the reactor, the
conversion, and the average carbon chain length of the products. It was desirable to minimize the cost,
maximize the conversion, and achieve an average chain length between 10 and 15 carbons. Minimizing
cost and maximizing the conversion of the reaction will maximize the profit. It is desirable to have an
average carbon chain length between 10 and 15 carbons because hydrocarbon products of chain lengths
between 5 and 21 require the least amount of processing to prepare the product for sale, and therefore
cost the least to process. Hydrocarbons longer than 21 carbons are wax and therefore must be
processed in the hydro-‐isomerization unit to convert them to lighter products. It was desirable to avoid
hydrocarbons of chain length 1 to 4 as these products are the least profitable. Hydrocarbons of a chain
length 5 to 20 maximized the potential profit while minimizing money and time required to process the
material.
The independent variables involved in the optimization were: temperature and pressure of the
inlet feed stream, the number of reactor tubes, the diameter of the reactor tubes, the weight of the
catalyst packing in each reactor tube, and the temperature of the cooling water stream. The fractional
factorial was set up as follows:
Design: =
Levels: 2
Number of Factors investigated, k: 6
Number of Generators, p: 2
Runs: 16
The number of generators indicates the number of factors that will not be considered
independently. This creates a factorial design that is a fraction of the full design. The final set of solution
parameters will not be as optimal as a full design would result in, but the fractional factorial design is
more efficient to analyze and will still give a decent estimate of the optimal parameters.
Minitab was used to create the design matrix. This matrix was transferred to an excel file and
then imported to Matlab and converted to a matrix that corresponds to the 16 test runs. For each
24. 23
independent variable initial high and low values of a testable range were selected. These values were
based on information given by the plant design specifications. Table (5) displays the first set of high and
low values selected for optimization.
Table 5: First set of FTR Optimization Parameters
Parameter High Low
Inlet Temperature (F) 450 390
Catalyst Weight (lb) 30 2
Number of Tubes 1000 800
Diameter of tube (in) 2 1
Pressure (psia) 330 270
Shell Temperature (K) 450 373
The Matlab simulation was run using this initial set of values and the cost, conversion, and
average carbon chain length of products were compared. All of the runs were checked to ensure that
the reactor length did not exceed 60 ft and the diameter did not exceed 20ft. Trends in the data were
analyzed and the range of the high and low values for each variable were adjusted. The first adjustment
that was made was increasing the shell temperature. It was evident from the first round of optimization
that higher shell temperatures corresponded to a significantly higher conversion. After the second
round of optimization, it was shown that the inlet pressure had little effect on any of the dependent
parameters and therefore was not changed again. The range for the catalyst weight, diameter of the
reactor tube, and inlet temperature were narrowed based on trends seen in the second round of
optimization runs. It was shown that lower inlet temperatures produced a higher average carbon chain
length, so the range of high and low values was change to 390 -‐ 400 . It was also shown that higher
catalyst weight packings contributed to a higher conversion and smaller reactor tube diameters lowered
cost. These ranges were adjusted to be 10 15 lbs of catalyst and 1 1.5 in diameter reactor tubes.
Several more rounds of optimization were completed, every time adjusting parameters in the same way.
Finally, a final set of optimal independent parameters was decided upon. At this point the reactor length
was 55.15 ft long. The maximum length of the reactor allowed is 60ft. The weight of the catalyst was
25. 24
increased until the maximum length was acquired. It was checked that this catalyst increase maximized
the conversion with little effect on the cost and no negative effect on the average carbon chain length of
the products. The optimal independent and dependent variables of the FTR are located in Table (6).
