Senior Design Project - Gas-to-Liquids Plant Design

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Design of a Gas-to-Liquids Plant

Soung Hun (Eric) Kim;Ryan Poling-Skutvik;Dhongik YoonThe Cooper UnionChE 161.2 – Process Evaluation and Chemical Systems Design IIProfessor Davis & Professor Okorafor

Table of ContentsList of Tables....................................................................................................................3

List of Figures..................................................................................................................5

Summary..........................................................................................................................6

Introduction.....................................................................................................................7

Conclusions....................................................................................................................10

Recommendations..........................................................................................................11

Project Premises............................................................................................................12

Process Flow Diagram(s)...............................................................................................14

Stream Attributes Table.................................................................................................18

Process Description.......................................................................................................36

Syngas Reactor...........................................................................................................36

Fischer-Tropsch Reactor.............................................................................................37

Hydroisomerization Unit.............................................................................................39

Pumps.........................................................................................................................40

Compressors...............................................................................................................41

Heat Exchangers.........................................................................................................42

Flash Tanks and Three-Phase Separators...................................................................45

Fired Heaters..............................................................................................................47

Distillation Columns....................................................................................................48

Heat Integration............................................................................................................49

Safety.............................................................................................................................52

Environmental................................................................................................................54

Utility Summary.............................................................................................................55

Operating Cost Summary...............................................................................................57

Fixed Costs.................................................................................................................57

Variable Costs.............................................................................................................57

Equipment Information Summary..................................................................................60

Compressors...............................................................................................................60

Drives..........................................................................................................................60

Furnaces.....................................................................................................................60

Process Vessels...........................................................................................................60

Heat Exchangers.........................................................................................................60

Reactors......................................................................................................................60

Towers & Trays...........................................................................................................61

Pumps.........................................................................................................................61

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Fixed Roof Tanks.........................................................................................................61

Source of Costs...........................................................................................................67

Capital Estimate.............................................................................................................72

Economic Analysis.........................................................................................................74

Innovation & Optimization.............................................................................................77

References.....................................................................................................................78

Engineering Calculations...............................................................................................80

Equipment Sizing........................................................................................................80

Computer Programs.......................................................................................................86

MatLab Code to Model FTR........................................................................................86

Computer Process Simulations......................................................................................89

PRO/II Keyword Input File..........................................................................................89

List of TablesTABLE 1 - SPECIFICATIONS FOR RAW MATERIALS........................................................................12

TABLE 2 - SPECIFICATIONS FOR PRODUCTS................................................................................12

TABLE 3 - QUOTED AND MARKET PRICES FOR RAW MATERIALS....................................................12

TABLE 4- UTILITY COSTS..........................................................................................................13

TABLE 5 - REACTIONS OCCURRING IN THE SYNGAS REACTOR.......................................................36

TABLE 6 - SUMMARY OF HAZARDOUS CHEMICALS [13]...............................................................53

TABLE 7- ANNUAL UTILITY USAGE............................................................................................56

TABLE 8- FIXED COSTS............................................................................................................57

TABLE 9- RAW MATERIALS COST...............................................................................................58

TABLE 10- TOTAL LABOR COSTS...............................................................................................58

TABLE 11- UTILITY COSTS........................................................................................................58

TABLE 12- VARIABLE COSTS.....................................................................................................59

TABLE 13- TOTAL EQUIPMENT COST.........................................................................................61

TABLE 14- COMPRESSOR SUMMARY..........................................................................................62

TABLE 15- DRIVE SUMMARY.....................................................................................................62

TABLE 16- FIRE HEATER SUMMARY..........................................................................................63

TABLE 17- PROCESS VESSELS SUMMARY...................................................................................63

TABLE 18- HEAT EXCHANGER SUMMARY...................................................................................64

TABLE 19 - REACTOR SUMMARY................................................................................................65

TABLE 20- TOWER SUMMARY...................................................................................................65

TABLE 21- TRAYS SUMMARY.....................................................................................................65

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TABLE 22- PUMP SUMMARY.....................................................................................................66

TABLE 23- FIXED ROOF TANK SUMMARY...................................................................................66

TABLE 24- BARE MODULE COST EQUATIONS.............................................................................69

TABLE 25- EQUIPMENT COST DATA TO BE USED WITH EQUATION 1.............................................70

TABLE 26- PRESSURE FACTORS FOR PROCESS EQUIPMENT.........................................................70

TABLE 27- MATERIAL FACTORS FOR HEAT EXCHANGERS, PROCESS VESSELS, AND PUMPS............71

TABLE 28- BARE MODULE FACTORS FOR EQUIPMENT.................................................................71

TABLE 29- WORKING CAPITAL..................................................................................................72

TABLE 30- DIRECT COSTS ESTIMATE.........................................................................................72

TABLE 31- EQUIPMENT INSTALLATION ESTIMATE.......................................................................73

TABLE 32- INDIRECT COSTS ESTIMATE......................................................................................73

TABLE 33- TOTAL CAPITAL INVESTMENT ESTIMATE....................................................................73

TABLE 34- SALES REVENUE.....................................................................................................74

TABLE 35- CUMULATIVE CASH FLOW........................................................................................94

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List of FiguresFIGURE 1- NATURAL GAS RESERVES...........................................................................................7

FIGURE 2- MARCELLUS SHALE FORMATION.................................................................................8

FIGURE 3- EXPECTED PRICE GAP BETWEEN NATURAL GAS AND OIL..............................................9

FIGURE 4 - PFD OF SYNGAS SECTION........................................................................................15

FIGURE 5 - PFD OF FTR AND HI SECTIONS...............................................................................15

FIGURE 6 - PFD OF TAIL GAS/LPG SEPARATION SYSTEM..............................................................16

FIGURE 7 - PFD OF NAPHTHA/DIESEL SEPARATION SYSTEM.........................................................17

FIGURE 8 - TEMPERATURE PROFILE OF THE FISCHER-TROPSCH REACTORS....................................38

FIGURE 9- HEAT INTEGRATION CURVE BEFORE ADJUSTMENT......................................................49

FIGURE 10- HEAT INTEGRATION CURVE AFTER ADJUSTMENT......................................................50

FIGURE 11- CUMULATIVE CASH FLOW......................................................................................75

FIGURE 12- ENTRAINMENT FLOODING CAPACITY IN A TRAYED TOWER.........................................84

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SummaryTo take advantage of the current rising price of oil and decreasing price of natural gas, a gas-to-liquids plant was designed. The plant will be located in Deposit, NY and will take 3 years to construct and will operate for 15 years. This location is ideal for the proposed plant, since the plant will be in close vicinity to the Marcellus Shale formation, a major source of natural gas.

The purpose of the plant is to convert the methane from natural gas into liquid fuels, while producing a significant amount of steam, which will also be sold. The process can be broken down into three parts: Syngas, Fischer-Tropsch, and Hydroisomerization. In the Syngas phase, the methane will be heated and reacted in an autothermal reformer to produce syngas, a composition of carbon monoxide and hydrogen. This syngas stream will then be introduced to the FTR unit, which with the presence of an ultra-stable Cobalt-based FTR catalyst, will be converted into hydrocarbons of various lengths. Lights and other useful products like LPG will be separated from this stream and the rest will be introduced to the hydroisomerization unit. In the HI unit, the paraffins are isomerized and the wax is converted to lighter products. A significant portion of the products will be generated from the HI unit and these oil products (naphtha and diesel) will be separated and stored for sale.

Since the design process produces a significant amount of carbon monoxide and other very flammable products at high pressures and high temperatures, many precautions will be taken. One system that will be employed is the gas vent system, which will vent hazardous gases like hydrogen and carbon monoxide, in times of leakage or failure in the system.

Apart from the safety of the plant and the workers, the environment must also be protected from the operation of the plant. Because the plant design produces significant amounts of carbon dioxide, carbon monoxide, and liquid fuels, it is important that these substances are not introduced into the atmosphere/environment. Thus, to minimize environmental impact, the carbon dioxide will be completely recycled within the plant and wastewater treatment systems will be installed to purify the water that may be reintroduced to the environment.

Following the first year of operation (50% capacity), the plant will operate at 100% capacity, producing approximately 55,000 barrels of products per day. This will generate annual revenue of approximately $1.64 billion on operating costs of approximately $1.37 billion. The total capital investment however, is significant: approximately $3.76 billion. With the expected revenue stream, the plant will breakeven at Year 12, with a discounted cash flow return on investment of 7.81%. The profitability of the process however, is heavily dependent on the market conditions. For example, when current prices specified in

are used to determine the profitability of the plant, the internal rate of return jumps to 17.97%. Thus, the plant design can be a profitable investment. At the given prices of raw materials and products however, this plant is not recommended.

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IntroductionThe Fischer-Tropsch process, which was developed in the 20th century, is a method to convert a mixture of carbon monoxide and hydrogen into liquid fuels. Although it has been decades since the introduction of this process, developments have been limited; only recently, has there been major advancement in the technology. This is mainly due to the fact that the Fischer-Tropsch process requires significant capital investment. Furthermore, cheap prices of oil made it economically unfeasible to invest a tremendous amount of money to develop a gas-to-liquids plant.

In recent years however, the price of oil has increased dramatically and thanks to new technologies like horizontal drilling and hydraulic fracturing, access to natural gas has become much easier and cheaper. Provided below is the Energy Information Administration’s (EIA) depiction of proven natural gas reserves in the United States.

FIGURE 1- NATURAL GAS RESERVES

Although only approximately 325 trillion cubic feet of natural gas has been proven as of 2010, the EIA estimates that there is 2,203 trillion cubic feet of natural gas that is technically recoverable in the United States. The EIA states that “at the rate of U.S. natural gas consumption in 2011 of about 24 Tcf per year, 2,203 Tcf of natural gas is enough to last about 92 years [1].” This abundance however, is not located in obscure parts of the United States. Rather, there are various locations with major shale formation, one of which is the Marcellus shale, which is located in the northeast, only a few miles from Deposit, NY. In this formation alone, there is expected to be 141 trillion cubic feet of natural gas [2]. Provided on the following page is the map of the United States with the various formations. Noted in blue is the Marcellus shale formation.

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FIGURE 2- MARCELLUS SHALE FORMATION

Although the sheer volume of natural gas available is important, it is also important to analyze the price at which natural gas is being sold. To take the analysis one step further, it is more important to understand the price of natural gas in comparison to the price of oil. As mentioned before, this is due to the fact that the price gap between natural gas and oil in previous years was not large enough for gas-to-liquids plants to profit. This gap between the prices however, has increased and is projected to hold for decades to come. Provided on the following page is a visual depiction of the expected gap between the two prices.

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FIGURE 3- EXPECTED PRICE GAP BETWEEN NATURAL GAS AND OIL

According to the EIA, it is estimated that the ratio between the price of oil and the price of natural gas will remain around 3.2 for the next 20 years. In previous years, this ratio was as low as 1.0, which provided no financial incentive for a gas-to-liquids plant.

Because the price gap between the two products is large and expected to hold for years to come, a gas-to-liquids plant may be economical and financially beneficial. Thus, the purpose of this project is to design a possible gas-to-liquids process that may be economically feasible.

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ConclusionsThis project would be technically feasible but does not produce enough revenue to justify the large capital investment. Additionally, it is only about 63% energy efficient, meaning that only 63% of the energy contained within the feed methane is transferred into the liquid fuels. This causes a significant release of carbon dioxide per unit energy in the liquid fuels.

The feed of methane will produce 55,000 barrels per day of liquid fuel – 1,900 of LPG, 19,100 of naphtha, and 34,000 of diesel. These products, along with steam and steam distillate, will generate an annual revenue of $1.64 billion with a net cash-flow of approximately $550 million. The project will have a DCFROR 7.81% on a total capital investment of $3.76 billion.

The project will be able to achieve the production capacity as specified, but further design will be necessary on large process vessels such as the distillation columns and the reactors. Additionally, environmental concerns would have to be addressed more rigorously with contingency plans prepared and processing facilities designed, especially since the plant operates with flammable and toxic components that cannot be released into the environment.

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RecommendationsIt is recommended that this project does not proceed into construction due to a low discounted rate of return and due to negative environmental effects. Overall, the process has a very high initial capital investment of $3.76 billion. With a construction time of three years, the project will have a payback period of approximately 12 years. This long payback period causes the discounted rate of return to be low, 7.81%. This is significantly below the rule-of-thumb return for the stock market of 10%, meaning that the project does not produce a large enough cash flow to be worthwhile. If current market prices, shown in

, for the products are used, the discounted rate of return jumps to 17.97%. If the current price for methane is also used, the discounted rate of return becomes 10.69%. Under these choices, the project should only be pursued if the products can be sold at current market prices while the raw materials are purchased at quoted costs. But at the current quoted prices, the project should not be pursued.

Some research and improvements could be conducted in order to attempt to make the project more economically feasible. These should focus on finding an improved catalyst to increase the rate of reaction for the FTR and the syngas unit, since these reactor systems drastically increase the capital cost of the project. Additionally, the FTR may be converted to a slurry bed or fluidized bed reactor in order to achieve better reaction rates. Furthermore, the process only currently achieves a total conversion of carbon monoxide into hydrocarbons of about 95% so a significant amount of carbon monoxide remains unconverted. If this could be recycled and returned to the feed of the FTR, then the yields would increase.

A very high portion of the variable costs of this plant is the purchase of oxygen from a third-party supplier. The plant is using a significant amount of oxygen on a daily basis, and the total variable cost for this raw material exceeds that of the methane. In order to improve the economics of this project, an air separation plant could be designed in-house rather than being contracted to a third party supplier. This may initially increase the total capital investment of the project, but the increase in operating profit should discount this and help to improve the rate of return.

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Project PremisesThis project will process 14.2 million standard cubic meters of methane to produce 55,000 barrels of liquid fuel. The construction phase of the project is expected to take three years, before a one-year start-up phase in which the plant will run at half capacity. Then the plant will operate for 14 years at 100% capacity before a year of shut-down. The raw materials for the process are methane and oxygen at the specifications found in Table 1. The products of the process are liquid petroleum gas (LPG), naphtha, and diesel. The specifications for these are found in Table 2.

TABLE 1 - SPECIFICATIONS FOR RAW MATERIALS

Temperature (°C)

Pressure (atm)

Purity (%)

Flow Rate (MSCMD)

Cost ($/MSCMD)

Methane

38 35 100 14.2 0.118

Oxygen

24 35 99 8.1 0.314

TABLE 2 - SPECIFICATIONS FOR PRODUCTS

Temperature (°C)

Pressure (atm)

Hydrocarbon Purity (%)

Flow Rate (MCMD)

LPG 43.3 12 98.7 281Napht

ha32 1 99.8 2,976

Diesel 43.3 2 100 5,305

The separation of the products is not as high as desired, especially with the LPG. In the table above, the purities listed are for the percent of hydrocarbons that are present. Any dissolved carbon dioxide, methane, or ethane account for the impurities. However, the amount of naphtha in LPG is quite high. The selling prices of all products are listed in Table 2.

TABLE 3 - QUOTED AND MARKET PRICES FOR RAW MATERIALS

Unit Quoted Price per Unit ($)

Market Price per Unit ($)

LPG Cubic Meter 436.1 490.60 [3]Naphtha Cubic Meter 629.00 754.80 [4]Diesel Cubic meter 754.80 1,015.60 [5]

Methane from Natural Gas

Std. Cubic Meter 0.071 0.118 [6]

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The costs for the utilities are listed in Table 4.

TABLE 4- UTILITY COSTS

Utilities Price Consumption (per

day)

Cost ($/day) Operating Time

(days/yr)

Cost ($/year)

Electricity $0.04/KW-h 162,050 KW-h

$6,482 330 $2,139,000

Fuel Gas $2.84/GJ 3,283 GJ $9,336 330 $3,081,000

Refrigeration $3.30/kg 124 kg $412 330 $136,000Process Water

$0.13/kL 26,339 kL $3,479 330 $1,148,000

Hydrogen $0.13/kg 33,830 kg $4,475 330 $1,477,000

LP Steam $7/ton 194 tons $1,360 330 $449,000Total $8,430,0

00

The economic parameters for the project are listed below

1. Operational period of 15 years

2. 15 year straight-line depreciation

3. 33% tax rate

4. Projected yearly inflation of 3%

The plant is not designed to release any environmental hazards because the cooling water that comes in contact with hydrocarbons is re-circulated and the only designed air emissions are carbon dioxide. However, because the furnaces will be burning hydrocarbons in the presence of nitrogen, nitrogen oxides will be formed. The limits on these emissions are 0.053 parts per million [7], so the plant should monitor emissions to be within this limit. Additionally, wastewater facilities should be on-site in case of any hydrocarbon contamination of the waste water. The cost of processing this contaminated water would be $1.59 per m3 of waste water. Furthermore, the plant produces a significant amount of carbon dioxide, and if a carbon tax were to be imposed, it may be beneficial to install carbon capture technology to reduce the emissions. Currently, however, this does not seem to be likely, so the carbon dioxide is release into the environment.

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Process Flow Diagram(s)Since the PFD for the overall process is too large to be displayed in one image, it will be split into major sections. The syngas section shown in Figure 4 follows the process from the feedstocks through the production of syngas. Methane, carbon dioxide, and steam are mixed together and enter the syngas preheater. In order to prevent coking in the preheater, the steam to methane molar ratio is set at 0.5. Although modeled as two separate reactors, the oxidation and syngas reactors are physically one module in which both the partial oxidation and steam reforming reactions occur simultaneously. The syngas product leaves the equilibrium reactor and passes through two heat exchangers that act as the reboilers for the distillation columns later in the process. The syngas is also used to produce high-pressure steam as a method of heat integration. After the heat recovery steps, the syngas passes through a cooling water heat exchanger and a refrigerated heat exchanger, with R134A refrigerant as the cooling fluid. The water present in the syngas has now condensed and is separated in the following flash tank. The gas is then preheated before it enters the Fischer-Tropsch reactor (FTR).

Figure 5 displays the process from the FTR through the hydro-isomerization unit (HI). The processed syngas enters the FTR in which it undergoes a reaction converting the carbon monoxide into hydrocarbon chains. These are then cooled in E2 and sent through a three-phase separator in which the light components exit the top as a gas, the heavy oils exit the bottom as a liquid, and the water separates from the heavy oils due to their immiscibility. The heavy oil stream passes through another flash tank to separate out most of the naphtha-length hydrocarbons, which by-pass the HI. The heaviest components are sent through the HI to crack them into smaller hydrocarbon chains. The light components leave the first flash tank and flow through a heat exchanger to make low-pressure steam to recover heat. The gas then passes through a refrigeration loop before the rest of the water is separated from another flash tank. The water from the first flash tank mixes with the water separated after the syngas unit and is used to provide the cooling duty on the FTR. This cooling jacket is represented by the heat exchanger labeled FTR_COOLING. Streams Lights_1 and 26 are sent to further separation in the light distillation column and heavy distillation column respectively.

The light ends from Lights_1 and from stream 35 are mixed together and sent through a heat exchanger that used propane as a refrigerant. This liquefies a significant percentage of the components that will end up in the LPG, namely propane and butane. The liquid products are sent to the light distillation column, while the gaseous components are warmed up using cooling water to be split into fuel gas for the syngas preheat furnace and for the steam boiler. The distillation column separates the LPG from additional tail gas. The LPG exits as a liquid from the distillation column but has a significant amount of naphtha, so it is separated using a flash tank in which the propane and butane are vaporized and then re-pressurized and condensed to give the final LPG product. The naphtha from the flash tank is sent to the light ends of the heavy distillation column.

The heavy distillation column and other equipment are displayed in Figure 7. The feeds from the HI unit are sent to this column to separate naphtha from the heavier diesel. The diesel coming out the bottom of the column is pure enough to send straight to storage. The naphtha, however, contains a decent amount of lighter components, so it is sent through a flash tank, with the gas coming off the flash tank sent through a series of compressors and flash tanks and then redirected back to the light distillation

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column system. The liquid naphtha is sent to storage after passing through this flash tank.

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FIGURE 4 - PFD OF SYNGAS SECTION

FIGURE 5 - PFD OF FTR AND HI SECTIONS

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FIGURE 6 - PFD OF TAIL GAS/LPG SEPARATION SYSTEM

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FIGURE 7 - PFD OF NAPHTHA/DIESEL SEPARATION SYSTEM

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Stream Attributes TableStream Name CO2 CW_1 CW_1_HP CW_2 CW_3 CW_3_LPTemperature C 37.78 15.00 15.34 15.56 15.56 15.57Pressure ATM 28.22 1.00 41.83 1.00 1.00 2.36Vapor Act. Vol. Rate M3/HR 1991.67 n/a n/a n/a n/a n/aLiquid Act. Rate (vol)

M3/HR n/a 863.99 862.38 3012.34 202.07 202.06

Total Mass Rate KG/HR 111898.44 863208.34 863208.34 3009340.96 201872.40 201872.40Total Weight Comp. Rates

KG/HR

HYDROGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 H2O KG/HR 0.00 863208.34 863208.34 3009340.96 201872.40 201872.40 OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 NITROGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 CO2 KG/HR 111898.44 0.00 0.00 0.00 0.00 0.00 CO KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 METHANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 ETHANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 PROPANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 BUTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 PENTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 HEXANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 HEPTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 OCTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 NONANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 DECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 UNDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 DODECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 TRIDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 TETDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 PENDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 HXDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 HDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 OCTDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 NONDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 EICOSANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 c2125 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 c2629 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 c3035 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 c3647 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 c48+ KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

Stream Name CW_HP DIESEL DIESEL_PROD

FTR_FEED

FTR_PROD

FTR_WATER

Temperature C 15.90 276.66 43.33 225.85 225.85 208.91Pressure ATM 41.83 2.00 2.00 28.15 28.10 20.34Vapor Act. Vol. Rate M3/

HRn/a n/a n/a 107789.8

744778.48 n/a

Liquid Act. Rate (vol) M3/HR

745.89 292.93 225.90 n/a 244.55 2224.28

Total Mass Rate KG/HR 746540.51

170202.94

170202.94 929724.23

930288.35 1900000.00

Total Weight Comp. Rates

KG/HR

HYDROGEN KG/HR 0.00 0.00 0.00 91873.46 4699.80 0.00 H2O KG/HR 746540.

510.00 0.00 303.88 389796.33 1900000.00

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

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NITROGEN KG/HR 0.00 0.00 0.00 4195.47 4195.47 0.00 CO2 KG/HR 0.00 0.00 0.00 188409.8

0188409.80 0.00

CO KG/HR 0.00 0.00 0.00 637459.08

31848.59 0.00

METHANE KG/HR 0.00 0.00 0.00 7482.52 36366.34 0.00 ETHANE KG/HR 0.00 0.00 0.00 0.00 2165.57 0.00 PROPANE KG/HR 0.00 0.00 0.00 0.00 3175.77 0.00 BUTANE KG/HR 0.00 0.00 0.00 0.00 4185.96 0.00 PENTANE KG/HR 0.00 0.00 0.00 0.00 8434.58 0.00 HEXANE KG/HR 0.00 0.00 0.00 0.00 9176.30 0.00 HEPTANE KG/HR 0.00 0.00 0.00 0.00 9705.92 0.00 OCTANE KG/HR 0.00 0.01 0.01 0.00 10056.90 0.00 NONANE KG/HR 0.00 8.60 8.60 0.00 10254.70 0.00 DECANE KG/HR 0.00 3395.48 3395.48 0.00 10330.25 0.00 UNDECANE KG/HR 0.00 15827.4

015827.40 0.00 10294.20 0.00

DODECANE KG/HR 0.00 17799.65

17799.65 0.00 10186.76 0.00

TRIDECAN KG/HR 0.00 17654.94

17654.94 0.00 10005.16 0.00

TETDECAN KG/HR 0.00 17416.81

17416.81 0.00 9766.94 0.00

PENDECAN KG/HR 0.00 17138.31

17138.31 0.00 9488.44 0.00

HXDECANE KG/HR 0.00 16824.86

16824.86 0.00 9174.97 0.00

HDECANE KG/HR 0.00 16488.09

16488.09 0.00 8838.64 0.00

OCTDECAN KG/HR 0.00 16134.75

16134.75 0.00 8484.84 0.00

NONDECAN KG/HR 0.00 15766.37

15766.37 0.00 8116.46 0.00

EICOSANE KG/HR 0.00 15396.17

15396.17 0.00 7746.25 0.00

c2125 KG/HR 0.00 315.15 315.15 0.00 33220.77 0.00 c2629 KG/HR 0.00 30.78 30.78 0.00 20447.88 0.00 c3035 KG/HR 0.00 5.24 5.24 0.00 22276.60 0.00 c3647 KG/HR 0.00 0.34 0.34 0.00 24086.77 0.00 c48+ KG/HR 0.00 0.00 0.00 0.00 15351.40 0.00 AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

Stream Name HEAVIES

HEAVY_FEED

HI_FEED

HI_FEED_P

HI_PROD

HP_STEAM

Temperature C 198.42 123.26 191.76 192.12 192.12 254.45Pressure ATM 27.10 2.00 1.35 27.50 27.50 41.83

Vapor Act. Vol. Rate M3/HR

n/a 15472.03 n/a n/a 6.78 35106.19

Liquid Act. Rate (vol) M3/HR

280.72 329.35 267.36 263.92 315.59 n/a


277840.83 185840.73

185840.73 185844.34

746540.51


KG/HR

HYDROGEN KG/HR 7.51 15.91 0.08 0.08 0.08 0.00

H2O KG/HR 1510.78 1510.80 58.13 58.13 58.13 746540.51

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 8.26 22.54 0.11 0.11 0.11 0.00

CO2 KG/HR 862.59 8373.46 26.63 26.63 26.63 0.00

CO KG/HR 68.01 222.22 0.96 0.96 0.96 0.00

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METHANE KG/HR 130.27 1784.49 3.11 3.11 1153.47 0.00

ETHANE KG/HR 15.58 779.02 0.76 0.76 575.94 0.00

PROPANE KG/HR 38.55 4911.75 3.19 3.19 4029.43 0.00

BUTANE KG/HR 85.41 6509.95 11.70 11.70 4037.93 0.00

PENTANE KG/HR 281.85 11912.64 60.31 60.31 4853.45 0.00

HEXANE KG/HR 494.32 13517.29 157.16 157.16 4950.31 0.00

HEPTANE KG/HR 810.05 14360.18 353.05 353.05 5146.24 0.00

OCTANE KG/HR 1291.54 14811.18 727.76 727.76 5520.94 0.00

NONANE KG/HR 1963.37 15037.35 1340.09 1340.09 6133.28 0.00

DECANE KG/HR 2827.20 15120.44 2205.69 2205.69 6998.87 0.00

UNDECANE KG/HR 3829.59 17943.29 3265.02 3265.02 10914.92 0.00

DODECANE KG/HR 4958.06 17836.45 4491.40 4491.40 12141.29 0.00

TRIDECAN KG/HR 5990.18 17655.00 5631.16 5631.16 13281.04 0.00

TETDECAN KG/HR 6849.10 17416.81 6591.71 6591.71 14241.58 0.00

PENDECAN KG/HR 7462.47 17138.31 7287.97 7287.97 14937.85 0.00

HXDECANE KG/HR 7818.64 16824.86 7705.42 7705.42 15355.30 0.00

HDECANE KG/HR 7933.86 16488.09 7860.98 7860.98 15510.42 0.00

OCTDECAN KG/HR 7884.49 16134.75 7838.03 7838.03 15487.93 0.00

NONDECAN KG/HR 7707.47 15766.37 7676.78 7676.78 15326.69 0.00

EICOSANE KG/HR 7527.29 15396.17 7511.63 7511.63 15161.55 0.00

c2125 KG/HR 32924.66 315.15 32905.62

32905.62 0.00 0.00

c2629 KG/HR 20418.66 30.78 20417.10

20417.10 0.00 0.00

c3035 KG/HR 22271.58 5.24 22271.36

22271.36 0.00 0.00

c3647 KG/HR 24086.45 0.34 24086.44

24086.44 0.00 0.00

c48+ KG/HR 15351.40 0.00 15351.40

15351.40 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

Stream Name HP_STM_FUEL

HPSTEAM_1

HPSTEAM1_SAT

LIGHT_FEED

LIGHTS LIGHTS_1

Temperature C 10.00 254.44 253.81 -29.95 198.42 -0.02Pressure ATM 27.00 41.83 41.83 27.00 27.10 27.10


6123.50 40592.17 26679.85 n/a 32836.95

7788.89

Liquid Act. Rate (vol)

M3/HR

n/a n/a 370.97 19.40 n/a n/a

Total Mass Rate KG/HR

199114.22 863208.34 863208.34 14347.23 587979.95

264206.02


KG/HR

HYDROGEN KG/HR

3427.86 0.00 0.00 4.03 4692.29 4683.89

H2O KG/HR

2.55 863208.34 863208.34 0.01 239386.30

40.42

OXYGEN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR

3057.94 0.00 0.00 6.43 4187.22 4172.94

CO2 KG/ 133734.80 0.00 0.00 5087.49 187547. 180036.3

20

HR 22 5 CO KG/

HR23205.70 0.00 0.00 59.16 31780.5

831626.37

METHANE KG/HR

27210.44 0.00 0.00 231.90 36236.07

35732.21

ETHANE KG/HR

1883.27 0.00 0.00 135.19 2149.99 1961.72

PROPANE KG/HR

3924.55 0.00 0.00 1279.38 3137.22 2290.26

BUTANE KG/HR

2099.79 0.00 0.00 3001.19 4100.55 1702.25

PENTANE KG/HR

540.33 0.00 0.00 3426.98 8152.73 1315.08

HEXANE KG/HR

25.58 0.00 0.00 874.60 8681.98 452.15

HEPTANE KG/HR

1.34 0.00 0.00 184.15 8895.87 138.93

OCTANE KG/HR

0.07 0.00 0.00 41.99 8765.36 38.90

NONANE KG/HR

0.00 0.00 0.00 10.69 8291.33 10.53

DECANE KG/HR

0.00 0.00 0.00 2.99 7503.05 2.98

UNDECANE KG/HR

0.00 0.00 0.00 0.80 6464.61 0.80

DODECANE KG/HR

0.00 0.00 0.00 0.20 5228.70 0.20

TRIDECAN KG/HR

0.00 0.00 0.00 0.04 4014.98 0.04

TETDECAN KG/HR

0.00 0.00 0.00 0.01 2917.84 0.01

PENDECAN KG/HR

0.00 0.00 0.00 0.00 2025.96 0.00

HXDECANE KG/HR

0.00 0.00 0.00 0.00 1356.33 0.00

HDECANE KG/HR

0.00 0.00 0.00 0.00 904.78 0.00

OCTDECAN KG/HR

0.00 0.00 0.00 0.00 600.35 0.00

NONDECAN KG/HR

0.00 0.00 0.00 0.00 408.98 0.00

EICOSANE KG/HR

0.00 0.00 0.00 0.00 218.96 0.00

c2125 KG/HR

0.00 0.00 0.00 0.00 296.12 0.00

c2629 KG/HR

0.00 0.00 0.00 0.00 29.22 0.00

c3035 KG/HR

0.00 0.00 0.00 0.00 5.02 0.00

c3647 KG/HR

0.00 0.00 0.00 0.00 0.32 0.00

c48+ KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

Stream Name LP_STEAM

LPG_PROD

METHANE

NAPHTHA

NAPHTHA_1

NAPHTHA_2

Temperature C 126.67 43.33 37.78 128.91 191.76 -0.02Pressure ATM 2.36 11.97 35.02 1.75 1.35 27.10

Vapor Act. Vol. Rate M3/HR 151506.56 n/a 17315.99 26749.56 4332.44 n/a

21


M3/HR n/a 12.41 n/a n/a n/a 116.77

Total Mass Rate KG/HR 201872.40 6892.87 400471.12 107637.89 7568.42 84428.07


KG/HR

HYDROGEN KG/HR 0.00 0.00 0.00 15.91 7.43 8.41

H2O KG/HR 201872.40 0.00 0.00 1510.80 1452.65 0.02

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 0.00 0.00 0.00 22.54 8.15 14.28

CO2 KG/HR 0.00 53.79 0.00 8373.46 835.96 7510.87

CO KG/HR 0.00 0.00 0.00 222.22 67.05 154.22

METHANE KG/HR 0.00 0.00 400471.12 1784.49 127.16 503.86

ETHANE KG/HR 0.00 34.22 0.00 779.02 14.82 188.27

PROPANE KG/HR 0.00 1182.39 0.00 4911.75 35.36 846.96

BUTANE KG/HR 0.00 2716.56 0.00 6509.95 73.71 2398.30

PENTANE KG/HR 0.00 2495.93 0.00 11912.64 221.54 6837.65

HEXANE KG/HR 0.00 372.88 0.00 13517.29 337.15 8229.83

HEPTANE KG/HR 0.00 34.23 0.00 14360.18 457.00 8756.94

OCTANE KG/HR 0.00 2.64 0.00 14811.17 563.78 8726.46

NONANE KG/HR 0.00 0.21 0.00 15028.76 623.27 8280.81

DECANE KG/HR 0.00 0.02 0.00 11724.96 621.51 7500.07

UNDECANE KG/HR 0.00 0.00 0.00 2115.89 564.57 6463.80

DODECANE KG/HR 0.00 0.00 0.00 36.80 466.66 5228.50

TRIDECAN KG/HR 0.00 0.00 0.00 0.06 359.03 4014.93

TETDECAN KG/HR 0.00 0.00 0.00 0.00 257.40 2917.83

PENDECAN KG/HR 0.00 0.00 0.00 0.00 174.50 2025.96

HXDECANE KG/HR 0.00 0.00 0.00 0.00 113.22 1356.33

HDECANE KG/HR 0.00 0.00 0.00 0.00 72.88 904.78

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 46.46 600.35

NONDECAN KG/HR 0.00 0.00 0.00 0.00 30.69 408.98

EICOSANE KG/HR 0.00 0.00 0.00 0.00 15.66 218.96

c2125 KG/HR 0.00 0.00 0.00 0.00 19.04 296.12

c2629 KG/HR 0.00 0.00 0.00 0.00 1.56 29.22

c3035 KG/HR 0.00 0.00 0.00 0.00 0.22 5.02

c3647 KG/HR 0.00 0.00 0.00 0.00 0.01 0.32

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

Stream Name NAPHTHA_3

NAPHTHA_4

NAPHTHA_PROD

O2 PROP_VAP_P1

PROP_VAP_P2


Vapor Act. Vol. Rate

M3/HR n/a n/a n/a 10232.75

12943.51 4563.74


M3/HR 3.93 3.11 126.31 n/a n/a n/a

Total Mass Rate KG/HR 2470.40 1953.32 85148.74 478648.05

83195.84 83195.84

22


KG/HR

HYDROGEN KG/HR 0.00 0.00 0.02 0.00 0.00 0.00

H2O KG/HR 0.64 0.00 11.76 0.00 0.00 0.00

OXYGEN KG/HR 0.00 0.00 0.00 474452.57

0.00 0.00

NITROGEN KG/HR 0.00 0.00 0.04 4195.47 0.00 0.00

CO2 KG/HR 12.17 0.18 121.98 0.00 0.00 0.00

CO KG/HR 0.05 0.00 0.45 0.00 0.00 0.00

METHANE KG/HR 1.09 0.00 9.79 0.00 0.00 0.00

ETHANE KG/HR 2.51 0.28 25.96 0.00 0.00 0.00

PROPANE KG/HR 48.33 32.95 578.92 0.00 83195.84 83195.84

BUTANE KG/HR 194.00 283.37 2615.72 0.00 0.00 0.00

PENTANE KG/HR 643.10 931.05 9990.64 0.00 0.00 0.00

HEXANE KG/HR 701.20 501.71 13562.05 0.00 0.00 0.00

HEPTANE KG/HR 483.57 149.92 14463.20 0.00 0.00 0.00

OCTANE KG/HR 251.52 39.35 14847.38 0.00 0.00 0.00

NONANE KG/HR 101.12 10.47 15039.09 0.00 0.00 0.00

DECANE KG/HR 29.22 2.97 11727.93 0.00 0.00 0.00

UNDECANE KG/HR 1.86 0.80 2116.69 0.00 0.00 0.00

DODECANE KG/HR 0.01 0.20 37.00 0.00 0.00 0.00

TRIDECAN KG/HR 0.00 0.04 0.10 0.00 0.00 0.00

TETDECAN KG/HR 0.00 0.01 0.01 0.00 0.00 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

Stream Name

PROPANE

PROPANE_COND

PROPANE_VAP

R134A_1

R134A_1_COND

R134A_1_VAP

Temperature C -41.85 26.64 -40.85 -9.71 24.23 -8.71Pressure ATM 1.00 9.90 1.00 2.00 6.48 2.00


M3/HR

14100.12 n/a 34824.34 22077.89

n/a 96822.07


M3/HR

84.10 169.52 n/a 557.41 795.50 n/a

Total Mass Rate

KG/HR

83195.84 83195.84 83195.84 961857.57

961857.57 961857.57


KG/HR

HYDROGE KG/ 0.00 0.00 0.00 0.00 0.00 0.00

23

N HR H2O KG/

HR0.00 0.00 0.00 0.00 0.00 0.00

OXYGEN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

CO2 KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

CO KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

METHANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

ETHANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

PROPANE KG/HR

83195.84 83195.84 83195.84 0.00 0.00 0.00

BUTANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

PENTANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

HEXANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

HEPTANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

OCTANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

NONANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

DECANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

UNDECANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

DODECANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

TRIDECAN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

TETDECAN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

PENDECAN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

NONDECAN

KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

c2125 KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

c48+ KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR

0.00 0.00 0.00 961857.57

961857.57 961857.57

24

Stream Name R134A_2

R134A_2_COND

R134A_2_P

R134A_2_VAP

R134A_VAP_P1

R134A_VAP_P2

Temperature C -25.90 26.67 48.56 -24.90 13.69 37.43Pressure ATM 1.00 6.97 6.97 1.00 3.60 6.48


M3/HR

31121.31

n/a 16193.12 94446.08 56892.61 32940.74


M3/HR

236.65 407.15 n/a n/a n/a n/a

Total Mass Rate KG/HR

488400.70

488400.70 488400.70

488400.70 961857.57 961857.57


KG/HR

HYDROGEN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

H2O KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

OXYGEN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

CO2 KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

CO KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

METHANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

ETHANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

PROPANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

BUTANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

PENTANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

HEXANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

HEPTANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

OCTANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

NONANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

DECANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

UNDECANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

DODECANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

TRIDECAN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

TETDECAN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

PENDECAN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

NONDECAN KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/ 0.00 0.00 0.00 0.00 0.00 0.00

25

HR c2125 KG/

HR0.00 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

c48+ KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR

0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR

488400.70

488400.70 488400.70

488400.70 961857.57 961857.57

Stream Name S1 S10 S11 S12 S13 S14 S15

Temperature C 167.95 200.00 132.22 43.33 -0.02 10.03 -29.95Pressure ATM 28.22 28.10 27.10 27.10 27.10 27.00 27.00


47050.15 31623.51 14501.05 9532.99 7788.89 8569.14 6846.54


M3/HR

78.15 454.93 281.55 342.98 355.83 0.05 19.53


930288.35

587979.95

587979.95

587979.95

287244.00

287244.00


KG/HR

HYDROGEN KG/HR 0.00 4699.80 4692.29 4692.29 4692.29 4699.78 4699.78

H2O KG/HR 247750.63

389796.33

239386.30

239386.30

239386.30

138.03 138.03

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 0.00 4195.47 4187.22 4187.22 4187.22 4195.44 4195.44