Table 6: Table of FTR Optimized Parameters
Variable Independent/Dependent Value Units
Inlet Temperature Independent 400
Catalyst Weight Independent 16.5 lb
Number of Tubes Independent 1000 n/a
Diameter of tube Independent 1 In
Pressure Independent 330 Psia
Shell Temperature Independent 480 K
Cost Dependent 7.4 Million USD
Average Carbon Chain
length of Products
Dependent 8.03 Carbons
Conversion Dependent 0.94 n/a
In order to characterize the optimization process of the FTR a sensitivity analysis was performed
exploring how sensitive the total cost of the reactor, the reaction conversion, and the average carbon
chain length of the products are to changes in the independent variables: temperature and pressure of
the inlet feed stream, the number of reactor tubes, the diameter of the reactor tubes, the weight of the
catalyst packing in each reactor tube, and the temperature of the cooling water stream. Tornado plots
were constructed for each dependent variable to demonstrate these relationships and represented in
Figure (2).
26. 25
Figure 2: Tornado plot for the sensitivity analysis of overall reactor cost
From Figure (2), it can be seen that the temperature of the cooling water had the greatest effect
on the cost of the reactor. Pressure had no impact, and the diameter of the reactor tube, number of
reactor tubes, and catalyst weight all equally effected the overall cost.
Figure 3: Tornado plot for the sensitivity analysis of the conversion of the reaction
0.0E+00 5.0E+06 1.0E+07 1.5E+07
Inlet Temp
Catalyst Weight
Number of tubes
Diameter of tubes
Pressure
Shell Temperature
Cost (million USD)
0.00 0.50 1.00 1.50
Inlet Temp
Catalyst Weight
Number of tubes
Diameter of tubes
Pressure
Shell Temperature
Conversion
27. 26
Figure (3) demonstrates a similar relationship between the independent variables and there
effect on the conversion of the reaction. Once again, pressure of the feed stream had very little effect,
and the temperature of the cooling water stream was the most important factor is maximizing the
conversion.
Figure 4: Tornado plot for the sensitivity analysis of the average carbon chain length of the products
Figure (4) shows that the independent variables that had the greatest effect on the average
carbon chain length of the products are the shell temperature, diameter of the reactor tubes and the
inlet temperature of the feed stream.
The products from the FTR unit require significant separation, a comprehensive step by step
walk through the separation of these products will be discussed in the following section. Figure (5) is the
PFD of the FTR unit and Table (7) is the corresponding stream tables with the unit.
0.0 2.0 4.0 6.0 8.0 10.0
Inlet Temp
Catalyst Weight
Number of tubes
Diameter of tubes
Pressure
Shell Temperature
Average Chain Length (# of carbon atoms)
28. 27
FTR PRODUCTSYNGAS
FISCHER TROPSCH REACTOR
Figure 5: A PFD of the FTR unit
Table 7: Stream Tables for the FTR Unit
Units SYNGAS FEED
Component Mole Flow
CO MSCFD 423.52 42.47
CO2 MSCFD 106.38 104.74
H2 MSCFD 847.04 54.26
WATER MSCFD 394.01 775.78
N2 MSCFD 3.59 5.00
CH4 MSCFD 4.48 13.50
ETHANE MSCFD 0.00 0.38
PROPANE MSCFD 0.00 0.38
N-‐BUT-‐01 MSCFD 0.00 0.38
NAPTHA MSCFD 0.00 7.04
DIESEL MSCFD 0.00 6.89
C21-‐C25 MSCFD 0.00 2.04
C26-‐C29 MSCFD 0.00 1.19
C30-‐C35 MSCFD 0.00 1.28
C36-‐C47 MSCFD 0.00 1.42
C48PLU MSCFD 0.00 0.66
Mole Flow MSCFD 1779.02 1017.42
Mass Flow LB/DAY 67257.43 67314
Temperature F 1929.46 382.21
Pressure PSIG 300.00 278.30
Total Inlet Mass (LB/DAY) 67257
Total Outlet Mass (LB/DAY) 67257
29. 28
Separation Unit Implementation
The products from the Fisher Tropsch Reactor consist of hydro carbons ranging from C1 to C60
along with some left over CO, CO2, H2, H2O, and N2. These products needed to be separated into
specific groups of hydrocarbons in order to be sold to the customer. The different groups of
hydrocarbons consist of Naptha (C5-‐C10), Diesel (C11-‐C20), wax (C21+), and the individual components
of methane, ethane, propane, and butane. The wax components were then sent to a hydroisomerization
unit in order to break those products down to smaller hydrocarbons to be sold. The following summary
is a comprehensive step by step walk through of the separation train to achieve the desired products.
The components in parentheses refer to specific streams in the process and can be seen in the overall
PFD of the process.
The feed from the FTR was sent through an initial flash drum (FLASH1) to separate as much of
the heavier wax components from the desired products. The top stream from this flash drum (DIST1A)
was then cooled by a heat exchanger (HX1) and created a new colder stream (DIST1B) before being put
into another flash drum (FLASH2). The bottom stream from the first flash drum (BOT1) was taken to a
mixer (MIX1) to be mixed with other bottom products. The second flash drum (FLASH2) has three
products and was used to further separate the heavier hydrocarbons. The top stream (DIST2A) was then
cooled by a separate heat exchanger (HX2) to make a colder stream (DIST2B) before being sent to
another flash drum (FLASH3). The non-‐aqueous bottom product from FLASH2 (BOT2) was then sent to
the same mixer (MIX1) as BOT1. The aqueous bottom stream (BOT2AQU) was sent to a different mixer
(MIX3) to be mixed with other aqueous bottom products. DIST2B was sent to FLASH3 to further
separate the heavier components. FLASH3 products include one vapor top stream (DIST3A), one
aqueous bottom stream (BOT3AQU), and one liquid bottom stream (BOT3). DIST3A was cooled by a
separate heat exchanger (HX3) to create another colder stream (DIST3B) and sent to another flash drum
(FLASH4). BOT3 was sent to a separate flash drum (FLASH6) to have its components be separated
30. 29
further. BOT3AQU was sent to the aqueous mixer (MIX3). DIST3B was sent to FLASH4 and produced
three products. The top product (DIST4A) was cooled by a heat exchanger (HX4) and the cooled stream
(DIST4B) and sent to another flash drum (FLASH5). The non-‐aqueous bottom product (BOT4) was sent
back to MIX1 to be mixed with BOT1 and BOT2. These three streams were mixed together to make a
stream consisting of all of their components (S1). S1 was then sent through a valve (VALVE1) to decrease
the pressure of the stream (S2). The aqueous bottom product of FLASH4 (BOT4AQU) was sent to MIX3
with the other previous aqueous bottom products. DIST4B entered FLASH5 and created three more
streams. The top vapor stream (DIST5) was sent to a mixer (MIX5). The non-‐aqueous bottom stream
from FLASH5 (BOT5A) was cooled by a heat exchanger (REFRIDGE) to create a colder stream (BOT5B).
BOT5B was then sent to a decanter (DECANT) to create two different streams (S11) and (S12). S11 was
sent back to MIX3 with the other aqueous streams. The aqueous bottom product (BOT5AQU) was sent
to MIX3 with the other aqueous bottom streams (BOT2AQU, BOT3AQU, BOT4AQU, BOT5AQU,
BOT8AQU) and S11 to be mixed. The product of MIX3 (S16) was used later in the process.
The products of FLASH6 consisted of two streams. The bottom product (BOT6) was sent to a
mixer (MIX2) to be mixed with other streams. The top product (DIST6A) was cooled by a heat exchanger
(HX5) to create a cooler stream (DIST6B) for further separation. DIST6B was sent through a flash drum
(FLASH7). FLASH 7 created two products to be sent elsewhere in the system. The bottom product (BOT7)
was sent to MIX2 to be mixed with S2 and BOT6. The top product from FLASH7 was sent to a mixer
(MIX4) to be mixed with certain products later in the separation. The product of MIX2 (S3) was cooled
by a heat exchanger (HX6) to make a cooler stream (S4). The pressure of S4 was decreased by a valve
(VALVE2) to create the stream S5.