CO2 KG/HR 111898.44

188409.80

187547.22

187547.22

187547.22

188287.88

188287.88

CO KG/HR 0.00 31848.59 31780.58 31780.58 31780.58 31848.14 31848.14

METHANE KG/HR 400471.12

36366.34 36236.07 36236.07 36236.07 37506.89 37506.89

ETHANE KG/HR 0.00 2165.57 2149.99 2149.99 2149.99 2715.04 2715.04

PROPANE KG/HR 0.00 3175.77 3137.22 3137.22 3137.22 6655.53 6655.53

BUTANE KG/HR 0.00 4185.96 4100.55 4100.55 4100.55 5877.65 5877.65

PENTANE KG/HR 0.00 8434.58 8152.73 8152.73 8152.73 4167.17 4167.17

HEXANE KG/HR 0.00 9176.30 8681.98 8681.98 8681.98 909.64 909.64

HEPTANE KG/HR 0.00 9705.92 8895.87 8895.87 8895.87 185.98 185.98

OCTANE KG/HR 0.00 10056.90 8765.36 8765.36 8765.36 42.09 42.09

NONANE KG/HR 0.00 10254.70 8291.33 8291.33 8291.33 10.69 10.69

DECANE KG/HR 0.00 10330.25 7503.05 7503.05 7503.05 2.99 2.99

UNDECANE KG/HR 0.00 10294.20 6464.61 6464.61 6464.61 0.80 0.80

DODECANE KG/HR 0.00 10186.76 5228.70 5228.70 5228.70 0.20 0.20

TRIDECAN KG/HR 0.00 10005.16 4014.98 4014.98 4014.98 0.04 0.04

TETDECAN KG/HR 0.00 9766.94 2917.84 2917.84 2917.84 0.01 0.01

PENDECAN KG/HR 0.00 9488.44 2025.96 2025.96 2025.96 0.00 0.00

HXDECANE KG/HR 0.00 9174.97 1356.33 1356.33 1356.33 0.00 0.00

HDECANE KG/HR 0.00 8838.64 904.78 904.78 904.78 0.00 0.00

OCTDECAN KG/HR 0.00 8484.84 600.35 600.35 600.35 0.00 0.00

26

NONDECAN KG/HR 0.00 8116.46 408.98 408.98 408.98 0.00 0.00

EICOSANE KG/HR 0.00 7746.25 218.96 218.96 218.96 0.00 0.00

c2125 KG/HR 0.00 33220.77 296.12 296.12 296.12 0.00 0.00

c2629 KG/HR 0.00 20447.88 29.22 29.22 29.22 0.00 0.00

c3035 KG/HR 0.00 22276.60 5.02 5.02 5.02 0.00 0.00

c3647 KG/HR 0.00 24086.77 0.32 0.32 0.32 0.00 0.00

c48+ KG/HR 0.00 15351.40 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

Stream Name S16 S17 S19 S2 S26 S34


Vapor Act. Vol. Rate M3/HR 8388.44 6128.41 2703.70 400644.42 24754.04 9350.63

Liquid Act. Rate (vol) M3/HR n/a 4.18 n/a n/a 314.87 118.63



KG/HR

HYDROGEN KG/HR 4695.75 15.95 0.00 0.00 15.91 15.91

H2O KG/HR 3.49 626.84 0.00 603897.71 1510.80 1510.80

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 4189.00 22.61 0.00 4195.47 22.54 22.54

CO2 KG/HR 183200.39 8497.51 53.79 111898.44 8373.46 8373.46

CO KG/HR 31788.98 222.98 0.00 276881.37 222.22 222.22

METHANE KG/HR 37274.99 1799.07 0.00 241890.28 1784.49 1784.49

ETHANE KG/HR 2579.85 802.02 34.22 0.00 779.02 779.02

PROPANE KG/HR 5376.15 5170.24 1182.39 0.00 4911.75 4911.75

BUTANE KG/HR 2876.46 6562.84 2716.56 0.00 6509.95 6509.95

PENTANE KG/HR 740.19 7744.35 2495.93 0.00 11912.64 11912.64

HEXANE KG/HR 35.04 3262.14 372.88 0.00 13517.29 13517.29

HEPTANE KG/HR 1.83 1104.91 34.23 0.00 14360.18 14360.18

OCTANE KG/HR 0.10 363.82 2.64 0.00 14811.18 14811.17

NONANE KG/HR 0.01 116.88 0.21 0.00 15037.35 15028.76

DECANE KG/HR 0.00 30.89 0.02 0.00 15120.44 11724.96

UNDECANE KG/HR 0.00 1.90 0.00 0.00 17943.29 2115.89

DODECANE KG/HR 0.00 0.01 0.00 0.00 17836.45 36.80

TRIDECAN KG/HR 0.00 0.00 0.00 0.00 17655.00 0.06

TETDECAN KG/HR 0.00 0.00 0.00 0.00 17416.81 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00 17138.31 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00 16824.86 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00 16488.09 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 16134.75 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00 15766.37 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00 15396.17 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00 315.15 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00 30.78 0.00

27

c3035 KG/HR 0.00 0.00 0.00 0.00 5.24 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00 0.34 0.00

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00




Liquid Act. Rate (vol) M3/HR n/a n/a n/a n/a n/a n/a



KG/HR

HYDROGEN KG/HR 15.90 15.95 0.00 91873.46 0.00 15.95

H2O KG/HR 97.61 626.77 0.00 309357.55 0.00 626.77

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 22.50 22.61 0.00 4195.47 0.00 22.61

CO2 KG/HR 8251.54 8497.63 53.79 188409.80 53.79 8497.63

CO KG/HR 221.77 222.98 0.00 637459.08 0.00 222.98

METHANE KG/HR 1774.68 1799.08 0.00 7482.52 0.00 1799.08

ETHANE KG/HR 753.32 802.05 34.22 0.00 34.22 802.05

PROPANE KG/HR 4365.27 5170.74 1182.39 0.00 1182.39 5170.74

BUTANE KG/HR 4175.40 6565.03 2716.56 0.00 2716.56 6565.03

PENTANE KG/HR 2852.09 7745.32 2495.93 0.00 2495.93 7745.32

HEXANE KG/HR 457.49 3261.60 372.88 0.00 372.88 3261.60

HEPTANE KG/HR 47.05 1104.76 34.23 0.00 34.23 1104.76

OCTANE KG/HR 3.19 363.77 2.64 0.00 2.64 363.77

NONANE KG/HR 0.17 116.86 0.21 0.00 0.21 116.86

DECANE KG/HR 0.01 30.88 0.02 0.00 0.02 30.88

UNDECANE KG/HR 0.00 1.90 0.00 0.00 0.00 1.90

DODECANE KG/HR 0.00 0.01 0.00 0.00 0.00 0.01

TRIDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

TETDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

28

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

Stream Name S5 S50 S56 S6 S64

Temperature C 1032.62 97.73 39.83 265.54 43.33Pressure ATM 28.22 9.00 1.75 28.22 1.00

Vapor Act. Vol. Rate M3/HR 344508.05 2330.28 10003.43 142217.80 n/a

Liquid Act. Rate (vol) M3/HR n/a n/a 133.88 n/a 3036.98

Total Mass Rate KG/HR 1238777.90 33629.40 120419.37 1238777.90 3009340.96


KG/HR

HYDROGEN KG/HR 91873.46 15.95 15.97 91873.46 0.00

H2O KG/HR 309357.55 381.04 1514.53 309357.55 3009340.96

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 4195.47 22.60 22.64 4195.47 0.00

CO2 KG/HR 188409.80 8485.34 8619.43 188409.80 0.00

CO KG/HR 637459.08 222.93 223.43 637459.08 0.00

METHANE KG/HR 7482.52 1797.98 1808.88 7482.52 0.00

ETHANE KG/HR 0.00 799.52 827.73 0.00 0.00

PROPANE KG/HR 0.00 5121.91 5716.72 0.00 0.00

BUTANE KG/HR 0.00 6368.84 8897.38 0.00 0.00

PENTANE KG/HR 0.00 7101.25 16804.91 0.00 0.00

HEXANE KG/HR 0.00 2560.93 16321.93 0.00 0.00

HEPTANE KG/HR 0.00 621.34 15418.04 0.00 0.00

OCTANE KG/HR 0.00 112.30 15171.80 0.00 0.00

NONANE KG/HR 0.00 15.77 15145.48 0.00 0.00

DECANE KG/HR 0.00 1.66 11755.84 0.00 0.00

UNDECANE KG/HR 0.00 0.04 2117.79 0.00 0.00

DODECANE KG/HR 0.00 0.00 36.81 0.00 0.00

TRIDECAN KG/HR 0.00 0.00 0.06 0.00 0.00

TETDECAN KG/HR 0.00 0.00 0.00 0.00 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00 0.00

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00

29

Stream Name S65 S67 S68 S69 S7

Temperature C 15.56 37.78 15.56 37.78 43.33Pressure ATM 1.00 1.00 1.00 1.00 28.22

Vapor Act. Vol. Rate M3/HR n/a n/a n/a n/a 68227.04

Liquid Act. Rate (vol) M3/HR 1135.62 1142.44 506.58 509.62 307.71

Total Mass Rate KG/HR 1134493.83 1134493.83 506072.56 506072.56 1238777.90


KG/HR

HYDROGEN KG/HR 0.00 0.00 0.00 0.00 91873.46

H2O KG/HR 1134493.83 1134493.83 506072.56 506072.56 309357.55

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 0.00 0.00 0.00 0.00 4195.47

CO2 KG/HR 0.00 0.00 0.00 0.00 188409.80

CO KG/HR 0.00 0.00 0.00 0.00 637459.08

METHANE KG/HR 0.00 0.00 0.00 0.00 7482.52

ETHANE KG/HR 0.00 0.00 0.00 0.00 0.00

PROPANE KG/HR 0.00 0.00 0.00 0.00 0.00

BUTANE KG/HR 0.00 0.00 0.00 0.00 0.00

PENTANE KG/HR 0.00 0.00 0.00 0.00 0.00

HEXANE KG/HR 0.00 0.00 0.00 0.00 0.00

HEPTANE KG/HR 0.00 0.00 0.00 0.00 0.00

OCTANE KG/HR 0.00 0.00 0.00 0.00 0.00

NONANE KG/HR 0.00 0.00 0.00 0.00 0.00

DECANE KG/HR 0.00 0.00 0.00 0.00 0.00

UNDECANE KG/HR 0.00 0.00 0.00 0.00 0.00

DODECANE KG/HR 0.00 0.00 0.00 0.00 0.00

TRIDECAN KG/HR 0.00 0.00 0.00 0.00 0.00

TETDECAN KG/HR 0.00 0.00 0.00 0.00 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00 0.00

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00

Stream Name S72 S73 S74 S76 S77 S78 S79

Temperature C 15.56 37.80 66.30 43.33 43.33 43.33 43.33Pressure ATM 1.00 1.00 27.10 3.00 9.00 9.00 9.00

Vapor Act. Vol. Rate M3/HR n/a n/a n/a 6128.41 1511.65 n/a 1511.65

30


M3/HR 33.38 33.58 466.82 n/a 17.82 17.53 n/a


33343.67

457952.91

33629.40

33629.40

10311.08

23037.98


KG/HR

HYDROGEN KG/HR 0.00 0.00 0.00 15.95 15.95 0.05 15.90

H2O KG/HR 33343.67

33343.67

457952.91

381.04 381.04 3.09 97.61

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 0.00 0.00 0.00 22.60 22.60 0.10 22.50

CO2 KG/HR 0.00 0.00 0.00 8485.34 8485.34 233.80 8251.54

CO KG/HR 0.00 0.00 0.00 222.93 222.93 1.16 221.77

METHANE KG/HR 0.00 0.00 0.00 1797.98 1797.98 23.30 1774.68

ETHANE KG/HR 0.00 0.00 0.00 799.52 799.52 46.20 753.32

PROPANE KG/HR 0.00 0.00 0.00 5121.91 5121.91 756.64 4365.27

BUTANE KG/HR 0.00 0.00 0.00 6368.84 6368.84 2193.44 4175.40

PENTANE KG/HR 0.00 0.00 0.00 7101.25 7101.25 4249.16 2852.09

HEXANE KG/HR 0.00 0.00 0.00 2560.93 2560.93 2103.44 457.49

HEPTANE KG/HR 0.00 0.00 0.00 621.34 621.34 574.29 47.05

OCTANE KG/HR 0.00 0.00 0.00 112.30 112.30 109.11 3.19

NONANE KG/HR 0.00 0.00 0.00 15.77 15.77 15.60 0.17

DECANE KG/HR 0.00 0.00 0.00 1.66 1.66 1.66 0.01

UNDECANE KG/HR 0.00 0.00 0.00 0.04 0.04 0.04 0.00

DODECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

TRIDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

TETDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00 0.00



Vapor Act. Vol. Rate M3/HR 58717.55 n/a n/a n/a n/a n/a


M3/HR 308.66 138.97 1905.49 525.74 137.84 526.29


Total Weight Comp. KG/HR

31

Rates HYDROGEN KG/HR 91873.46 0.00 0.00 0.00 0.00 0.00

H2O KG/HR 309357.55 137705.58 1900000.00 525218.39 137705.58 525218.39

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 4195.47 0.00 0.00 0.00 0.00 0.00

CO2 KG/HR 188409.80 0.00 0.00 0.00 0.00 0.00

CO KG/HR 637459.08 0.00 0.00 0.00 0.00 0.00

METHANE KG/HR 7482.52 0.00 0.00 0.00 0.00 0.00

ETHANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

PROPANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

BUTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

PENTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HEXANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HEPTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

OCTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NONANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

DECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

UNDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

DODECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

TRIDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

TETDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00


Temperature C -25.90 -24.90 1.00 15.72 123.17 15.56Pressure ATM 1.00 1.00 28.15 20.34 23.68 1.00

Vapor Act. Vol. Rate M3/HR n/a 8357.11 58857.69 n/a n/a n/a

Liquid Act. Rate (vol) M3/HR 31.30 n/a n/a 1442.21 19.77 747.28



KG/HR

HYDROGEN KG/HR 0.00 0.00 91873.46 0.00 0.00 0.00

H2O KG/HR 0.00 0.00 303.88 1442047.09 0.00 746540.51

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

32

NITROGEN KG/HR 0.00 0.00 4195.47 0.00 0.00 0.00

CO2 KG/HR 0.00 0.00 188409.80 0.00 53.97 0.00

CO KG/HR 0.00 0.00 637459.08 0.00 0.00 0.00

METHANE KG/HR 0.00 0.00 7482.52 0.00 0.00 0.00

ETHANE KG/HR 0.00 0.00 0.00 0.00 34.50 0.00

PROPANE KG/HR 0.00 0.00 0.00 0.00 1215.34 0.00

BUTANE KG/HR 0.00 0.00 0.00 0.00 2999.93 0.00

PENTANE KG/HR 0.00 0.00 0.00 0.00 3426.98 0.00

HEXANE KG/HR 0.00 0.00 0.00 0.00 874.60 0.00

HEPTANE KG/HR 0.00 0.00 0.00 0.00 184.15 0.00

OCTANE KG/HR 0.00 0.00 0.00 0.00 41.99 0.00

NONANE KG/HR 0.00 0.00 0.00 0.00 10.69 0.00

DECANE KG/HR 0.00 0.00 0.00 0.00 2.99 0.00

UNDECANE KG/HR 0.00 0.00 0.00 0.00 0.80 0.00

DODECANE KG/HR 0.00 0.00 0.00 0.00 0.20 0.00

TRIDECAN KG/HR 0.00 0.00 0.00 0.00 0.04 0.00

TETDECAN KG/HR 0.00 0.00 0.00 0.00 0.01 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 43216.41 43216.41 0.00 0.00 0.00 0.00

Stream Name STEAM SYN_FUEL

SYNGAS SYNGAS_FEED

TAIL_1 TAIL_2

Temperature C 231.21 10.00 1065.55 537.78 -20.67 -29.95Pressure ATM 28.22 27.00 28.22 28.22 23.00 27.00



M3/HR n/a n/a n/a n/a n/a n/a


73648.02 1238777.90

760120.18 5501.04 272762.24


KG/HR

HYDROGEN KG/HR 0.00 1267.89 91873.46 0.00 4.03 4695.75

H2O KG/HR 247750.63

0.94 309357.55 247750.63 0.01 3.49

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 0.00 1131.07 4195.47 0.00 6.43 4189.00

CO2 KG/HR 0.00 49465.59 188409.80 111898.44 5033.52 183200.39

33

CO KG/HR 0.00 8583.28 637459.08 0.00 59.16 31788.98

METHANE KG/HR 0.00 10064.55 7482.52 400471.12 231.90 37274.99

ETHANE KG/HR 0.00 696.58 0.00 0.00 100.69 2579.85

PROPANE KG/HR 0.00 1451.61 0.00 0.00 64.04 5376.15

BUTANE KG/HR 0.00 776.67 0.00 0.00 1.26 2876.46

PENTANE KG/HR 0.00 199.86 0.00 0.00 0.00 740.19

HEXANE KG/HR 0.00 9.46 0.00 0.00 0.00 35.04

HEPTANE KG/HR 0.00 0.49 0.00 0.00 0.00 1.83

OCTANE KG/HR 0.00 0.03 0.00 0.00 0.00 0.10

NONANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.01

DECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

UNDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

DODECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

TRIDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

TETDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

Stream Name WATER_1 WATER_2 WATER_3 WATER_4 WATER_5 WATER_6

Temperature C 1.00 198.42 -0.02 43.33 43.33 -29.95Pressure ATM 28.15 27.10 27.10 3.00 9.00 27.00

Vapor Act. Vol. Rate M3/HR n/a n/a n/a n/a n/a n/a

Liquid Act. Rate (vol) M3/HR 308.66 171.65 239.07 0.25 0.28 0.13



KG/HR

HYDROGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

H2O KG/HR 309053.67 148899.25 239345.86 245.15 280.34 134.53

OXYGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NITROGEN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

CO2 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

CO KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

METHANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

ETHANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

PROPANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

34

BUTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

PENTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HEXANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HEPTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

OCTANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NONANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

DECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

UNDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

DODECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

TRIDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

TETDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3035 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

c48+ KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00 0.00 0.00

Stream Name WATER_7 WATER_DIST_1 WATER_DIST_2 WATER_MAKEUP

Temperature C 31.22 -20.67 128.91 15.56Pressure ATM 1.00 23.00 1.75 1.00

Vapor Act. Vol. Rate M3/HR n/a n/a n/a n/a

Liquid Act. Rate (vol) M3/HR 0.88 n/a n/a 1443.48

Total Mass Rate KG/HR 876.01 n/a n/a 1442047.09


KG/HR

HYDROGEN KG/HR 0.00 0.00 0.00 0.00

H2O KG/HR 876.01 0.00 0.00 1442047.09

OXYGEN KG/HR 0.00 0.00 0.00 0.00

NITROGEN KG/HR 0.00 0.00 0.00 0.00

CO2 KG/HR 0.00 0.00 0.00 0.00

CO KG/HR 0.00 0.00 0.00 0.00

METHANE KG/HR 0.00 0.00 0.00 0.00

ETHANE KG/HR 0.00 0.00 0.00 0.00

PROPANE KG/HR 0.00 0.00 0.00 0.00

BUTANE KG/HR 0.00 0.00 0.00 0.00

PENTANE KG/HR 0.00 0.00 0.00 0.00

HEXANE KG/HR 0.00 0.00 0.00 0.00

HEPTANE KG/HR 0.00 0.00 0.00 0.00

35

OCTANE KG/HR 0.00 0.00 0.00 0.00

NONANE KG/HR 0.00 0.00 0.00 0.00

DECANE KG/HR 0.00 0.00 0.00 0.00

UNDECANE KG/HR 0.00 0.00 0.00 0.00

DODECANE KG/HR 0.00 0.00 0.00 0.00

TRIDECAN KG/HR 0.00 0.00 0.00 0.00

TETDECAN KG/HR 0.00 0.00 0.00 0.00

PENDECAN KG/HR 0.00 0.00 0.00 0.00

HXDECANE KG/HR 0.00 0.00 0.00 0.00

HDECANE KG/HR 0.00 0.00 0.00 0.00

OCTDECAN KG/HR 0.00 0.00 0.00 0.00

NONDECAN KG/HR 0.00 0.00 0.00 0.00

EICOSANE KG/HR 0.00 0.00 0.00 0.00

c2125 KG/HR 0.00 0.00 0.00 0.00

c2629 KG/HR 0.00 0.00 0.00 0.00

c3035 KG/HR 0.00 0.00 0.00 0.00

c3647 KG/HR 0.00 0.00 0.00 0.00

c48+ KG/HR 0.00 0.00 0.00 0.00

AIR KG/HR 0.00 0.00 0.00 0.00

R134A KG/HR 0.00 0.00 0.00 0.00

36

Process DescriptionThis section will describe the various process changes that occur throughout the flow sheet. They will not be described according to their placement in the flow sheet, but instead are lumped together under their respective process changes. The reactors are significant pieces of equipment and will be discussed individually at length.