S5 was sent to a distillation column (DIST1) with a total condenser and a partial reboiler to
separate the product into 2 streams. The bottoms product is the WAX stream that was sent to the
hydroisomerization unit. The distillate (S6) was sent through a heat exchanger (HX7) to cool the stream
31. 30
to create a colder stream S7. S7 was put into a distillation column (DIST2) with a total condenser and a
partial reboiler to create two more product streams. The bottoms stream is the DIESEL product stream.
The distillate (S8) was sent to a mega compressor (B3) to compress the stream (S9). MIX4 mixed the
streams DIST7 from FLASH7, S12 from the decanter, and S9 from the mega compressor (B3). The
product from MIX4 (S13) was sent to a heat exchanger (HX10) to cool the stream and create S14. S14
was sent to a flash drum (FLASH8) and created three product streams. The non-‐aqueous bottom stream
was the NAPTHA desired product stream. The aqueous bottom stream was sent to MIX3 to be mixed
with the other aqueous bottom streams and S11. The top product (DIST8) was sent to a compressor
(COMP2) to compress the stream and create a highly pressurized stream (S10).
S10 was sent to a mixer (MIX5) to be mixed with the DIST5 product from FLASH5. The product
from MIX5 (S15) was sent to a valve (VALVE4) to decrease the pressure of the stream and create (S16).
S16 was then cooled by a heat exchanger (HX8) to create the cooler product (S17). The product of from
MIX3 (S18) was sent to a valve (VALVE3) to decrease the pressure in the stream and create S19. S19 was
cooled by a heat exchanger (HX9) to create the cool stream S20. S17 and S20 were sent to a stripping
column (STRIP1) to create 2 streams. The top stream (FLUEGAS1) is a product stream. The bottom
product (BOTSTRIP1) was sent to another stripping column (STRIP2). Another air stream (AIR) necessary
for the stripping column was added and sent into STRIP2. The two products of STRIP2 are FLUEGAS2 and
WASTEH2O.
The product streams from this separation process that contain desired products are NAPTHA,
DIESEL, and WAX streams. The NAPTHA and DIESEL streams are ready for packaging and sale. The WAX
stream requires further separation in order to achieve the desired products. A Hydroisomerization Unit
was employed in order to achieve the desired separation. This will be discussed in detail in the following
section. Figure (6) is the PFD for the separation unit, Table (8) is the stream tables of the inlet and outlet
32. 31
streams for the separation unit, and Table (9) displays the mass balance values for the Separation Unit.
The full stream tables can be seen in Appendix B.
34. 33
Table 8: Inlet and Outlet Stream Tables for the Separation Unit
Units FEED AIR NAPTHA DIESEL WAX WASTE
H2O
FLUE
GAS1
FLUE
GAS2
From FLASH8 DIST2 DIST1 STRIP2 STRIP1 STRIP2
To FLASH1 STRIP2
Substream:
MIXED
Phase: Mixed Vapor Liquid Liquid Liquid Liquid Vapor Vapor
Component
Mole Flow
CO MSCFD 42.47 0.00 0.00 0.00 0.00 0.00 42.44 0.03
CO2 MSCFD 104.74 0.00 0.03 0.00 0.00 0.00 104.32 0.39
H2 MSCFD 54.26 0.00 0.00 0.00 0.00 0.00 54.26 0.00
WATER MSCFD 775.78 0.00 0.01 0.00 0.00 254.76 278.71 242.31
N2 MSCFD 5.00 19.13 0.00 0.00 0.00 0.01 5.00 19.12
CH4 MSCFD 13.50 0.00 0.00 0.00 0.00 0.00 13.48 0.02
ETHANE MSCFD 0.38 0.00 0.00 0.00 0.00 0.00 0.38 0.00
PROPANE MSCFD 0.38 0.00 0.00 0.00 0.00 0.00 0.37 0.01
N-‐BUT-‐01 MSCFD 0.38 0.00 0.01 0.00 0.00 0.00 0.36 0.01
NAPTHA MSCFD 7.04 0.00 5.86 0.00 0.00 0.00 1.18 0.00
DIESEL MSCFD 6.89 0.00 0.26 6.39 0.21 0.00 0.02 0.00
C21-‐C25 MSCFD 2.04 0.00 0.00 0.00 2.04 0.00 0.00 0.00
C26-‐C29 MSCFD 1.19 0.00 0.00 0.00 1.19 0.00 0.00 0.00
C30-‐C35 MSCFD 1.28 0.00 0.00 0.00 1.28 0.00 0.00 0.00
C36-‐C47 MSCFD 1.42 0.00 0.00 0.22 0.95 0.00 0.26 0.00
C48PLU MSCFD 0.66 0.00 0.00 0.00 0.66 0.00 0.00 0.00
Mole Flow MSCFD 1017.42 91.08 6.17 6.61 6.34 254.81 500.79 333.79
Temperature F 382.14 100.00 140.00 345.83 460.33 196.55 183.93 196.31
Pressure PSIG 278.30 0.00 0.00 -‐13.54 -‐13.54 0.00 0.00 0.00
Table 9: Mass Balance for the Separation Unit.
FEED AIR NAPTH
A
DIESEL WAX WASTE
H2O
FLUE
GAS1
FLUE
GAS2
Mass
Flow
LB/DA
Y
67313.9
5
7478.8
5
1748.93 3848.5
8
7532.5
2
12098.1
8
30536.3
7
19028.1
1
Total Inlet
Mass
(LB/DAY)
74793
Total
Outlet
Mass
(LB/DAY)
74793
* The blue boxes are inlet mass and green are the outlet mass
35. 34
Hydro-‐isomerization Unit
The hydroisomerization unit took the wax produced from the separation train and broke it down
into smaller hydrocarbon chains. The amount of each carbon chain produced from the
hydroisomerization unit were assumed by a weight percent of the total weight input into the unit. The
total weight of the WAX stream was 313.85 . Table (10) demonstrates the amount of each
hydrocarbon leaving the hydroisomerization unit based on the weight percent given in the plant
specifications.
Table 10: Mass Flow rates of Hydrocarbons from the Hydroisomerization Unit
Hydrocarbon Mass flow (lb/hr) Wt% from total Wax
C1 3.14 1
C2 1.57 0.5
C3 10.98 3.5
C4 10.98 3.5
Naptha 78.46 25
Diesel 208.71 66.5
These components are mixed and require separation in order to recover the naptha and diesel portions
of the stream (WAX). The first step is to increase the pressure of the mixture from 1.16 psia to 14.7 psia.
This was completed by using a compressor. The unit operation table of the compressor can be seen in
the equipment specifications section. After the compressor the mixture was sent to a distillation column
(HIDIST1) to separate the C1 C4 components from the naptha and diesel components of the stream.
The first distillation column was operated at atmospheric pressure with a partial condenser and a partial
reboiler. The partial condenser was used because the components being produced in the distillate (C1-‐
C4) are all vapor and do not require any more separation. The unit operations table for the first
distillation column can be seen in equipment specifications. The summary of distillation column 1 can be
seen in Table (11).
36. 35
Table 11: A Summary of the Separation in Distillation Column 1
Mole Fraction in
distillate
Mole Fraction in
bottoms
Temperature
(F)
Pressure
(psia)
C1-‐C4 0.9919 0.0062 31.3 14.7
Naptha and
Diesel
0.0081 0.9938 247.9 14.7
As is demonstrated in Table (11), a very good separation is achieved from the first distillation
column unit. The C1-‐C4 stream is a product and no longer required any more separation. The bottoms
product (NAPDIE) was sent to another distillation column because the mixture needed further
separation to create more product to be sold. Another distillation column (HIDIST2) was used to
separate the naptha and diesel into each pure component. A total condenser and a partial reboiler were
implemented for this column because both the distillate and the bottoms need to be liquid for transport
and sale. The unit operations for the second distillation column can be seen in equipment specifications.