Syngas ReactorThe syngas reactor processes 248,000 kg/hr of water, 112,000 kg/hr of carbon dioxide, 400,000 kg/hr of methane, and 479,000 kg/hr of oxygen at 28 atm and approximately 535°C. The oxygen reacts with the methane to produce a significant amount of energy to raise the temperature in the reactor to 1065°C and to drive the steam reforming reaction. The three reactions are given, along with their associated changes in enthalpy, in Table 5. It can be seen that the steam reforming reaction is endothermic, the shift reaction is moderately exothermic, and the partial oxidation of methane is extremely exothermic.

TABLE 5 - REACTIONS OCCURRING IN THE SYNGAS REACTOR

Chemical Reaction Enthalpy Change Equilibrium Constant at 1065

°CSteam

ReformingCH 4+H 2O⟷CO+3H 2 ΔH R=226.1kJ /mol

[8]K=0.036 [8]

Shift Reaction

CO+H 2O⟷CO2+H2 ΔH R=−41kJ /mol [8]

K=28.27 [8]

Partial Oxidation CH 4+

32O2⟶CO+2H 2O

ΔH R=−346.6kJ /mol K ⟶∞

The partial oxidation reaction continues to completion and drives the endothermic reforming reaction. Approximately 40% of the methane reacts in the partial oxidation reaction, with the 96.9% of the rest reacting in the reforming reaction until it reaches equilibrium. This produces 1,240,000 kg/hr of syngas at 1065 °C and 28 atm. The exact composition can be found in the Stream Attributes Table.

The syngas reactor must achieve equilibrium, so the residence time in the reactor is likely to be long. In addition, the reactor is processing a large volume of gas, approximately 1,371,000 m3/hr. Therefore, the syngas reactor will be very large. Because the syngas reactor was not rigorously designed, an approximation was made that set the volume of the syngas reactor to that of the FTR. With this approximation, the syngas unit comprises of 15 reactors in parallel, each of which has 24,293 tubes with a diameter of 3.175 cm. The reactors are 18.3 m in length and have a diameter of 6.1 m. Due to the high temperatures that the reactors must withstand, it will be constructed out of 310 Stainless Steel.

The oxygen feed to the reactor must be well controlled because this fuels the exothermic reaction. If any failure were to occur, the oxygen supply would have to be cut off so that the reaction self-extinguishes by using all of the available energy. Additionally, since the reaction raises the temperature of the feed entering and produces more moles than are consumed, there need to be pressure controls that relieve any build-up of pressure to prevent the reactor from failing.

37

Fischer-Tropsch ReactorThe FTR comprises the main reaction step that determines the efficiency of our process. It converts the carbon monoxide from the syngas unit into single-chain alkanes of various lengths. The reaction that occurs within the reactor is

2H 2+CO⟶ H❑2O+x (CH 2 )n H 2

Where, x represents some stoichiometric factor that depends on the length of the hydrocarbon chain. The feed enters the reactor at 225.85 °C and 28 atm at a flow rate of 930,000 kg/hr. The exact composition can be found in the Stream Attributes Table. The reaction’s stoichiometry depends on the temperature of the reactor according to a number of relations. The rate of reaction is best fit by a Langmuir-Hinshelwood form of

−rCO=k1T1 PH 2

PCO

(1+k2T 2PCO )2, ΔH R=−163.3kJ /mol

where

T 1=exp(−4492×( 1T −1473 ))

T 2=exp(8237×( 1T −1473 ))

k 1=0.0173gmolCO

hr ×cc cat×atm2

k 2=4.512at m−1

The pressures are partial pressures of the subscripted components, and the temperature is given in Kelvin. The selectivity of the products is very important. For methane, ethane, propane, and butane, the selectivity can be given by

SC H 4=0.03exp (−10,000×( 1T −

1473 ))

SC2 H6=SC3H 8

=SC4H 10=0.04SCH 4

The distribution of longer chain hydrocarbons follows the Anderson-Shulz-Flory probability distribution, which is given by

W n=n (1−α )2αn−1

α=0.93exp (250×( 1T −1473 ))

Where, Wn is the weight fraction of a hydrocarbon with n carbons.

There are 15 packed-bed reactors in parallel to accommodate this flow rate. Each reactor is 6.1 m in diameter and 18.3 m in length, containing 24,293 tubes with diameters of 3.175 cm. Each tube is packed with an ultra-stable cobalt-based catalyst. This catalyst has an equivalent diameter of 1.59 mm, a void fraction of 0.4, and a total packed density of 0.8 g/cc. This arrangement achieves a conversion of 95%. The reactor is extremely exothermic, but has a pressurized water jacket surrounding the

38

reactor so that the external water boils at 216°C, corresponding to a pressure of 21 atm. The reactors release a total of 3.42×109 kJ/hr. Each individual reactor releases

2.28×108kJ/hr. Since the enthalpy of vaporization of water at this temperature is 1,873 kJ/kg [9], the heat duty of the reactor requires 1.9×106 kg/hr of water through the cooling jackets of the fifteen reactors, with a flow rate of 35.3 L/s (127 m3/hr) per reactor. The pressure of the water jacket was determined to acceptably control the temperature of the reactors without quenching the reactor. The overall heat transfer coefficient depends on the process side flow rate as given by

U=0.385× G0.8

D0.2

where U is the overall heat transfer coefficient given in BTU /(hr ×° F× f t2), G is the

inlet gas mass velocity in g/ (cm2×hr ), and D is the tube diameter in cm. The MatLab code that was used to model the reactor system, found in Computer Programs, gave the temperature profile shown in Figure 8.

FIGURE 8 - TEMPERATURE PROFILE OF THE FISCHER-TROPSCH REACTORS

The reactor was modeled isothermally in PRO/II since the stoichiometry of the reaction depended on the temperature of the reactor, and there was no way to implement the temperature dependence in a manageable fashion in PRO/II. However, this does not discredit the results of the model because the reactor is relatively isothermal even when the heat transfer is modeled. Additionally, the model in PRO/II operates at the reaction-weighted temperature given by

39

T avg=∑i=1

N

rCO ,iT i

∑i=1

N

T i

Where, i represents a step along the reactor. Operating isothermally at this average temperature will give the same product distribution as evaluating the reaction products for every step in the reactor.

The FTR is extremely exothermic because many bonds are formed as carbon monoxide converts into hydrocarbon chains. If the temperature is not controlled well, the reactor will experience temperature run away, and this would pose a significant health and safety hazard not only to the operators but also to the surrounding community. In order to prevent this, a flow controller must be installed on the gas inlet to regulate the feed of gas entering the reactor at any one time. Additionally, there need to be pressure controls on the water jacket that can relieve pressure, bringing the temperature of the water down and quench the reaction if the temperature gets out of hand. This would require a temperature sensor inside the reactor that controls a pressure relief valve. Additionally, the flow of water through the jacket must be guaranteed to meet the heat duty, meaning that a flow controller should be installed on the water inlet to the reactor. In order to increase safety, a purge stream of inert nitrogen could also be set up so that in case the reaction exceeds a certain threshold temperature, the flow of the reactants will stop and the purge stream will flood the reactor, lowering the rate of reaction and stopping the reaction from occurring. These scenarios should not occur if the reactor runs according to the above specifications, but these contingency plans need to be in place in case of any disturbances in the reactor.

Hydroisomerization UnitHydroisomerization reaction is used to crack longer chains (C20

+¿¿) and yield lighter

hydrocarbons (C20−¿¿

). Heavier oil separated from the product of FTR will be fed into the HI unit. There, longer hydrocarbon chains are isomerized to form branched isomers and then the branched hydrocarbons are split by beta scission. The reactor itself will be a plug flow reactor with tubes inside. The tubes are to be filled with ssz-32 Zeolite which works best for the given situation and desired conversion. The conversion using this specific type of catalyst often goes up as high as 99%. For design purpose, 100% conversion was assumed and the product specifications were also assumed as the following: every 100 kg of C20

+¿¿ feed yields 1 kg ofC1, 0.5 kg ofC2, 3.5 kg ofC3, 3.5 kg of

C4, 25 kg of Naphtha and 66.5 kg of Diesel.

The size of the HI unit was determined from the amount of catalyst needed for the given feed flow rate. Typically, the contact time of hydroisomerization of heavy wax/paraffin with zeolite as the reaction catalyst is around 4000 kg mol/s for maximum conversion [10]. Then, with the feed flow rate of 605 kmol/hr and zeolite density of 5.25 g/cm3 [11], the reactor size was calculated as follows:

V=W o×Fo

ρ=4000

kgmols

×605kmolhr

5.25g

c m3

≈130m3

40

Where, Wo, Fo and ρ represent contact time, feed flow rate and density of catalyst, respectively.

The HI unit is operating at 192 °C and 27.5 atm. High pressure is necessary for catalyst life and high temperature is to maximize conversion. The hydroisomerization reaction is slightly exothermic. However, reactor temperature rise must be limited to 28 °C or below [12]. The reactor temperature is controlled by continuous feed of relatively cooler hydrogen gas.

PumpsPumps exist in the plant design to pressurize fluids. Other pumps will be needed to operate the plant, such as pumps on distillation columns to move the condensate back to the top of the column from the condenser, but only pumps with head of over 20 m will be discussed here, as these represent significant process changes. All of the pumps are powered by electricity.

REF_1The first significant pump in the process is called ‘Ref_1’. Although represented on the flow sheet as a single pump, the mass flow rate is too large for a single pump. The overall process change takes cooling water at atmospheric pressure and 15 °C to 42 atm in order to produce high pressure steam in the heat exchanger. The pump system operates at 77% adiabatic efficiency, which causes a temperature rise of 0.34 °C. The mass flow rate into this pump system is 863,000 kg/hr. In order to achieve this process change, the system of pumps will operate with four in series and two sets in parallel for a total of 8 pumps. Each of these pumps will have a head of 106 m and a work of 162.25 kW.

REF_2The second major pump pressurizes water to provide cooling water to the jacket of the FTR. This takes cooling water at 15.6°C and 1 atm to 20 atm. This pressurized water mixes with the separated water from the products of the syngas unit and the FTR. The flow rate through this pumping system is 1,442,000 kg/hr, which requires more pumps. The system contains a total of 6 pumps, two sets in series to reduce the head and three sets in parallel to reduce the required power. The pumps were assumed to operate with 77% efficiency, with a head of 100 m each, and a required power of 170 kW each. The total head for the system is 200 m, with a total power requirement of 1020 kW.

REF_3This pump pressurizes water to produce low pressure stream. The inlet flow rate is 202,000 kg/hr at 15.6°C and 1 atm. The pump system raises the pressure to 2.4 atm. The relatively low pressure rise in the pump means that a single pump can be used to achieve this pressure rise, with a head of 14 m, a required power of 10 kW, and an adiabatic efficiency of 77%.

HL_1This pump pressurizes the output of a series of flash tanks to enter the hydro-isomerization unit. The flow rate through this pump is 186,000 kg/hr at 1.4 atm and 192°C of rather heavy hydrocarbon chains, giving the stream a volumetric flow rate of 233 m3/hr. The overall head for the pump system is 389 m, with a total power requirement of 256 kW. In order to reduce the head of the pump, two pumps will operate in series, with individual heads of 194.5 m.

41

P1The final significant pumping system pressurizes water to enter a fired heater to produce high pressure steam, which requires an outlet pressure of 42 atm. The water enters the pumping system at 15.6°C and atmospheric pressure, with a mass flow rate of 747,000 kg/hr. The total head for the pumping system is 422.5 m, and the pumping system requires 1,115 kW to operate. In order to achieve this process change, three pumps will operate in series and two sets will operate in parallel for a total of six pumps, each with an individual head of 141 m and an individual power requirement of 186 kW.

CompressorsCompressors are placed throughout the process design to achieve pressure changes in gas, mainly existing in refrigeration loops and to repressurize separated gas. All of the compressors are powered by condensing high-pressure steam generated throughout the process.

H2O_1This is the first of two compressors in a refrigeration loop used to pressurize R134A. This compressor has an inlet feed of 962,000 kg/hr of refrigerant 134A at 2 atm and -8.7 °C. The compressor pressurizes the refrigerant to a pressure of 3.6 atm with an associated temperature of 13.7 °C. This compressor has a head of 1,619 m and a required power of 4,243 kW. Since the power requirement is significant, two compressors will be operated in parallel, each with a required power of 2,121.5 kW. Each compressor is assumed to operate with an adiabatic efficiency of 77%. The output of this compressor system is the inlet feed to H2O_2.

H2O_2This compressor system has two compressors in parallel and takes the output of H2O_1 as its input. The gas has a mass flow rate of 962,000 kg/hr of refrigerant at 3.6 atm and 13.7 °C. Each compressor in parallel has a head of 1,700 m, with a required power output of 2,228 kW and an adiabatic efficiency of 77%. The total required power for this compressor system is 4,456 kW. The outlet of these compressors is gaseous refrigerant at 6.5 atm and 37.4 °C. This gas is passed through a condenser before it is partially vaporized and used to chill the products of the syngas unit.

3PH_1This compressor also exists in a refrigerant loop, but it processes refrigerant 134A at a mass flow rate of 488,000 kg/hr at -24.9 °C and 1 atm. The compressor pressurizes the propane to 6.9 atm. In order to accomplish this, three compressors will be staged in parallel, each with a head of 5,400 m and a required power of 2,400 kW for a total required power of 7,200 kW. The gas exits these compressors at a temperature of 48.6 °C. The compressors operate with an adiabatic efficiency of 77%.

LIGHTS_1This compressor is one of three in series that repressurizes the light components that are separated from the final naphtha stream. These light components have a mass flow rate of 36,300 kg/hr at 1 atm and 31.2 °C. The process stream is mainly made up of CO2, propane, and butane. The compressor has a head of 8,200 m and a required power of 812 kW. This pressurizes the gas to 3 atm and a temperature of 82.2 °C with an adiabatic efficiency of 77%. The pressurized gas is then cooled down and some of the liquefied water is removed before it enters the Lights_2 compressor.

42

LIGHTS_2This takes the pressurized gas from Lights_1 at a temperature of 43.3 °C and 3 atm with a mass flow rate of 33,600 kg/hr. The compressor raises the pressure of this gas to 9 atm and a temperature of 97.7 °C. The compressor has an overall adiabatic efficiency of 77%. The compressor has a head of 8,500 m and a required power of 777 kW. The gas must then be cooled again with any liquid removed and then sent to Lights_3.

LIGHTS_3This compressor takes the pressurized gas from Lights_2 at a temperature of 43.3 °C and a pressure of 9 atm. It raises the pressure to 27 atm with a required power of 578 kW at an adiabatic efficiency of 77% and a head of 9,210 m. The resulting gas has a temperature of 110.5 °C. It mixes with the light components separated from the products of the FTR.

PROP_C1This compressor exists in a recycle loop processing gaseous propane. This propane was vaporized when condensing the light components to enter the light distillation column. The propane enters the compressor at 1 atm and -40.9 °C at a mass flow rate of 83,200 kg/hr. The compressor has a head of 6,850 m and a required work of 1,553 kW. Operating at 77% adiabatic efficiency, the output gas reaches a temperature of 7.1 °C and a pressure of 3.15 atm. The output of the compressor then enters the feed of Prop_C2.

PROP_C2This compressor takes the output of Prop_C1 and pressurizes it to a pressure of 9.9 atm and a temperature of 59.5 °C with an adiabatic efficiency of 77%. The compressor has a head of 7,800 m and a power of 1,772 kW. The output of this compressor is then condensed before being partially vaporized to start the cycle again.

Heat ExchangersThe proposed design requires temperature control mainly because reactors and separation equipment are very temperature-sensitive. Appropriate temperatures as well as pressures allow desired reaction conversion or separation degree. Heat exchangers exist in the design to achieve such desired temperature change. Each heat exchanger’s specific role, operating conditions and size are discussed below. The size of a heat exchanger is specified by its heat-exchanging area. All the heat exchangers are modeled as shell and tube type unless specified otherwise.

FTR_HPSTMThis heat exchanger serves twofold purpose: cooling down hot process stream out of syngas unit and to generate steam which will be used as utility streams later in the process. The flowrate of process stream is 90,399 kmol/hr at 28 atm and the stream is cooled down from 1033 °C to 266 °C. The cold stream is water, of which the flowrate is 47,916 kmol/hr at 42 atm. In the heat exchanging process, the cold water is boiled up to 254 °C and becomes HP steam. The outlet temperature of hot stream is was determined so that just enough energy can be exchanged between the cold and hot stream in order to heat up the water to its dew point. The size of this heat exchanger is 3,790 m2.

E1This heat exchanger further cools down the syngas stream directly out of FTR_HPSTM down to near ambient temperature (43 °C) using cooling water at atm. The inlet and

43

outlet temperature for cooling water are 15.56 °C and 37.78 °C, respectively. The outlet temperature of the hot stream was determined so that the water discharge temperature is specifically 37.78 °C. This specification for the temperatures and pressure of cooling water stays the same throughout unless specified otherwise. The flowrate of cooling water is 809,541 kmol/hr. The size of this heat exchanger is 37,000 m2.

R134A_1_VAPThis heat exchanger further cools down the syngas stream out of E1 down to 1 °C in order to separate out decant water in the process stream. The cold stream is refrigerant R134A at its saturation point at 2 atm (-9.71 °C), flowrate of which is 9,427 kmol/hr. The necessary heat exchange comes from vaporizing the refrigerant to its dew point. The size of this heat exchanger is 2,812 m2.

R134A_1_CONDThis heat exchanger is part of a R134A recycle loop. R134A vaporized at R134A_1_VAP is compressed via a series of compressors and condensed back to saturated liquid state through this heat exchanger. Gaseous R134A at 28.22 atm and 43 °C enters this heat exchanger and leaves at 24 °C as saturated liquid. Cooling water of 450,133 kmol/hr is used to achieve desired cooling of refrigerant. The size for this heat exchanger is 4,890 m2.

FTR_PREHEATThis heat exchanger heats up the dry syngas stream (73,244 kmol/hr) out of R134A_1_VAP from 1 °C to 226 °C. Desired heating is achieved by HP steam produced at FTR_HPSTM. Superheated HP steam is condensed so that the outlet temperature of process stream is 226 °C. The size of this heat exchanger is 4,090 m2.

E2This heat exchanger cools down the stream directly out of FTR from 226 °C prior to entering a 3-phase flash tank separating lighter gas, heavy oil and decant water. Cooling water of 167,046 kmol/hr is used to ensure that the outlet temperature of process stream is 200 °C. The size of this heat exchanger is 3,492 m2.

FTR_LPSTMThis heat exchanger serves similar purpose as FTR_HPSTM. It cools down the light gas stream from FTR products, and at the same time, produces LP steam. The process stream, flowrate of which is 24,263 kmol/hr, at 27 atm and 198 °C, is cooled down to 132 °C. The cold stream is water, flowrate of which is 11,206 kmol/hr at 2.36 atm, is heated up to 127 °C becoming superheated steam. The outlet temperature of process stream was determined so that the cold water can be boiled to its dew point. The size of this heat exchanger is 3,313 m2.

E3This heat exchanger further cools the process stream directly out of FTR_LPSTM down to 43 °C. Cooling water of 119,271 kmol/hr is used to provide necessary cooling. The size of this heat exchanger is 6,262 m2.

R134A_2_VAPThis heat exchanger further cools the process stream directly out of E3 down to 0 °C. R134A at its saturation point at 1 atm (-26 °C), flowrate of which is 4,787 kmol/hr, is vaporized to its dew point provide cooling duty. The size of this heat exchanger is 490 m2.