The summary of distillation column 2 can be seen in Table (12).
Table 12: A Summary of the Separation in Distillation Column 2.
Mole Fraction
in distillate
Mole Fraction
in bottoms
Temperature (F) Pressure (psia)
Naptha 0.9856 0.0001 182.6 14.7
Diesel 0.0001 0.9999 506.1 14.7
Figure (7) demonstrates the separation of the naptha and diesel components and is the PFD of the
Hydroisomerization Unit. Table (13) are the stream tables for the HI unit.
37. 36
Hydroisomerization Unit
HI PROD1
HI COMPRESSOR
HIDIST1
HIDIST2
HI PROD2
C1-‐C4
NAPDIE
NAPTHA2
DIESEL2
WAX
Figure 7: Hydroisomerization unit PFD
Table 13: Stream tables and Mass Balance for the Hydroisomerization Unit
Units HI PROD 1 HI PROD 2 C1-‐C4 NAPDIE NAPTHA DIESEL
From HI
CIOMPRESSOR
HIDIST 1 HIDIST 1 HIDIST 2 HIDIST 2
To HI
CIOMPRESSOR
HIDIST 1 HIDIST 2
Substream:
MIXED
Phase: Vapor Vapor Vapor Liquid Liquid Liquid
Component
Mole Flow
CH4 MSCFD 1.78 1.78 1.78 0.00 0.00 0.00
ETHANE MSCFD 0.48 0.48 0.47 0.00 0.00 0.00
PROPANE MSCFD 2.27 2.27 2.24 0.03 0.03 0.00
N-‐BUT-‐01 MSCFD 1.72 1.72 1.66 0.07 0.07 0.00
NAPTHA MSCFD 6.90 6.90 0.05 6.85 6.85 0.00
DIESEL MSCFD 9.09 9.09 0.00 9.09 0.00 9.09
Mole Flow MSCFD 22.24 22.24 6.20 16.04 6.95 9.09
Mass Flow LB/DAY 7532.50 7532.50 640.26 6892.23 1883.04 5009.19
Temperature F 460.33 543.10 31.28 247.89 182.57 506.10
Pressure PSIG -‐13.54 0.00 0.00 0.00 0.00 0.00
Total Inlet Mass
(LB/DAY)
7532.50
Total Outlet
Mass (LB/DAY)
7532.50
38. 37
Heat Integration
In industrial processes there often exists the opportunity to integrate the streams involved using
energy released from some streams to heat other streams. This heat integration can decrease utility
costs. For this system streams from the syngas unit, FTR unit and the separation unit were considered
for heat integration. A compilation of the selected streams and the associated parameters are located in
Table (14). Each stream was determined to be a hot stream, which gives off heat, or a cold stream,
which requires heat.
39. 38
Table 14: Streams Considered for Heat Integration
Stream Label Condition Source
Temperature
( )
Target
Temperature ( )
Cp (Btu/ ) Duty, Q (Btu)
Heat
Exchanger
1
H1 Hot 409 365 17116.34 -‐753,119
Heat
Exchanger
2
H2 Hot 365 365 n/a -‐515,546
Heat
Exchanger
9
H3 Hot 323 90 1706.44 -‐397,601
Flash
Column 2
H4 Hot 365 365 n/a -‐140,997
Heat
Exchanger
4
H5 Hot 250 95 610.95 -‐94,697
HI Unit
Distillation
Column 2
reboiler
H6 Hot 543 247 260.79 -‐77,194
Heat
Exchanger
6
H7 Hot 390 200 312.14 -‐59,307
Syngas
Heat
Exchanger
H8 Hot 1938 400 1790 -‐2,753,020
Flash
Column 1
C1 Cold 382 409 1736.63 46,889
HI Unit
distillation
column 2
reboiler
C2 Cold 247 506 221.86 57,240
Distillation
Column 1
reboiler
C3 Cold 199 460 498.07 127,646
Stripper 2
(top stage)
C4 Cold 150 196 10517.61 483,810
Stripper 1
(bottom
stage)
C5 cold 90 190 7178.58 717,858
40. 39
In order to determine which streams to integrate, the absolute value of the energy duties for all
streams were compared. The streams located in Table (14) were selected because the associated heat
duties were significantly larger than the duties of other streams.
The temperature interval method(13)
was used in the preliminary design of the heat exchanger
network. The minimum approach temperature was selected to be 10 . Then, the temperature for the
hot streams were adjusted down 10 accordingly. All of the source and target temperatures were
sorted from largest to smallest and Figure (8) was constructed to show the temperature intervals that
exist in the process.
Figure 8: Diagram of the Temperature Interval Method
41. 40
Using the heat capacities of each stream, the change in enthalpy for each interval was calculated
using the Equation (31).
) (31)
Then, the residual enthalpy of each interval was calculated by adding to each enthalpy the value before
it. This gives an indication of where the pinch point in the process exists. Table 15 lists the change in
enthalpy and residual enthalpy for each interval.
Table 15: Calculated Enthalpies for the Temperature Interval Method
Interval
1 1395 2497050 2497050
2 27 55371 2552421
3 46 84131 2636552
4 51 68333 2704885
5 10 -‐3968 2700918
6 9 -‐3571 2697347
7 8 119437 2816784
8 2 33332 2850116
9 25 1081002 3931118
10 42 -‐5796 3925322
11 65 101949 4027272
12 8 14322 4041594
13 3 7204 4048798
14 38 81338 4130136
15 3 7889 4138024
16 6 -‐47328 4090696
17 40 -‐615152 3475544
18 60 -‐291671 3183872
19 5 11587 3195459
20 5 8532 3203992
In this case, the pinch point exists before interval 1, or at 1938 . This means that there is no
hot utility for the process. The cold utility for the process is 3,203,992 btu/hr, the residual enthalpy
42. 41
associated with interval 20, the last temperature interval being evaluated. No external heating sources
will be necessary for the integration of these streams, because the hot utility equals zero.
The next step was to match streams based on temperatures of the streams and energy duties available.
In order for a hot stream to be able to heat a cold stream, the temperatures of the hot stream must be
greater than the temperatures of the cold streams. Only three of the hot streams were necessary to
supply enough energy to heat all of the cold streams in question. Figure 9 shows the final heat
integration of the process; all temperatures listed below are in units of
Figure 9: Final Heat integration Diagram
Direction of heat transfer is from the hot (red) streams to the cold (blue) streams and are
depicted by the green circles. Stream H8 is the feed to the FTR unit from the Syngas unit. The stream
exits the Syngas Unit at 1938 and needs to enter the FTR Unit at 400 . The amount of energy
associated with this stream is more than sufficient to heat cold streams C1, C2, and C3. The
temperatures labeled above the heat transfer streams represent is the end of the temperature interval
43. 42
that that hot stream needs to decrease to supply enough energy for that cold stream. For example, C3
requires 127,646 btu/hr to continuously heat it from 199 to 460 . 127,646 btu/hr is available from
stream H8 by cooling it from 1938 to 1866.69 . Sample calculations are shown below.
(2)
Energy required:
is the intermediate temperature that stream H8 will decrease to after heating stream C3. This value
must be great that the target temperature for H8, which in this case is 400 H8 has more than enough
energy to heat C3, as (1866.69 ) is much greater than (400 This analysis was performed for
the other four cold streams to produce the final heat integration design. A PFD of the process is shown
in Figure 10.
44. 43
Figure 10: Process flow diagram of the integrated heat streams
Equipment Specifications
This section (Tables (16)-‐(29)) lists all of the equipment specifications and operating conditions
for the equipment in the entire process. The material used is carbon steel. Carbon steel is a cheap
material and although we have some corrosive materials in the process, the amount is negligible and
will not affect the equipment.