44

R134A_2_CONDThis heat exchanger is part of a R134A recycle loop. R134A vaporized at R134A_2_VAP is compressed via a compressor and condensed back to saturated liquid state through this heat exchanger. R134A at 6.97 atm and 49 °C enters this heat exchanger and leaves at 27 °C as a saturated liquid at bubble point. Cooling water of 235,585 kmol/hr is used to achieve desired cooling of refrigerant. The size for this heat exchanger is 1,247 m2.

AFTERCOOL1This heat exchanger is part of a series of compressors and heat exchangers to further remove decant water and heavy oils from the light gas out of FTR unit. Process stream, flowrate of which is 774 kmol/hr at 82 °C and 3 atm, enters this heat exchanger, is cooled down to 43 °C by cooling water. 2282 kmol/hr of cooling water is used to provide necessary cooling. The size of this heat exchanger is 241 m2.

AFTERCOOL2This heat exchanger is also part of a series of compressors and heat exchangers to further remove decant water and heavy oils from the light gas out of FTR unit. Process stream, flowrate of which is 730 kmol/hr at 98 °C and 9 atm, enters this heat exchanger, is cooled down to 43 °C by cooling water. 4,334 kmol/hr of cooling water is used to provide necessary cooling. The size of this heat exchanger is 296 m2.

PROP_VAPThis heat exchanger further cools dry light gas, flowrate of which is 10,654 kmol/hr at 10 °C and at 27 atm down to -30 °C. Propane at its saturation point at 1 atm (-42 °C), flowrate of which is 1,887 kmol/hr, is vaporized to its dew point in the heat exchanger to provide cooling duty. The size of this heat exchanger is 540 m2.

PROP_CONDThis heat exchanger is part of a propane recycle loop. Propane vaporized at PROP_VAP is compressed via a series of compressors and condensed back to saturated liquid state through this heat exchanger. Propane at 9.9 atm and 52 °C enters this heat exchanger and leaves at 27 °C as saturated liquid at bubble point. Cooling water of 79,786 kmol/hr is used to achieve desired cooling of refrigerant. The size for this heat exchanger is 840 m2.

E4This heat exchanger heats up the tail gas separated right before the light gas out of PROP_VAP enters a distillation column, DC_LIGHT. Tail gas, flowrate of which is 10,367 kmol/hr at 27 atm and -30 °C, is heated up to 10 °C by liquid water at 1 atm and 43 °C. The water, flowrate of which is 7,644 kmol/hr, leaves the heat exchanger at 15.55 °C. The amount of water needed was determined so that the outlet temperature of process stream is 10 °C. The size of this heat exchanger is 370 m2.

LPG_COOLThis heat exchanger condenses the gaseous LPG product out of a distillation column, DC_LIGHT, to liquid state. The process stream, having the flowrate of 115 kmol/hr at 12 atm and 109 °C, is cooled down to 43 °C. Cooling water, flowrate of which is 1,851 kmol/hr is used to provide necessary heat exchange. The size of this heat exchanger is 360 m2

NAPHTHA_COOLThis heat exchanger cools down the naphtha product out of a distillation column, DC_HEAVY. Naphtha product stream, flowrate of which is 1,460 kmol/hr at 129°C and

45

1.75 atm, is cooled down to 43 °C. Cooling water of 28,092 kmol/hr provides the necessary cooling.

DIESEL_COOLThis heat exchanger cools down the diesel product out of a distillation column, DC_HEAVY. Naphtha product stream, flowrate of which is 814 kmol/hr at 277°C and 2 atm, is cooled down to 43 °C. Cooling water of 62,975 kmol/hr provides the necessary cooling.

DC_L_CThis heat exchanger is a partial condenser for DC_LIGHT. It partially condenses the top tray’s content, mostly tail gas, to achieve the desired top product specifications and reflux ratio. R134A, flowrate of which is 424 kmol/hr at its saturation point at 1 atm (-26 °C) is used. The size of this condenser is 146 m2.

DC_H_CThis heat exchanger is a partial condenser for DC_HEAVY. It partially condenses the top tray’s content, mostly naphtha, to achieve the desired top product specifications and reflux ratio. Cooling water of 29,155 kmol/hr is used to provide necessary cooling. The size of this condenser is 238 m2.

DC_H_RThis heat exchanger is a reboiler for DC_HEAVY. It heats up the bottom tray’s content, mostly diesel, to achieve the desired bottoms product specifications and reflux ratio. Syngas stream, flowrate of which is 90,399 kmol/hr at 1066 °C and 28 atm, directly out of syngas unit is used to provide necessary heating duty for this reboiler. The syngas stream is cooled down to 1034 °C along the process. The type of this heat exchanger is kettle reboiler, not shell and tube. The size is 150 m2.

DC_L_RThis heat exchanger is a reboiler for DC_LIGHT. It heats up the bottom tray’s content, mostly LPG, to achieve the desired bottoms product specifications and reflux ratio. Syngas stream directly out of DC_H_R is used to provide necessary heating duty for this reboiler. The syngas stream is cooled down to 1033 °C along the process. This heat exchanger is also kettle reboiler, not shell and tube, having the size of 25 m2.

Flash Tanks and Three-Phase SeparatorsFlash tanks are used to separate liquid phases from gaseous phases. Three-phase separators are used to separate two immiscible liquids – such as water and hydrocarbons – and a gas. All of the flash tanks and three-phase separators need level controls to ensure that a liquid phase exists in the flash tank and to prevent the vapor from exiting the vessel through the liquid exit. This consists of a liquid level measurement connected to a flow controller. For a three-phase separator, the hydrocarbons and water are separated due to density differences, so level controllers must ensure that the proper levels exist in order to achieve this separation.

H2O_RMVLThis is a flash tank that removes the liquid water from the syngas. This flash tank processes a total mass flow rate of 1,239,000 kg/hr of material, but only 30,900 kg/hr are liquid. This liquid is mainly water but with some dissolved components. The flash tank was sized to have a volume of 25.8 m3. The flash tank operates adiabatically and isochorically at 1 °C and 28 atm. The liquid leaves in a volumetric flow rate of 309.5 m3/hr, and the vapor leaves at a volumetric flow rate of 59,000 m3/hr.

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3P_1This is a three-phase separator after the FTR that is used to separate the water produced in the Fischer-Tropsch reaction from the hydrocarbons while also separating the lighter, gaseous molecules from the longer chain hydrocarbons existing as a liquid. This separator operates at a pressure of 27.1 atm has a 1 atm pressure drop into it in order to vaporize the some of the lighter components. It operates adiabatically at 198.5 °C. It processes a total mass flow rate of 930,000 kg/hr, with a water flow rate of 149 m3/hr, a light hydrocarbon vapor flow rate of 32,800 m3/hr, and a heavy hydrocarbon flow rate of 243 m3/hr. The volume of this three-phase separator was determined to be 295 m3/hr in order to achieve sufficient separation between the two liquid phases and the gaseous phase.

L_H_SPLITThis flash tank splits the heavy stream out of 3P_1 into a naphtha-rich stream that by-passes the hydro-isomerization unit and a heavy stream that enters the hydro-isomerization unit. The flash tank has a large pressure drop to prevent dissolved carbon dioxide from entering the HI unit. The flash tank is maintained at a pressure of 1.35 atm. The entering fluid experiences a pressure drop of 25.75 atm when entering the tank. The tank operates adiabatically at a temperature of 192 °C. A total mass flow rate of 193,000 kg/hr, with a vapor flow rate of 4,332 m3/hr and a liquid flow rate of 233 m3/hr.

3P_2This is a three-phase separator processing the light ends out of 3P_1 after they have been cooled. The water and heavier components that were present liquefied and settle to the bottom of the tank where the two immiscible phases are separated by density differences. Water exits out of the bottom of the tank and the hydrocarbons leave out of the middle. This tank processes 588,000 kg/hr, with a water flow rate of 240 m3/hr, a liquid hydrocarbon flow rate of 120 m3/hr, and a gaseous flow rate of 7,790 m3/hr. The tank operates at a pressure of 27.1 atm without a pressure drop entering the tank and adiabatically at 0 °C.

NAPH_SEPThis three-phase separator combines the naphtha-rich streams with the light ends from the heavy distillation column. There is still some water in the streams so this must be separated from the hydrocarbons. The tank operates at a pressure of 1 atm and a temperature of 31.2 °C. It processes a mass flow rate of 122,000 kg/hr, with a water outflow of 0.88 m3/hr, a liquid hydrocarbon outflow of 124 m3/hr, and a vapor hydrocarbon outflow of 19,100 m3/hr.

F1This three-phase exists to remove any condensed vapor after compression and cooling. It only has a small amount of liquid hold-up due to the small liquid flow rates. The flash tank operates at 3 atm and 43.3 °C, with a total mass flow rate of 36,300 kg/hr. The water flow rate out of the tank is 245 kg/hr and the liquid hydrocarbon flow rate is 2,470 kg/hr. The vapor hydrocarbon flow rate is 33,600 kg/hr of light hydrocarbons.

F2This is similar to F1, separating the condensed vapor from the light hydrocarbons that will be sent to the light distillation column. The three-phase separator operates at 9 atm and 43.3 °C with a total mass flow rate of 33,600 kg/hr. The water flow rate out of the tank is 280 kg/hr and the liquid hydrocarbon flow rate is 10,300 kg/hr due to the higher pressure. The vapor exiting the three-phase separator is the very light hydrocarbons with a mass flow rate of 23,000 kg/hr.

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LIGHT_SEPThis flash tank exists to separate out non-condensables such as carbon monoxide and hydrogen from the tail gas and LPG components that are sent to the light distillation column. Since the flow rate of non-condensables are much greater than that of the condensables, a significant amount of the light hydrocarbons exist in the vapor phase. The flash tank operates adiabatically and isochorically at 27 atm and -30 °C. The flow rate into the flash tank is 287,000 kg/hr, with a flow rate of vapor at 273,000 kg/hr and a flow rate of liquid of 14,000 kg/hr. The vapor is sent to a furnace to produce high-pressure steam and the liquid is sent to the lights distillation column.

LPG_SEPThis flash tank exists to separate naphtha from LPG that exited the light distillation column. The LPG entering the flash tank was not pure enough, and a higher recovery of naphtha is beneficial economically. The flash tank operates at 1 atmosphere with a pressure drop into the flash tank of 22.7 atm. The tank operates adiabatically at 21.4 °C. With an inlet flow rate of 8,850 kg/hr, it produces a vapor flow rate of 6,890 kg/hr and a liquid flow rate of 1,950 kg/hr. The vapor is then compressed and condensed to produce the LPG product while the liquid is sent to be further processed with other naphtha-rich streams.

Fired HeatersIn this process, there are two fired heaters – the syngas preheater and a high-pressure steam boiler. These fired heaters have very high temperatures in the radiative section. The flue gas then enters the conductive section and transfers more heat to the process fluid through convection. In designing these fired heaters, 85% efficiency was assumed, meaning that only 85% of the energy contained in the fuel gas is transferred to the process streams. Additionally, it was assumed that 60% of the energy was transferred in the radiative section, with the remaining 40% transferred in the convective section. All furnaces need controls to prevent pressure from building up due to hot spots in the furnace. Additionally, the flow rate of the fuel must be controlled to prevent the furnace from reaching temperatures that are outside the range of operability.

SU_PREHEATThis fired heat heats the syngas feed from 168 °C to 538 °C at a constant pressure of 28 atm. This preheat furnace uses the tail gas that is separated from the light ends of the light distillation column and from the flash tank immediately before that column. The syngas preheater needs to process a feed rate of 760,000 kg/hr of methane, steam, and carbon dioxide. The flue gas from the preheater is collected and cooled before being recycled to the feed as the carbon dioxide feed. The overall heat duty of this furnace is 290,000 kW. In order to supply the necessary heat, a flow rate of 73,650 kg/hr of tail gas must be fed to the burners. The size of this furnace can be determined by the heat exchange area in the radiative and convective sections. The radiative section has a heat transfer area of 5,800 m2 and the convective section has a heat transfer are of 10,680 m2.

HP_STM_BOILThis furnace uses the excess tail gas to generate high-pressure steam. The heat duty of this furnace is 666,200 kW, which requires a flow rate of 199,000 kg/hr of tail gas. The furnace processes 746,500 kg/hr of water at 15.9 °C and 41.8 atm. The steam leaves at

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254.4 °C as high-pressure steam to power compressors or to be sold. The size of furnace can be determined in a similar fashion to the preheat furnace, giving areas in the radiative and convective sections of 13,324 m2 and 24,540 m2 respectively.

Distillation ColumnsMost of the separations necessary in this design are able to be achieved by flash tanks. However, when the mixture in need of separation cannot be split in terms of physical means – phase and/or density difference, using distillation column can provided necessary separation power, taking advantage of difference in vapor pressures of the components.

DC_LIGHTThis distillation column separates the dry stream of light gas into tail gas and LPG. The liquid feed of light gas at -30 °C and 27 atm, with flowrate of 279.5 kmol/hr, is adiabatically flashed into the distillation column and liquid overflow flows downward and vapor overflow rises to top. Vapor overflow at the top tray then flows into a partial condenser where it is partially condensed. Condensed liquid is refluxed into the column while remaining vapor comes out of the system. Liquid overflow at the very bottom tray flows into a reboiler where some of it is boiled and refluxed into the column while the rest exits the column as bottoms product. The duties of the condenser and the reboiler are determined to satisfy the following specifications: 5 mol. % of propane in the feed exits through top and the composition of hydrocarbons heavier than propane in the bottoms product is 99 mol. %. As a result, 138 kmol/hr of gaseous tail gas at -21 °C and 23 atm exits through the top and 141.5 kmol of liquid LPG at 123.17 °C and 23.68 atm exits through the bottom. Each product then goes through a series of minor separation system and/or heat exchanging step to be completely taken out of the system as final products. To achieve desired separation, the column has total of 15 equilibrium stages with 0.61 m tray spacing. To account for the number of trays and total amount of overflow rate in the column, the column needs to be 12.19 m tall and 2.18 m in diameter. It is operating at 23 atm with total pressure drop of 0.68 atm. The temperatures at the top and bottom are -18.4 °C and 98.7 °C, respectively. Details on the condenser and reboiler can be found in the Heat Exchangers section above.

DC_HEAVYThis distillation column separates the heavy oil stream out of HI unit into naphtha and diesel. The feed, containing both vapor and liquid phase, at 123 °C and 2 atm, with flowrate of 2274 kmol/hr, enters the distillation column and liquid overflow flows downward and vapor overflow rises to top. Vapor overflow at the top tray then flows into a partial condenser where it is partially condensed. Condensed liquid is refluxed into the column while remaining vapor comes out of the system. Liquid overflow at the very bottom tray flows into a reboiler where some of it is boiled and refluxed into the column while the rest exits the column as bottoms product. The duties of the condenser and the reboiler are determined to satisfy the following specifications: top product stream contains 98 mol. % of hydrocarbons lighter than decane and bottom product stream contains 98 mol. % of hydrocarbons heavier than undecane. As a result, 1459.5 kmol/hr of gaseous naphtha at 128.9 °C and 1.75 atm exits through the top and 814.5 kmol of liquid diesel at 276.7 °C and 2 atm exits through the bottom. Each product then goes through a series of minor separation system and/or heat exchanging step to be completely taken out of the system as final products. To achieve

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desired separation, the column has total of 25 equilibrium stages with 0.61 m tray spacing. To account for the number of trays and total amount of overflow rate in the column, the column needs to be 18.29 m tall and 5.36 m in diameter. It is operating at 1.75 atm with total pressure drop of 0.25 atm. The temperatures at the top and bottom are 141.2 °C and 253.8 °C, respectively. Details on the condenser and reboiler can be found in the Heat Exchangers section above.

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Heat IntegrationThe proposed GTL process requires significant heat transfer to be operational. Reactors operate in high temperatures and the products for those reactors are well above ambient temperature. Because the final products should be in room temperature (about 25 °C), the streams need to be eventually cooled down. Also, for separation purposes, some of the streams need to be chilled well below freezing point. In other words, the process requires large amount of cooling and heating duties, and without such heat transfer system, proposed design remains incomplete. This as a whole poses one engineering problem: heat integration. Of course, the process is viable and feasible without studying of heat integration. Nevertheless, well-heat-integrated plant operating with optimized heat transfer system minimizes heat loss, and therefore, saves significant amount of utility costs.

Looking at the overall process, the streams going into the syngas are heated to about 750 °C. Then the product out of the syngas unit is chilled to room temperature for water removal and then heated up to about 220 °C before it goes into the FTR unit. The product out of the FTR unit then separates into two streams: light gas and heavy oil. Light gas stream goes through a series of compressors and chillers to remove non-hydrocarbon gases and water. The remaining hydrocarbon gas is then fed to a distillation column where tail gas is separated from LPG. Heavy oil is heated up to about 200 °C upon entering HI unit. The product of HI unit is fed to a distillation column where naphtha and diesel products are separated out. Required heating and cooling duties for this heat transfer process is illustrated on the Figure 9 below.

FIGURE 9- HEAT INTEGRATION CURVE BEFORE ADJUSTMENT

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As shown above, the process requires far more cooling duties (about 1,480,000 kW) than heating duties (about 440,000 kW), which indicates minimum cooling duties of about 1,000,000 kW. It also indicates that the process possesses excess amount of heat. In original design of the plant, the excess heat is to be removed by cooling water.

Studying the proposed GTL process’s heat integration, it was concluded that using the excess heat, removal of which introduces significant amount of cost for cooling water, to produce HP steam and provide heat for some cold streams builds better heat-integrated plant. Significant amount of HP steam is needed for various parts for the process, and HP steam can be sold at a decent price. Also, some of the cold streams can be heat up using the hot streams.

As a result, the hot streams, each out of syngas unit and FTR, are used to produce HP steam. Also, the hot stream out of syngas unit provides necessary heating duties for the two reboilers attached at the distillation column. The plant’s heat integration after such adjustment is shown on Figure 10 below.

FIGURE 10- HEAT INTEGRATION CURVE AFTER ADJUSTMENT

Now, the minimum cooling duties are reduced to about 800,000 kW. Assuming the excess heat is removed by cooling water, the adjusted design saves about $894,450 every year. For equipment, two steam generators, each after syngas unit and FTR, are to be added in the process design. However, by doing so, four existing heat exchangers can be significantly reduced in size. As a result of such plus and minus on equipment cost, it was predicted that this adjustment on design actually saves $2,000,000 on capital cost.

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In short, studying of heat integration of the proposed GTL process resulted in saving $2,000,000 on capital cost and additional $894,450 annually on cooling water. Additionally, significant amount of tail gas is separated out from the distillation column. That tail gas can be used as fuel gas, burning of which provides heat for syngas preheater. In fact, the process is producing more tail gas than it needs for its syngas preheater. The rest is burnt to produce additional HP steam.

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SafetyOperation of the proposed GTL process includes significant safety hazards. Significant parts of the process operate at high-temperature and high-pressure. Also, various chemicals can be very dangerous.

As an oil-processing plant, the proposed GTL plant mainly processes hydrocarbons with various chain-lengths. Toxicity of such hydrocarbons is very low, if not none. However, flash points of lighter hydrocarbons can go lower than room temperature, therefore must be handled with extreme care. Reactors go up very high in temperature and pressure; therefore there is always a chance for explosion or fire, given that the process streams are very flammable.

Included safety equipment is mostly for maintenance purpose. High Pressure Coolant Injection System (HPCI) and vacuum purging system are necessary to clean up the reactors prior to maintenance or immediately clear out any toxic contaminants in case of leakage. HPCI is necessary because of the high pressure in the reactor. If unpredicted error occurs and the reactors need to be shut down immediately, enough coolant should be pumped with enough pressure into the high pressure reactor to flush out and cool down the reactor. Vacuum purging system purges the reactor with inert gas to ensure the safety of workers.

Significant portion of process streams is gaseous, and some of them can be hazardous. Carbon monoxide and hydrogen gas are two main hazards. Leakage of those gases can result in detrimental situations. Therefore, gas vent systems must be installed with care such that failure of a part will not compromise the whole system.

In addition, ignition sources must be kept at minimum to prevent fire and explosion hazards. Smoking should be prohibited within significant distance from the plant as well as matches, lighters and other ignition sources must be regulated and handled with care, if necessary.

The plant will be located in Deposit, NY. The location is relatively free from natural disaster such as earthquakes or typhoons. However, some equipment will be installed to prevent loss due to unexpected events such as power outage: diesel generators and flywheels. Because the process produces diesel and the plant will have significant amount of diesel stored near, having diesel generators is one of the best way to eliminate the loss caused by power outage. Also, flywheels store significant amount of rotational energy therefore can be used to operate pumps or compressors for a brief amount of time in case of unexpected power loss situation.

A detailed information on the chemicals processed in the plant is available in the table below.

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TABLE 6 - SUMMARY OF HAZARDOUS CHEMICALS [13]

NFPA number Flammability Limit (%)

Flash Point Temperature (°C)Blu

eRed

Yellow

White

CO 4 4 2 12 to 75 n/aCO2 2 0 0 n/a n/aH2 3 4 0 4 to 75 n/aH2O 0 0 0 n/a n/aN2 3 0 0 n/a n/aO2 3 0 0 OX n/a n/aR134A

1 0 1 n/a n/a

C1 1 4 0 5 to 15 -188.15C2 1 4 0 3 to 12.5 -135.15C3 1 4 0 2.1 to 9.5 -104C4 1 4 0 1.6 to 8.5 -60.15C5 1 4 0 1.4 to 8 -57.15C6 2 4 0 1.1 to 7.5 -22C7 1 3 0 1.1 to 6.7 -4.15C8 1 3 0 1 to 6.5 13.33C9 1 3 0 0.8 to 2.9 31C10 1 2 0 0.8 to 2.6 46.11C11 1 2 0 0.7 n/aC12 1 2 0 0.6 n/aC13 1 1 0 n/a n/aC14 1 1 0 n/a n/aC15 1 1 0 n/a n/a

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EnvironmentalEPA regulates emission of harmful chemicals, especially those affiliated with ozone, global warming, heavy metals and other air pollutants. Other than water and nitrogen gas, current design of GTL plant emits no other chemicals out of the system. Carbon dioxide gas is completely recycled inside of the plant and all the organic compound stays in the system and removed from the system as products. Also, the purity of wastewater is above 99 wt. %. In general, proposed design of GTL plant poses no direct and eminent threat to the environment.

Nevertheless, the purity of wastewater cannot be always guaranteed as it is separated from the heavy oils due to immiscibility. Slight error in temperature or pressure control can affect the purity of decant water in a negative way. Also, the purity of water used to flush the reactors for maintenance purpose or emergency shutdown of reactors or equipment can be too low to discharge without treatment. Wastewater resulting from such flush activity can contain large concentrations of liquid oil which can be detrimental to the environment. Therefore, wastewater treatment facilities should be included in the detailed plant design.

Environmental concerns could rise from the use of refrigerants. Especially, prolonged emission of significant amount of R134A should be avoided as R134A has relatively high global warming rating. However, such environmental concerns due to refrigerants are minimal because the amount of refrigerants present in the system is relatively low and refrigerants are trapped in recycle loops. Also, both R134A and propane pose no threat to ozone layer. Care should be exercised with handling and waste-treatment of refrigerants as well as maintenance of recycle loops.

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Utility SummaryFor steady-state operation for designed GTL process, utility steams are necessary for the following purposes: heating/cooling of streams or equipment, powering equipment and providing necessary components for the reactions.

Temperature control – heating or cooling of process streams – is achieved by heat exchangers including reboilers and condensers attached to the distillation columns. Three types of utility streams were used for this purpose: cooling water, steams and refrigerants. Cooling water was used to cool down or heat up streams to the near ambient temperature. It was assumed to be provided from the river near the plant with the feed temperature of 15.56 °C. The state law of New York regulates that the water discharged from the plant must be no hotter than 70 °F (21.11 °C). However, since the discharged water will be collected into a pond where the water will stay and cool down to ambient temperature, the discharge water temperature was assumed to be 37.78 °C.

Steams were used to heat up the streams above ambient temperature where cooling water cannot provide necessary heat transfer. Three types of superheated steams were provided: HP (41 atm and 255 °C), MP (8.5 atm and 178 °C) and LP steam (1.4 atm and 127 °C). Only HP steams were deemed appropriate for the designed process. It was assumed that, at the heat exchangers with steams, the inlet superheated steam is condensed to its bubble point so that the heat released from steam condensation is used to heat up the process stream.

In case where cooling down below freezing point was necessary, two types of refrigerants were used: R134A and propane. Appropriate refrigerants were chosen so that the outlet temperature of process stream is about 5 to 10 °C above the saturation temperature of the refrigerant at 1 atm as the necessary cooling duty comes from vaporization of saturated refrigerant. Saturation temperatures for R134A and propane at 1 atm are -26.11 °C and 43.81 °C, respectively.

Pumps and compressors require power source to operate. For the current design, all pumps are powered by electricity and all compressors are powered by HP steam (steam turbine driver).

For syngas unit and hydro-isomerization unit, some utility streams are necessary to operate and carry on the reaction with desired conversion and rate. For syngas unit, HP steam, carbon dioxide and oxygen streams are necessary; and for HI unit, fuel gas, hydrogen gas, LP steam, electricity and cooling water are needed. Table below shows utility type, its unit cost, usage rate and annual cost for each.

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TABLE 7- ANNUAL UTILITY USAGE

Utility Unit Cost Annual Usage Annual Cost (in $1,000s)

Water $0.13/kL 8,688,997 kL 1,148Refrigeration n/a n/a 136

Electricity $0.04/kW-h 53,463,459 kW-h 2,139

Fuel Gas $2.84/GJ 1,083,465 GJ 3,081

H2 $0.13/kg 11,167,361 kg 1,477

LP STM $7.72/Mg 58,230 MG 449

Total 8,430

The designed GTL process produces more HP steam than it requires. Therefore, there is no need to purchase HP steam from outside vendors. Also, refrigerants are almost completely recycled as they are trapped in a refrigerant loops. However, it was assumed that there is 5% annual loss of refrigerants that needs to be refilled. The annual cost of $136,000 for refrigerant takes care of that 5% annual loss.

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Operating Cost SummaryThe operating cost can be estimated from analyzing the individual components of both fixed and variable costs. In more detail, fixed costs refer to costs that are not dependent on the actual manufacturing of the product, whereas variable costs refer to costs that are dependent on production. The fixed and variable costs can be separated into its following components as follows.

Fixed CostsThe total fixed cost can be separated into three major components: fixed charges, plant-overhead costs, and general expenses. Fixed charges refer to costs related to depreciation, insurance, and rent. These costs are generally constant from year to year and are unrelated to the manufacturing of the product. Plant-overhead costs refer to costs related to the general functioning of the plant. These services include providing hospital and medical services; payroll overhead, which includes pensions, social security, and vacations; property protection, and other services. Lastly, general expenses refer to administrative costs, distribution and selling costs, research and development, and financing. The exact costs for each component are provided in Table8 below.

TABLE 8- FIXED COSTS

Fixed ChargesDepreciation $264,242,00

0Insurance $18,718,000Rent $0

Fixed Charges Subtotal $282,960,000

Plant-overhead Costs

Subtotal $21,910,000

General ExpensesAdministrative Costs $5,056,000Distribution and Selling Costs

$33,600,000

Research and Development $32,766,000Financing $79,272,000

General Expenses

Subtotal $150,694,000

Total $455,564,000

Variable CostsVariable costs refer to costs directly related to the manufacturing of the products. Thus, raw material costs, operating labor costs, utility costs, maintenance and repair costs, operating supplies costs, and laboratory charges were determined and included in the overall calculation of the variable cost. It is important to note that the cost

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estimates are based on the assumption that the plant will operate 330 days/year. This approximation takes into account the downtime that will occur every four years to change the catalyst and for other expected and unexpected delays and shutdowns. The total raw materials costs required for production is specified in Table 9.

TABLE 9- RAW MATERIALS COST

Raw Materials

Price Consumption (per

day)

Cost ($/day) Operating Time

(days/yr)

Cost ($/year)

Methane $70,621/MSCM

14.16 MSCM

$1,000,000 330 $330,000,000

Oxygen $110.23/ton 11,488 ton $1,266,300 330 $417,874,000

Catalyst $22.05/kg - - 330 $13,961,000

Total 761,835,000

To determine the cost of the catalyst, the total cost for the catalyst that is required every four years was evenly distributed for the four years.

The operating labor cost was also estimated to more accurately understand the variable costs. To determine the total number of workers related to production, general heuristics were referred to determine that with approximately 55,000 barrels of product per day, 125 employees will be required every hour. Furthermore, it was estimated that there will be 2 common workers, workers that are easier to acquire, for every 1 skilled worker, workers that are more technical and experienced. Their hourly rate and expected labor costs are provided below. It is important to note that these labor costs only refer to costs related to manufacturing. Salaries for personnel in administration, research and development, and laboratories are calculated in other cost streams. A detailed breakdown of labor costs are presented in Table 10.

TABLE 10- TOTAL LABOR COSTS

Type of Employee

# of Employees

Pay ($/hr)

Total Labor Cost per day

Total Labor Cost per year

Skilled 42 46 $46,368 $15,301,440Common 83 28 $55,776 $18,406,080

Total $33,707,520

To more accurately estimate the variable costs, the utility costs were also calculated. It is important to note that the refrigeration costs in the table below refer to the refrigeration added every year due to losses. This was determined by taking 5% of the total refrigeration required. Provided is the utility estimate in Table 11.

TABLE 11- UTILITY COSTS

Utilities PriceConsumpti

on (per day)

Cost ($/day)Operating

Time (days/yr)

Cost ($/year)

Electricity $0.04/KW-h 162,050 KW-h

$6,482 330 $2,139,000

Fuel Gas $2.84/GJ 3,283 GJ $9,336 330 $3,081,00

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0Refrigeration $3.30/kg 124 kg $412 330 $136,000

Process Water

$0.13/kL 26,339 kL $3,479 330 $1,148,000

Hydrogen $0.13/kg 33,830 kg $4,475 330 $1,477,000

LP Steam $7/ton 194 tons $1,360 330 $449,000

Total $8,430,000

With the estimated raw materials cost, operating labor cost, and the utility cost, the overall variable cost was determined. It is important to note that maintenance and repairs refers to costs related to keeping the plant working efficiently. These costs include labor, supplies, and supervision. Operating supplies refers to miscellaneous supplies required to keep the plant operating efficiently. These costs include lubricants, testing chemicals, and similar supplies that are not included in raw materials and maintenance and repairs. The laboratory charges refer to costs related to testing the products to ensure quality.

The variable costs and its breakdown are in Table 12.

TABLE 12- VARIABLE COSTS

Variable CostsRaw Material $761,835,000Operating Labor $33,708,000Utilities $8,430,000Maintenance and Repairs

$98,000,000

Operating Supplies $25,000,000Laboratory Charges $3,528,000

Total $930,501,000

With the calculated fixed costs and the variable costs, the total operating cost is $1,386,065,000/year.

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Equipment Information SummaryFor the successful operation of the plant, various types of equipment need to be sized and installed. The costs provided in the tables below are bare module costs that were calculated with a CEPCI value of 724.1 [14]. Detailed calculations of the estimated costs for the equipment are provided in the section “Source of Costs,” which follows the summary of the equipment.

Compressors It is important to note that the maximum capacity of a centrifugal compressor is 3000 kW. Thus, for compressors with fluid power requirements greater than 3000 kW, separate compressors were calculated. For example, for compressor H2O_1, two compressors of size 2121.5 kW were sized and costed individually and summed. The same procedure was applied for all equipment, depending on the maximum capacity.

The minimum capacity of a centrifugal compressor is 450 kW. Thus, a rotary compressor was used for the two LPG compressors. The specifications for the compressors are on Table 14.

DrivesIt is also important to note that the compressor costs do not take into account the drives that are required. Thus, the drives were specified and costed separately and for drives that exceeded the maximum capacity of 7500 kW, the drives were separated accordingly. The drives were designed to operate with steam, since significant steam is generated in the plant process. The specifications of the drives are provided on Table15.

FurnacesIn the design, there are also two nonreactive fire heaters. The first heater is to ensure that the syngas entering the syngas reactor is at the optimal temperature and the second heater is to generate steam. Since the plant design generates more fuel gas than required, the excess will be burned to provide heat to generate more steam. The maximum capacity for a nonreactive fire heater is 100,000 kW. The specifications of the two nonreactive hire heaters are provided on Table 16.

Process VesselsThere are also 9 vertical process vessels, 4 of which are three-phase separators. Since CAPCOST was used to estimate the costs for the vessels, to account for the internals in a three-phase separator, a multiplier factor of 1.2 was used. The maximum capacity of a vertical process vessel is 520 m3. The specifications of the vessels are provided on Table 17.

Heat ExchangersThe most common equipment in the plant design is the heat exchanger. Although most of the heat exchangers are U-tube, there are a few kettle reboilers as well. It is important to note that the maximum capacity of a U-tube heat exchanger is 1,000 m2, whereas the maximum capacity of a kettle reboiler is 100 m2. The specifications for the heat exchangers are provided on Table 18.

62

ReactorsThe three reactors present in this process are all packed-bed reactors, meaning that within a large diameter shell are many tubes containing catalyst. Since there are no direct correlations for the cost of a packed-bed reactor, the costs were estimated by sizing each reactor as a horizontal process vessel and a heat exchanger. The horizontal process vessel estimation will cost the outer shell of the reactor while the heat exchanger part will cost the internals. The costs for the two parts were summed together for each reactor. The total cost of the reactors is given in Table 19.

Towers & TraysThere are two separation columns in this plant design; one column for the separation of the lights into tail gas and LPG and another column for the separation of the heavies into naphtha and diesel. The trays were costed separately from the tower and are summed in the following page. The maximum capacity of a tower is 520 m3 and the maximum area for a sieve tray is 12.3 m2. Tower and tray specifications are provided on Table 20 and Table 21, respectively.

PumpsThere are 5 major pumps that are required for this plant design. The head and shaft power required for the pumps however, are very large. Thus, the pumps were separated in series to reduce the head required and in parallel to reduce the shaft power required. The maximum shaft power available for a centrifugal pump is 300 kW. The specifications for the pumps are provided on Table 22.

Fixed Roof TanksThere will be numerous fixed roof tanks to store the oil products. Despite the higher cost of employing fixed roof tanks instead of floating roof tanks, fixed roof tanks can withstand snowy conditions. Since the plant is located in NY, snow may collapse the floating roof tanks. The maximum capacity for a fixed roof tank is 30,000 m3. The specifications for the tanks are provided on Table 23.

The total bare module costs for all equipment are provided on Table 13.

TABLE 13- TOTAL EQUIPMENT COST

Equipment CostCompressors $25,234,475

Drives $17,408,584Fire Heaters $281,564,908

Process Vessels $20,737,096Heat Exchangers $73,167,486

Reactors $640,600,881Towers $7,364,907Trays $1,729,774Pumps $26,550,409Tanks $75,353,867

63

Total $1,169,712,387

64

TABLE 14- COMPRESSOR SUMMARY

Compressor

Material Type Size (Fluid Power kW)

Process Conditions

Cost # of Compress

ors

Total Cost

H2O_1 Carbon Steel

Centrifugal 2,122 13.7 C @ 3.6 ATM $2,350,184 2 $4,700,368

H2O_2 Carbon Steel


3PH_1 Carbon Steel


LIGHT_1 Carbon Steel

Centrifugal 813 82.1 C @ 3.0 ATM $1,177,790 1 $1,177,790


Centrifugal 777 97.7 C @ 9.0 ATM $1,138,260 1 $1,138,260


Centrifugal 578 110.5 C @ 27 ATM

$903,950 1 $903,950

PROP_C1 Carbon Steel


PROP_C2 Carbon Steel


LPG_1 Carbon Steel

Rotary 127 62.7 C @ 3.4 ATM $392,079 1 $392,079

LPG_2 Carbon Steel

Rotary 133 109.1 C @ 11.9 ATM

$425,993 1 $425,993

Total $25,234,475

TABLE 15- DRIVE SUMMARY

Drive Material Type Size (Shaft Power kW)

Process Conditions

Cost # of Drives

Total Cost

H2O_1 Carbon Steel

Steam Turbine

5,510 13.7 C @ 3.6 ATM $2,146,652 1 $2,146,652

H2O_2 Carbon Steel

Steam Turbine

5,787 37.4 C @ 6.5 ATM $2,157,949 1 $2,157,949

3PH_1 Carbon Steel

Steam Turbine

4,676 48.6 C @ 6.9 ATM $2,103,566 2 $4,207,133

65


Steam Turbine

1,055 82.1 C @ 3.0 ATM $1,477,538 1 $1,477,538


Steam Turbine

1,009 97.7 C @ 9.0 ATM $1,423,866 1 $1,423,866


Steam Turbine

751 110.5 C @ 27 ATM

$1,266,768 1 $1,266,768

PROP_C1

Carbon Steel

Steam Turbine

2,020 7.1 C @ 3.1 ATM $1,776,948 1 $1,776,948

PROP_C2

Carbon Steel

Steam Turbine

2,306 59.5 C @ 10 ATM $1,838,077 1 $1,838,077

LPG_1 Carbon Steel

Steam Turbine

165 62.7 C @ 3.4 ATM $562,726 1 $562,726

LPG_2 Carbon Steel

Steam Turbine

173 109.1 C @ 12 ATM

$580,926 1 $580,926

Total $17,408,584

66

TABLE 16- FIRE HEATER SUMMARY

Fire Heater

Material Type Capacity (Duty kW)

Process Conditions

Cost # of Heaters

Total Cost

SU_Preheat

Stainless Steel

Nonreactive

88,750 168 C @ 28 ATM $17,980,381

8 $143,843,049

Steam Generato

r

Stainless Steel

Nonreactive

95,170 254 C @ 42 ATM $19,674,551

7 $137,721,859

Total $281,564,908

TABLE 17- PROCESS VESSELS SUMMARY

Process Vessel

Material Type Capacity (Volume

m3)

Process Conditions

Cost # of Vessels

Total Cost

3P_1 CS Vertical 295 198 C @ 27 ATM $15,943,555

1 $15,943,555

3P_2 CS Vertical 55 -0.02 C @ 27 ATM

$1,650,828

1 $1,650,828

LIGHT_SEP CS Vertical 8 -30 C @ 27 ATM $302,865 1 $302,865NAPH_SEP CS Vertical 29 31 C @ 1 ATM $615,291 1 $615,291H2O_RMVL CS Vertical 26 1 C @ 28 ATM $1,202,00

41 $1,202,00

4L_H_SPLT CS Vertical 19 192 C @ 1 ATM $386,683 1 $386,683LPG_SEP CS Vertical 3 21 C @ 1 ATM $117,944 1 $117,944

FLASH_COMP1

CS Vertical 15 43 C @ 3 ATM $351,212 1 $351,212

FLASH_COMP2

CS Vertical 4 43 C @ 9 ATM $166,712 1 $166,712

Total $20,737,096

67

TABLE 18- HEAT EXCHANGER SUMMARY

Heat Exchanger

Material Type Process Conditions Capacity (Area m2)