47. 46
Syngas Unit
Table 24: Syngas Reactor Specifications
Operation Value
Liquid volume (gal) 35.25
Vessel diameter (ft) 1
Vessel tangent to tangent height (ft) 6
Design gauge pressure (psig) 550
Design temperature 1979.46
HI Unit
Table 25: HI Distillation Column: Condenser/Top Stage Performance
Column # Temperature Heat Duty
(Btu/hr)
Distillate
Rate
(lbmol/hr)
Reflux Rate
(lbmol/hr)
Reflux
Ratio
HI DIST 1 31.27 -‐27198.90 0.68 1.36 2
HI DIST 2 182.59 -‐21840.30 0.76 0.61 0.8
Table 26: HI Distillation Column: Reboiler/ Bottom Stage Performance
Column # Temperature ( Heat Duty
(Btu/hr)
Bottoms Rate
(lbmol/hr)
Boilup Rate
(lbmol/hr)
Boilup
Ratio
HI DIST 1 247.88 -‐64574.00 1.76 2.16 1.22
HI DIST 2 506.10 54409.38 1.00 2.60 2.61
Table 27: HI Distillation Column Specifications
Column # Tray type Vessel
diameter
(ft)
Vessel
tangent to
tangent
height (ft)
Design gauge
pressure (psig)
Operating
temperature
Number
of trays
Tray
spacing
(in)
HI DIST 1 SIEVE 1.5 30 15 247.91 9 24
HI DIST 2 SIEVE 1.5 30 15 506.13 9 24
Table 28: HI Compressor Equipment Specifications
Operation Value
Design gauge pressure Inlet (psig) 0.30
Design temperature Inlet 460.33
Design gauge pressure Outlet (psig) 0.30
Driver power (hp) 6.38
48. 47
FTR Unit
Table 29: FTR Equipment Specifications
Operation Value
Inlet Temperature 400
Catalyst Weight (lb) 16.5
Number of Tubes 1000
Diameter of tube (in) 1
Pressure (psia) 330
Shell Temperature (K) 480
Utility Summary
A summary of the utility requirements for each unit will now be presented. Values for the
consumption of each utility were read from ASPEN. The heat duties required for the vertical pressure
vessels were obtained by taking excess energy from hot streams from various units in the process. Table
(30) shows heat available and required from the hot streams and cold streams, respectively, as well as a
credit heat utility for left-‐over heat.
Table 30: Utility Summary: Heat integration
Hot Streams
Heat Available
(BTU/hr) Cold Streams
Heat Required
(BTU/hr)
Credit
(BTU/hr)
HX1
753,119 FLASH1
46,889
HX2
515,546
HIdist2(reboil)
57,240
HX9
397,601 DIST1(reboil)
127,646
FLASH2
140,997
STRIP2(reboil)
483,810
HX4
94,697 STRIP1(cond)
717,858
HIdist1(reboil)
77,194
49. 48
HX6
59,307
Syngas HX
2,753,020
Total
4,791,481 Total
1,433,443
3,358,038
Heat from heat integration was used to calculate the flow rate of the cooling water to steam,
which was used as a credit and thus as ensuing profit. This conversion of heat power to steam mass was
used by the following, equation:
(32)
Where is the left-‐over heat, is the heat capacity of water, and are the initial and final
temperatures respectively, chosen as room temperature (77 degrees Fahrenheit) and 353 degrees
Fahrenheit (which represents mid-‐pressure steam). represents the flow rate of steam, which was
used to calculate both the utility costs and credits for different units of the reaction process. Table (31)
shows the utility summary for the Syngas Unit. All utility tables presented exist on a per-‐year basis, more
specifically, for the 300 days the plant is operating.
Table 31: Utility Summary for Syngas Unit
Utility Consumed Units Price Cost
Total Cost per
Year
Methane Feed 150,000 MSCF $ 2.00 $ 300,000 $ 322,000
600# 490 F HP
Steam 3,560 klb $ 5.00 $ 18,000
Carbon Dioxide
500psig & 100F 10,314 MSCF $ 0.40 $ 4,000