Cost # of Exchange

rsTotal CostInlet

TempOutlet Temp

Pressure

FTR_HPSTM CS-shell/CS-tube

U-Tube 1,032 C 265 C 28 ATM 951 $2,038,310

4 $8,153,240

E1 CS-shell/CS-tube

U-Tube 265 C 43 C 28 ATM 963 $1,001,107

27 $27,029,897

R134_1_VAP CS-shell/CS-tube

U-Tube 43 C 1 C 28 ATM 938 $977,337 4 $3,909,350

R134_1_COND CS-shell/CS-tube

U-Tube 37 C 24 C 6.5 ATM 935 $923,088 6 $5,538,531

FTR_PREHEAT CS-shell/CS-tube

U-Tube 1 C 226 C 28 ATM 754 $759,653 4 $3,038,612

E2 CS-shell/CS-tube

U-Tube 226 C 200 C 28 ATM 885 $926,236 3 $2,778,709

FTR_LPSTM CS-shell/CS-tube

U-Tube 198 C 132 C 27 ATM 859 $899,367 5 $4,496,834

E3 CS-shell/CS-tube

U-Tube 132 C 43 C 27 ATM 981 $1,016,050

6 $6,096,299

R134A_2_VAP CS-shell/CS-tube

U-Tube 43 C 0 C 27 ATM 844 $884,650 1 $884,650

R134A_2_COND CS-shell/CS-tube

U-Tube 48 C 27 C 7 ATM 958 $993,684 2 $1,987,369

AFTERCOOL1 CS-shell/CS-tube

U-Tube 82 C 43 C 3 ATM 229 $298,920 1 $298,920

AFTERCOOL2 CS-shell/CS-tube

U-Tube 98 C 43 C 9 ATM 278 $343,618 1 $343,618

68

PROP_VAP CS-shell/CS-tube

U-Tube 10 C -30 C 27 ATM 450 $518,121 1 $518,121

PROP_COND CS-shell/CS-tube

U-Tube 59 C 27 C 10 ATM 545 $579,119 1 $579,119

E4 CS-shell/CS-tube

U-Tube -30 C 10 C 27 ATM 329 $407,222 1 $407,222

NAPHTHA_COOL

CS-shell/CS-tube

U-Tube 129 C 43 C 1.8 ATM 755 $760,829 3 $2,282,488

DIESEL_COOL CS-shell/CS-tube

U-Tube 277 C 43 C 2 ATM 857 $852,121 2 $1,704,242

LPG_COOL CS-shell/CS-tube

U-Tube 109 C 43 C 12 ATM 174 $250,575 1 $250,575

DC_L_C CS-shell/CS-tube

U-Tube -26 C -25 C 1 ATM 146 $234,866 1 $234,866

DC_H_C CS-shell/CS-tube

U-Tube 15.6 C 21 C 1 ATM 238 $306,680 1 $306,680

DC_H_REBOIL CS-shell/CS-tube

Kettle 1,065 C 1,034 C 28 ATM 75 $957,815 2 $1,915,630

DC_L_REBOIL CS-shell/CS-tube

Kettle 1,034 C 1,032 C 28 ATM 25 $412,514 1 $412,514

Total $73,167,486

69

TABLE 19 - REACTOR SUMMARY

Reactor

Material

HX Area (m2)

Vessel Volume (m3)

Process Conditions

Cost per Reactor

# of Reactors

Total Cost

Syngas

SS Clad

5276 616 1976 C @ 28 ATM

$17,238,130 13 $258,571,956

FTR CS 5276 533 226 C @ 28 ATM $28,702,399 15 $373,131,187

HI CS 215 215 192 C @ 27 ATM $8,897,738 1 $8,897,738

Total $640,600,881

TABLE 20- TOWER SUMMARY

Tower Material

Type Capacity (Volume

m3)

Nmin Nactual Min. Reflux Ratio

Actual Reflux Ratio

Diameter

Cost # of Towers

Total Cost

DC_HEAVY

Carbon Steel

Trayed

412 9 25 0.1 0.14 5.4 m $6,104,417

1 $6,104,417

DC_LIGHT Carbon Steel

Trayed

46 4 15 0.5 0.6 2.2 m $1,260,490

1 $1,260,490

Total $7,364,907

TABLE 21- TRAYS SUMMARY

Trays Material Type Capacity (Area m2)

Cost # of Trays Total Cost

DC_HEAVY Carbon Steel

Sieve 22.5 $69,049 23 $1,588,130

DC_LIGHT Carbon Steel

Sieve 3.7 $10,893 13 $141,643

Total $1,729,774

70

TABLE 22- PUMP SUMMARY

Pump Material Type Capacity (Shaft Power

kW)

Cost # of Pumps

Total Cost

REF_1 Carbon Steel

Centrifugal 161 $1,236,199 8 $9,889,590

REF_2 Carbon Steel

Centrifugal 255 $1,684,060 4 $6,736,241

REF_3 Carbon Steel

Centrifugal 10 $145,430 1 $145,430

HI_1 Carbon Steel

Centrifugal 256 $1,755,184 1 $1,755,184

P1 Carbon Steel

Centrifugal 279 $2,005,991 4 $8,023,964

Total $26,550,409

TABLE 23- FIXED ROOF TANK SUMMARY

Tanks Type Capacity (Volume m3)

Cost # of Tanks Total Cost

Naphtha Fixed Roof 22,712 $1,328,450 4 $5,313,800LPG Fixed Roof 27,793 $1,584,878 1 $1,584,878

Diesel Fixed Roof 29,467 $1,669,638 41 $68,455,188

Total $75,353,867

71

Source of Costs To estimate the capital costs for the required equipment, Appendix A “Cost Equations and Curves for the CAPCOST Program” from Analysis, Synthesis, and Design of Chemical Processes by Turton, et al. was referenced. The literature referred a CEPCI value of 397, which is lower than the current equipment CEPCI value of 724.1. Thus, after determining the equipment cost with the original CEPCI value, the cost estimates were projected.

Purchased Equipment CostsTo determine the purchased equipment cost, which assumes ambient operating pressure and carbon steel construction, Equation 1 was used.

log10Cpo=K1+K2 log10 ( A )+K3 ¿¿¿¿ Equation

1

Where, A refers to the capacity or size parameter for the equipment and K1, K2, and K3

refers to equipment specific constants. Provided in Table 25 are the size parameters and constants for the various equipment that were costed. It is also important to note that when the size/capacity of an equipment exceeded those specified in Table 25, the equipment was split into separate equally sized parts.

For example, to cost compressor H2O_2 with a fluid power of 4456 kW, the compressor was split into two equal sections with capacity of 2228 kW, each. The equipment was then costed:

log10Cpo=2.2897+1.3604 log10 (2228 )−0.1027¿¿¿¿

C po=$493,273

This equipment cost however, refers to just one of two compressors required for H2O_2. Thus, the total equipment cost is $986,546. It is also important to note that this is only the purchased equipment cost and does not include pressure and material factors and the current CEPCI value, which will be further calculated in the following pages.

The same procedure was taken to determine the purchased equipment costs for all other equipment.

Pressure FactorsBecause the purchased equipment cost assumes ambient pressure conditions, pressure factors were applied to more accurately estimate the costs for the equipment. Two expressions were used to determine this pressure factor. Equation 2 refers to pressure factors for process vessels whereas, Equation 3 refers to pressure factors for other process equipment.

72

FP, vessel=

( P+1 ) D2[850−0.6 ( P+1 )]

+0.0315

0.0063for tvessel>0.0063m

Equation 2

Where, D (meters) refers to the diameter of the process vessel, P (barg) refers to the pressure. For FP, vessel less than 1, the pressure factor was determined to be 1 and for pressure below -0.5 barg, the pressure factor was determined to be 1.25.

For example, to determine the pressure factor for the vertical vessel FLASH_COMP1 operating at 2.4 barg with a diameter of 2 meters, the following was calculated.

FP, vessel=

(2.4+1 )22[850−0.6 (2.4+1 )]

+0.0315

0.0063=5.63

For other process equipment, Equation 3 was used to determine the pressure factors.

log10 FP=C1+C2 log10P+C3 ¿¿¿¿ Equation 3

Where, C1, C2, and C3 refer to the equipment specific constants and P refers to the pressure (barg). The constants and pressure ranges are provided on Table 26.

To determine the pressure factor for PROP_COND, the following was calculated.

log10 FP=0.03881−0.11272log109+0.08183¿¿¿¿

FP=1.013

Material factors and Bare Module FactorsFollowing the calculation of the pressure factors, the material and bare module factors were also determined. It is important to note that there are various expressions for each equipment type. Provided in Equation 4 is the expression for determining the bare module factor FBM from the material factor, FM and the pressure factor, FP.

CBM=C po FBM=C p

o (B1+B2 FM F P ) Equation 4

The material factor was determined from a figure in Appendix A of Turton et al. (Figure A.8). For convenience, the material factors for the heat exchangers, process vessels, and pumps for carbon steel and stainless steel are provided on

. The constants B1 and B2 are also provided on the same table.

73

Thus, to determine the bare module cost for the vertical vessel, FLASH_COMP1, the following can be calculated.

CBM=C po FBM=C p

o (B1+B2 FM F P )

FBM=(2.25+1.82 (1 ) (5.63 ) )FBM=12.5

Thus,

CBM=C po FBM= ($15,394 ) (12.51 )=$ 192,558

Thus, the bare module cost for the vertical vessel FLASH_COMP1 is $192,558, assuming a CEPCI value of 397. To account for the current CEPCI value of 724.1, the following calculation was done.

$192,558×( 724.1397 )=$ 351,212

Similar calculations were performed to determine the bare module costs for other heat exchangers, process vessels, and pumps. To determine the bare module costs for compressors, drives, fired heaters, and sieve trays, the following equations were used.

TABLE 24- BARE MODULE COST EQUATIONS

Equipment Type Equation for Bare Module CostCompressor without drives CBM=C p

o FBMEquation 5

Drives CBM=C po FBM

Equation 6

Fired Heaters CBM=C po FBM FPFT

Equation 7

Sieve Trays CBM=C po NFBM Fq

Equation 8

It is important to note that FT is the superheat correction factor for steam boilers. Since steam boilers were not used in this design, FT is 1. N for sieve trays also refers to the number of trays whereas, Fq refers to the quantity factor for trays,

where, log10 Fq=0.4771+0.08516 log10N−0.3473¿¿for N<20

and Fq=1 for N>20.

The bare module factors, FBM, for the equipment not covered in

are provided in Table 28. It is important to note that with the exception of the nonreactive heaters, all equipment was estimated to be made of carbon steel.

74

TABLE 25- EQUIPMENT COST DATA TO BE USED WITH EQUATION 1

Equipment Type Equipment Description

K1 K2 K3 Capacity, Units Min Size

Max Size

Compressors Centrifugal 2.2897 1.3604 -0.1027 Fluid power, kW 450 3000Rotary 5.0355 -1.8002 0.8253 Fluid power, kW 18 950

Drives Steam Turbine 2.6259 1.4398 -0.1776 Shaft power, kW 70 7500Furnaces Nonreactive fired

heater7.344 -1.1666 0.2028 Duty, kW 1000 100,000

Heat Exchangers u-tube 4.1884 -0.2503 0.1974 Area, m2 10 1000Kettle reboiler 4.4646 -0.5277 0.3955 Area, m2 10 100

Process Vessels Horizontal 3.5565 0.3776 0.0905 Volume, m3 0.1 628Vertical 3.4974 0.4485 0.1074 Volume, m3 0.3 520

Pumps Centrifugal 3.3892 0.0536 0.1538 Shaft power, kW 1 300Towers Trayed 3.4974 0.4485 0.1074 Volume, m3 0.3 520Tanks API-fixed roof 4.8509 -0.3973 0.1445 Volume, m3 90 30000Trays Sieve 2.9949 0.4465 0.3961 Area, m2 0.07 12.30

TABLE 26- PRESSURE FACTORS FOR PROCESS EQUIPMENT

Equipment Type Equipment Description

C1 C2 C3 Pressure Range (barg)

Compressors Centrifugal 0 0 0 -Rotary 0 0 0 -

Drives Steam Turbine 0 0 0 -Furnaces Nonreactive fired

heater0 0 0 P<10

0.1347 -0.2368 0.1021 10<P<200Heat Exchangers u-tube/kettle reboiler 0 0 0 P<5

0.03881 -0.11272 0.08183 5<P<140Pumps Centrifugal 0 0 0 P<10

-0.3935 0.3957 -0.00226 10<P<100Tanks API-fixed roof 0 0 0 P<0.07Trays Sieve 0 0 0 -

75

TABLE 27- MATERIAL FACTORS FOR HEAT EXCHANGERS, PROCESS VESSELS, AND PUMPS

Equipment Type

Equipment Description

Material of Construction

Material Factor FM

B1 B2

Heat Exchanger

U-tube/kettle reboiler CS-shell/CS-tube 1 1.63 1.66

CS-shell/SS-tube 1.8SS-shell/SS-tube 2.7

Process Vessels Horizontal/vertical CS 1 (horizontal) 1.49

(horizontal) 1.52

(including towers) SS clad 1.7 (vertical) 2.25 (vertical) 1.82SS 3.1

Pumps Centrifugal CS 1.6 1.89 1.35SS 2.3

TABLE 28- BARE MODULE FACTORS FOR EQUIPMENT

Equipment Type

Equipment Description

Material of Construction

Bare Module Factor FBM

Compressors Centrifugal CS 2.7SS 5.7

Rotary CS 2.4SS 5.0

Drives Steam turbine - 3.5Fired Heater Nonreactive heater CS 2.1

Alloy Steel 2.5SS 2.8

Trays Sieve CS 1SS 2

76

Capital EstimateThe total capital estimate can be separated into two components: the working capital and the fixed capital investment. The working capital refers to the money invested in 1-month supply of the raw materials, finished products in stock, cash kept for monthly payments, and accounts payable. Essentially, the working capital is the cash required on a monthly basis to keep the plant operating efficiently. Thus, the working capital was estimated by summing the monthly costs related to utilities, labor, raw materials, products, general expenses, plant overhead, and operating supplies. A detailed breakdown of the working capital is provided on Table 29.

TABLE 29- WORKING CAPITAL

1-Month of Amount (in $1,000’s)

Utility Costs 703Operating Labor Costs 2,809

Raw Materials 63,486Products 136,526

General Expenses 12,221Plant Overhead 1,826

Operating Supplies 2,083Total 219,654

The fixed capital investment can be separated into two components: direct costs and indirect costs. The direct costs refer to the material and labor required for the installation of the facility. The indirect costs on the other hand, refer to costs not directly related to the actual expenses for the completion of the facility. A breakdown of both direct and indirect costs is provided on Table 30 and Table 32.

TABLE 30- DIRECT COSTS ESTIMATE

Direct Costs Amount (in $1,000’s)

Equipment Installation, etc.

1,555,717

Buildings, process and auxiliary

466,715

Service facilities and yard improvements

513,387

Land 3,889Total 2,539,709

It is important to note that “Equipment Installation, etc.” refers to not only the total bare module equipment cost specified in the previous section, but also refers to the costs related to instrumentation and controls, piping, and electricals. The breakdown of “Equipment Installation, etc.” is provided on Table 31. Note that the costs also include installation.

77

TABLE 31- EQUIPMENT INSTALLATION ESTIMATE

Equipment Installation, etc.

Amount (in $1,000’s)

Bare Module 1,169,712Instrumentation 116,971

Piping 128,668Electrical 140,365

Total 1,555,717

Also included in fixed capital investment are the indirect costs. A breakdown of indirect costs is provided on Table 32.

TABLE 32- INDIRECT COSTS ESTIMATE

Indirect Costs Amount (in $1,000’s)

Engineering and Supervision

507,942

Legal Expenses 25,397Construction

Expenses/Contractor’s Fee

253,971

Contingency 214,801Total 1,002,110

Where, “Engineering and Supervision” refer to costs related to the designing and engineering of the plant facilities, drafting, cost engineering, consultant fees, and engineering supervision and inspection. “Construction Expenses/Contractor’s Fee” refers to costs related to field expenses that are not directly related to the installation of equipment/ facilities. “Contingency” refers to costs that will compensate for unexpected events, such as floods, strikes, and price changes.

With the estimation of the working capital, direct costs, and indirect costs, the total capital investment was determined. Provided below is a detailed description of the total capital investment on Table 33.

TABLE 33- TOTAL CAPITAL INVESTMENT ESTIMATE

Total Capital Investment

Amount (in $1,000’s)

Working Capital 219,654Direct Costs 2,539,709

Indirect Costs 1,002,110Total 3,761,473

78

Economic AnalysisAfter estimating the total capital investment and the operating costs, the sales revenue was also projected to develop a discounted cash flow rate of return. The sales revenue breakdown is provided on Table 34.

TABLE 34- SALES REVENUE

Oil Products Price ($/bbl)

Volume (bbl/day)

Revenue ($/day)

Operating Time

(days/yr)

Revenue ($/year)

LPG 51.78 1,877 $97,209 330 $32,079,571Naphtha 75.00 19,069 $1,430,175 330 $471,958,911Diesel 90.00 34,101 $3,069,090 330 $1,012,799,655

Utility Products

Price ($/klb)

Weight (klb/day)

Revenue ($/day)

Operating Time

(days/yr)

Revenue ($/year)

HP Steam 4.00 68,783 $275,136 330 $90,794,570Steam

Condensate2.00 46,483 $92,967 330 $30,678,784

Total $1,638,311,490

To determine the quality of the investment, the cash flow for the plant was projected. Provided below are the assumptions taken into account when determining the cash flow.

a. Construction phase of 3 years

b. Projected plant life of 15 years

c. 15 year straight line depreciation

d. Projected 3% yearly inflation

e. First year of production at 50% capacity

f. Shutdown cost equals 10% of total capital investment

A time phase diagram of the cumulative cash flow is provided on Figure 1 and a detailed table (Table 35) is provided at the end of the report. It is important to note that the first year of construction refers to only 10% of the total capital investment while the third year refers to 60% of the total capital investment. This is due to the fact that most of the capital investment will be used during the last year, when the majority of the equipment will be installed.

79

0 2 4 6 8 10 12 14 16 18 20

-5000000

-4000000

-3000000

-2000000

-1000000

0

1000000

2000000

3000000

4000000

Cumulative Cash Flow

Cumulative Cash Flow Discounted Cumulative Cash Flow

Years since Beginning of Construction

Billions

of

$

FIGURE 11- CUMULATIVE CASH FLOW

80

Following the three years of construction, operations begin at 50% capacity for year 4. From the end of year 4, there is a straight line increase in cash flow. Despite the positive cash flow, since the initial capital investments are so high, it take approximately 12 years for the plant to break even. Thus, the plant needs to operate for 9 years to make up for the losses in the first three years. Following year 12, there is a net positive cumulative cash flow and at the end of year 18, there is shut down. The shutdown cost was estimated to be 10% of the total capital investment. This negative impact however is slightly mitigated by the return of the working capital at the end of production. Hence, the slight negative cash flow for year 19. A detailed table of the year-by-year cash flow and discounted cash flow analysis are provided in Table 35 at the end of the report.

Despite the positive cumulative cash flow by the end of plant life, the discounted cash flow rate of return is only 7.8%, due to the large capital costs and the long payback period. Furthermore, the project is very sensitive to market changes in the price of oil, natural gas, and environmental laws. From preliminary research on the sensitivity on the rate of return for the proposed GTL process suggests that the DCFROR is most sensitive to diesel market price. 10% change in sales price of diesel could increase or decrease current DCFROR as high as 35%. DCFROR is also very sensitive to oxygen price and naphtha price because current design includes large amount of continuous oxygen feed for syngas unit. Change in oxygen price, therefore, can make or deny current process. Same for naphtha, as large portion of sales revenue will come from selling naphtha.

Because the plant is only 63% energy efficient and produces large CO2 emissions, regulations may force a less profitable venture. Since there are other investment options (stocks, bonds, mutual funds) that often times generate a better return on investment, it is advised that this design is not implemented.

81

Innovation & OptimizationCurrent preliminary design covers the basics of GTL process. Proposed design would work with no problem as it is; however, there are further optimization and refinement issues that might increase the plant’s integration as a whole and productivity. Such improvements on plant design, as a result, could change the plant’s economic return.

Fisher-Tropsch reaction is central to the GTL process. Improvements on FTR unit, therefore, can bring significant changes to the GTL process as a whole. FTR unit design can be improved in terms of size and conversion. Designing FTR requires intense optimization because the conditions for high conversion contradict each other. High operating temperature gives high rate of reaction while low operating temperature gives good selectivity towards desired products. In order to accommodate for such contradicting conditions, current design of FTR unit includes 15 packed-bed reactors in parallel with total of 24,293 tubes. However, this geometry might not be the best choice and further optimization may allow smaller reactor system, which could decrease capital investment, and increased conversion, which could increase product sales revenue.

Also, the GTL process produces significant amount of methane with methane feed. 4 wt. % of FTR product stream and 6 wt. % of HI product stream is methane. Considering the large capacity of the plant, the amount of methane produced by the system cannot be overlooked. Current design of GTL plant uses all of produced methane as fuel gas. Some of methane is burnt to provide heat for syngas preheater and the rest is burnt to produce HP steam which can be used in the plant as well as sold to outside purchaser. Recycling produced methane and redirecting it to the syngas unit can actually increase the amount of final products. However, in order to do so, more sophisticated separation system is necessary. The balance between the capital investment increase due to additional separation system and the increase in sales revenue coming from recycling methane must be found prior to decision making.

Additionally, existing separation system could be further refined and optimized to yield more desirable products. However, as mentioned above, the balance between the increase in the capital investment due to building better separation systems and the increase in sales revenue due to having better separation system must be further investigated for consideration.

82

References

[1] EIA, "U.S. Energy Information Administration," 29 August 2012. [Online]. Available: http://www.eia.gov/tools/faqs/faq.cfm?id=58&t=8. [Accessed May 2013].

[2] EIA, "Annual Energy Outlook," 2012.

[3] ICIS, "Aramco cuts April propane, butane prices by $60-105/tonne," 1 April 2013. [Online]. Available: http://www.icis.com/Borealis/Article.asp?p=1&q=BFB3C6D1D8BDE2B6CCAD8DB96EB0D9CAAFDCC1D48DAEE7B281AED7B8E0B4D5D6B0EC&id=B28496A19C79AA. [Accessed 12 May 2013].

[4] ICIS, "Europe naphtha trades up $50/t in a sustained week-long bull run," 10 May 2013. [Online]. Available: http://www.icis.com/Articles/2013/05/10/9667049/europe+naphtha+trades+up+50t+in+a+sustained+week-long+bull+run.html. [Accessed 13 May 2013].

[5] ICIS, "US gasoline, diesel prices to decline through end of year – EIA," 10 October 2012. [Online]. Available: http://www.icis.com/Articles/2012/10/10/9602866/us-gasoline-diesel-prices-to-decline-through-end-of-year-eia.html. [Accessed 13 May 2013].

[6] U.S. Energy Information Administration, "Natural Gas Prices," December 2012. [Online]. Available: http://www.eia.gov/dnav/ng/ng_pri_sum_dcu_nus_m.htm. [Accessed 13 May 2013].

[7] United States Environmental Protection Agency, "Nitrogen Dioxide," 4 April 2013. [Online]. Available: http://www.epa.gov/oar/nitrogenoxides/. [Accessed 13 May 2013].

[8] X. Wang and M. Economides, Advanced Natural Gas Engineering, Gulf Publishing Company, 2009.

[9] Design Institute for Physical Properties, "DIPPR Project 801," 2012. [Online]. Available: http://www.knovel.com/web/portal/browse/display?_EXT_KNOVEL_DISPLAY_bookid=1187&VerticalID=0.

[10] G. H. E. G. J. M. C. Bouchy, "Fisher-Tropsch Waxes Upgrading via Hydrocracking and Selective Hydroisomerization," Oil & Gas Science and Technology, pp. 91-112, 2009.

[11] C. G. J. W. C. F. E.L. First, "Computational Characterization of Zeolite Porous Networks: An Automated Approach," Computer-Aided Systems Laboratory, Princeton University, 2013.

[12] S. Parkash, Refining Processes Handbook, Boston: Elsevier, 2003.

[13] "CAMEO Chemicals: Database of Hazardous Materials," National Oceanic and Atmospheric Administration, [Online]. Available: http://cameochemicals.noaa.gov/.

[14] C. Engineering, "Chemical Engineering Plant Cost Index," November 2011. [Online]. Available: http://www.scribd.com/doc/121019752/CEPCI. [Accessed May 2013].

83

[15] K. A. Maurice Stewart, Gas-Liquid and Liquid-Liquid Separators, Elsevier, 2008.

[16] E. J. H. D. K. R. J.D. Seader, Separation Process Principles: Chemical and Biochemical Operations, John Wiley & Sons, Inc..

84

Engineering CalculationsThis section shows the hand calculation conducted for design purposes. Each and every calculation for each problem is not shown. Rather, this section will cover calculations for each type of problem.

Equipment SizingProposed design of GTL plant includes number of equipment of different type. Each one was designed and optimized through a simulation program; however, its size has to be calculated additionally. This section shows one sample calculation for sizing each type of equipment used in the design of GTL plant.

Fired HeatersThere are two fired heaters in this design: SU_PREHEAT and HP_STM_BOIL. They are the same kind and their sizes were calculated using the same method. For this section, sizing calculation for SU_PREHEAT was shown.

The size of fired heater is determined by its heat-exchanging areas in convective and radiative section. The fundamental equation for heat transfer is as follows:

Q=Ah

where Q , A and h represent heat duty, heat-exchanging area and heat transfer coefficient, respectively.

It was assumed that 60% of total heat transfer necessary occurs in the radiative section while the other 40% occurs in the convective section. Also, typical heat transfer coefficient values for convective and radiative sections were assumed to be 10.86 kW/m2 and 30 kW/m2, respectively. Also, total heat duty necessary for this stage was obtained to be 290,000 kW from the necessary temperature rise for the process stream upon passing through the preheater.

Therefore,

Aconv= (40% )× Qhconv

=0.4×290,000kW10.86kW /m2

=10,682m2

Arad= (60%)× Qhrad

=0.6×290,000kW30kW /m2

=5,800m2

CompressorsAll the compressors present in the current design are of same type, and therefore, this section will show calculations conducted to size 3PH_1

The size of a compressor is determined by its BHP (brake horse power) which indicated the initial power output the compressor needs to provide before any power loss due to the mechanical system. BHP is calculated by the following equation:

BHP=GHPηM

where

GHP=ρ×Q×H ×g×10−6

3.6×ηv

85

GHP is gas horse power which indicates the real power output that the compressor applies to the fluid. ρ, Q, H, g, ηv and ηM represent density and volumetric flow rate of the fluid through the compressor, head of the compressor, gravitational acceleration, efficiency to account for the heating effect during suction stroke and for gas leak and mechanical efficiency, respectively.

Therefore,

BHP=

ρ×Q×H ×g×10−6

3.6×ηv

ηM

=

5.171kgm3×46,825

m3

hr×4,167m×9.81

m2

s×10−6

3.6×0.950.8

=3,617 kW

PumpsSimilarly, all the pumps in the design are of same type. Sample calculation for REF_1 was shown in this section.

The size of pump is determined by its NPSHA value. It is obtained by the following equation:

NPS H A=po−pv

ρg

where po, pv, ρ and g represent atmospheric pressure, vapor pressure of the liquid pumped, density of fluid and gravitational acceleration.

Therefore,

NPS H A=101,325

N

m2−6,556 N

m2

990.0kgm3×9.81

ms2

=9.7m

Heat ExchangersHeat exchangers in the current GTL process design are all calculated in the same way. In this section, calculation to size R134A_2_COND was shown.

Similar to fired heater, size of a heat exchanger is determined by its heat-exchanging are. The fundamental heat transfer equation is shown below:

Q=UAF ΔT LM

where Q, U, A, F and ΔT LM represent heat duty, overall heat transfer coefficient, heat-exchanging area, correction factor and log mean temperature difference. Appropriate heat duty, correction factor and log mean temperature difference were provided by the simulation program in order to provided desired heat exchange. For this case, U of 850 W/m2K was assumed.

Therefore,

A= QUFΔT LM

= 12,761kW

850W

m2K×0.947×12.73K

=1,247m2

86

Flash TanksThere are two general types of flash tanks used in the proposed design: two-phase and three-phase flash tanks. Two-phase flash tanks provide separation due to phase difference (vapor and liquid phase) while three-phase flash tanks provide separation due to both phase difference and immiscibility of two different types of liquids.

For sizing two-phase flash tanks, volume of the tank capable of holding the feed for the given residence time must be calculated. Below calculations show sizing calculations conducted for H2O_RMVL.

V=(Q¿¿G+QL)×τ ¿

where V, QG, QL and τ represent volume of a tank, gas volumetric flow rate, liquid volumetric flow rate and residence time, respectively.

Therefore,

V=(58850 m3

hr×309

m3

hr )×0.15min≈148m3

Baseline to determine diameter and height was that the height must not exceed 4 times the diameter. Accordingly, the diameter and height were determined to be 3.75 m and 13.39 m, respectively.

For sizing three-phase flash tanks, diameter was calculated first. Three different diameters were calculated: minimum diameter to settle liquid droplets through a gas, minimum diameter to settle water droplets from liquid oil phase and minimum diameter to settle oil droplets from water. The largest among those three minimum diameters was chosen for the diameter of tank [15].

d1=√3500×(TZ Qg

P )[( ρg

ρl−ρg)CD]

1 /2

d2=√2550×(Qoμo

ΔSG )d3=√1.59×104×(Qo μw

ΔSG )where d, T, Z, Qg, P, ρg, ρl,CD, Qo, μ and Δ SG represent vessel inside diameter (mm), operating temperature (K), gas flow rate (scm/h), operating pressure (kPa), density of gas, density of liquid (kg/m3), drag coefficient, liquid flow rate (scm/h), viscosity of liquid (Pa s) and difference in specific gravity of water and liquid oil. For the preliminary calculations, compressibility of 1 and drag coefficient of 0.1 were assumed.

Therefore,

d1=√3500×( 472.21K×8987.6scmhr

3947.5kPa )[( 29.127kg

m3

662.4+867.42

kgm3−29.127

kgm3 )×0.1]

1 /2

=486.5739mm

87

d2=√2550×( 355.11 scmhr

×0.981Pa s

1−0.784 )=2,027.921mm

d3=√1.59×104×( 355.11 scmhr

×0.1344 Pas

1−0.784 )=5,063.835mm

d3 is the largest of all three. Therefore, the diameter was determined to be 5.06 m.

To calculated length (height) of the tank, following correlation was used [15].

Lss=ho+hw+d+1016

where

ho+hw=τ (Qo+Qw )

4.713×10−8d2

Lss, ho + hw and τ represent seam-to-seam length of a tank (mm), liquid (oil and water) level in the tank (mm) and residence time (hr). One thing to note is that the d2 term for calculating ho + hw must be in meters, not millimeters.

Therefore,

Lss=( 1060 hr )(355.11 scm

hr+389.88 scm

hr )4.713×10−8×(5.06m)2

+5,063.835+1,016=12,238mm

Therefore, the diameter and height of this three-phase flash tank are 5.06 m and 12.24 m, respectively.

Distillation ColumnsDistillation columns are sized by specifying its column and diameter. Below calculations show sizing of DC_LIGHT.

The diameter of a distillation column is determined so that the column is large enough to hold the liquid and vapor overflows at a given point. Such diameters were calculated at two points: at the top and at the bottom and the larger one was chosen as a final diameter for a column.

Following equation was used to calculate the diameter [16].

d=[ 4V MV

f U f π (1−AD /A ) ρV ]1/2

where

U f=C ( ρL−ρV

ρV)1/2

C=FST FFCF

FST= (σ /20 )0.2

88

and CF is determined by using the graph below.

FIGURE 12- ENTRAINMENT FLOODING CAPACITY IN A TRAYED TOWER

where

FLV=( LM L

V MV)( ρV

ρL)1/2

V, L, MV, ML, f, Uf, Ad/A, ρV , ρL, C, FST, FF, CF, σ and FLV represent vapor overflow, liquid overflow, molar weight of vapor flow, molar weight of liquid flow, flooding velocity correction factor, flooding velocity, fraction of effective tray area, vapor density, liquid density, capacity of Souders and Brown, surface tension factor, foaming factor, entrainment flooding capacity and abscissa ratio.

All the calculations were conducted in British unit because the entrainment flooding information is given in British unit.

Tray spacing of 24 inches (61 cm) was chosen to optimize column height and molar overflow.

For minimum diameter at the bottom,

FLV=( 3.42×103lbmol

h×52.245

lblbmol

2.87×103lbmol

h×47.58

lblbmol

)( 3.35lb

ft3

28.7lbft3

)1/2

=0.4465

89

CF=0.2 ft /s

FST=(3.68 dynecm

/20)0.2

=0.7128

C=0.4465×0.2×0.7128=0.1426 ft / s

U f=0.1426fts×(28.7

lb

ft3−3.35

lb

ft3

3.35lbft 3

)1 /2

=0.3923 ft /h

dbot=[ 2.87×103lbmol

h×47.58

lblbmol

0.8×0.3923fth×π × (1−0.1 ) ×3.35 lb

f t3]1/2

=7.15 ft=2.18m

For minimum diameter at the top,

FLV=( 1.88×103lbmol

h×43.78

lblbmol

3.55×103lbmol

h×42.44

lblbmol

)( 3.68lb

ft3

62.3lbft3

)1/2

=0.1330

CF=0.35 ft /s

FST=(8.02 dynecm

/20)0.2

=0.8330

C=0.1330×0.35×0.8330=0.2915 ft / s

U f=0.2915fts×( 62.3

lb

ft3−3.68

lb

ft3

3.68lbft3

)1 /2

=1.1641 ft /h

d top=[ 3.5 .5×103lbmol

h×42.44

lblbmol

0.8×1.1461fth×π × (1−0.1 )×3.68 lb

f t3]1/2

=4.15 ft=1.27m

Because minimum diameter required at the bottom is greater, the column diameter was determined to be 2.18 m.

Column height is calculated using the following equation:

h=N trays×h tray+3.048

h, Ntrays and htray represent column height, number of trays and tray spacing, respectively.

90

In this case,

h=15×¿Thus, the column diameter and height for DC_LIGHT were calculated to be 2.18 m and 12.19 m, respectively.

91

Computer Programs

MatLab Code to Model FTR

%FTR DESIGN %These are the inlet conditions that come from the outlet of the SynGas%Unit. Any changes to the syngas will change these inputs clear allformat longclc z =linspace(0,(60*2.54*12/100),1e3); %Generates a vector representing the length of the reactor in 1e3 stepsdz = diff(z); %Step size v = zeros(1,length(z)); %velocity of the gas (m/s) mol = zeros(7,length(z)); %Generates a vector to hold the molar flow ratesmol(:,1) = [4.6856; %Sets the inlet molar flow rates 129.5567; 12658.9316; 6321.5155; 1189.1856; 41.6024; 0]; %H2O CH4 H2 CO CO2 N2 HC and all in mol/s Cp = 1e-3*[35.12; %These represent the Cp values for the above components 45.95; 29.31; 39.74; 0; 29.54; 0]; %kJ/molK Cp_avg = sum(Cp.*mol(:,1))/sum(mol(:,1)); %Calculates average heat capacity D = 1.25*2.54/100; %diameter of the tubes (m)Ac = pi()*D^2/4; %Cross-sectional area of a pipe (m^2) P = zeros(1,length(z)-1); %Pressure VectorP(1) = 27.5; %Inlet pressure (Atm) T = zeros(1,length(z)-1); %Temperature VectorT(1) = 495; %Inlet temperature (K)

T_ext = 490; %Cooling water temperature (K) Rg = 0.08206/10^3; %Ideal gas constant m^3*atm/molK Dp = 0.08*2.54/100; %Catalyst particle diameter in m mu_constants = [0.000000017096, 1.1146, 0, 0, 0; %Viscosity Constants 0.00000052546, 0.59006, 105.67, 0, 0; 0.0000001797, 0.685, -0.59, 140, 0; 0.0000011127, 0.5338, 94.7, 0, 0; 0.000002148, 0.46, 290, 0, 0; 0.000000017096, 1.1146, 0, 0, 0; 0,0,0,0,0]; mu = zeros(size(mu_constants,1),length(z)-1); %Viscosity vectormu(:,1) = (mu_constants(:,1).*T(1).^mu_constants(:,2))./(1+mu_constants(:,3)./T(1)+mu_constants(:,4)./T(1).^2);MW = [18.01528; %Molecular weights 16.042; 2.01588; 28.0101; 44.0095; 28.0134; 0]; mu_avg = zeros(1,length(z)-1); MW_avg = zeros(1,length(z)-1); mu_avg(1) = sum(mu(:,1).*MW.*mol(:,1))/sum(MW.*mol(:,1)); %Average viscosityMW_avg(1) = sum(MW.*MW.*mol(:,1))/sum(MW.*mol(:,1)); %Average Molecular weight rho_mol = zeros(1, length(z)); %Molar density mol/m^3rho_kg = zeros(1,length(z)); %Density kg/m^3 x = zeros(7,length(z)); %mol fraction rho_mol(1) = P(1)/Rg/T(1); %mol/m^3rho_kg(1) = rho_mol(1)*MW_avg(1)/1000; %kg/m^3Q = sum(mol(:,1))/rho_mol(1); %Volumetric flow rate m^3

92

x(:,1) = mol(:,1)/sum(mol(:,1)); %Reaction Constantsk1 = 0.0173/3600*100^3*.4; %gmol CO/s/m^3/atm^2k2 = 4.512; %1/atm^2T1 = zeros(1,length(T));T2 = zeros(1,length(T));T3 = zeros(1,length(T)); T1(1) = exp(-4492*(1/T(1)-1/473));T2(1) = exp(8237*(1/T(1)-1/473));T3(1) = exp(-10000*(1/T(1)-1/473)); %Rate for each chemical compoundrco = zeros(1, length(T));rh2 = zeros(1, length(T));rch4 = zeros(1, length(T));rhc = zeros(1, length(T));rh2o = zeros(1, length(T)); rco(1) = -k1.*T1(1).*x(3,1).*x(4,1).*P(1).^2./(1+k2.*T2(1).*x(4,1).*P(1)).^2; %gmolCO/s/m^3rh2(1) = 2.*rco(1);rch4(1) = -rco(1).*0.03.*T3(1);rhc(1) = 0.7.*rch4(1); %Hydrocarbon rate assumed to be 0.7 mol/mol of CH4 (other HC in liquid phase)rh2o(1) = -rco(1); Nt = 364397; %Number of tubes specified N_reac = Nt/(0.319*(20*12/(D*100/2.54))^2.142); %Number of reactors v(1) = Q/Ac/Nt; %Velocity of the gas (m/s) Re_p = zeros(size(rho_kg,1),size(rho_kg,2)); %Particle reynold's number describing flow through packed bed Re_p(1) = rho_kg(1)*v(1)*Dp/mu_avg(1); Flux = zeros(7,length(z)); %Molar flux of compoundsFlux(:,1) = P(1)/Rg/T(1) * x(:,1)*v(1); G = MW_avg(1)*sum(Flux(:,1))*3600/100^2; %Inlet gas mass velocity g/hr/cm^2 U_0 = 0.385*G^0.8/(D*100)^0.2*1.055*5/9/3600*(1/0.3048)^2; % Overall heat transfer coefficient kJ/m^2/s/K Hr = 70200*1.055/.4536/1000; %Heat of reaction kJ/gmol for i=1:1:length(z)-1 %Enacting Euler's method through the reactor with step size dz Flux(4,i+1) = Flux(4,i) + dz(i)*rco(i); %Flux after one step through the reactor Flux(3,i+1) = Flux(3,i) + dz(i)*rh2(i); Flux(2,i+1) = Flux(2,i) + dz(i)*rch4(i); Flux(7,i+1) = Flux(7,i) + dz(i)*rhc(i);

Flux(1,i+1) = Flux(1,i) + dz(i)*rh2o(i); Flux(5,i+1) = Flux(5,i); Flux(6,i+1) = Flux(6,i); P(i+1) = P(i) -1.5*dz(i)*rho_kg(i)*v(i)^2/Dp*(1-0.4)/0.4^3*(150*(1-0.4)/Re_p(i)+1.75)/101325; %Pressure after one step through reactor T(i+1) = T(i)+ dz(i)*(-0.4*Hr*rco(i)/v(i)/rho_mol(i)/Cp_avg+2*U_0/v(i)/rho_mol(i)/Cp_avg/(D/2)*(T_ext-T(i))); %Temperature after one step through reactor %Viscosity evaluated at new T mu(:,i+1) = (mu_constants(:,1).*T(i+1).^mu_constants(:,2))./(1+mu_constants(:,3)./T(i+1)+mu_constants(:,4)./T(i+1).^2); %New average fluid properties mu_avg(i+1) = sum(mu(:,i+1).*MW.*Flux(:,i+1))/sum(MW.*Flux(:,i+1)); MW_avg(i+1) = sum(MW.*MW.*Flux(:,i+1))/sum(MW.*Flux(:,i+1)); rho_mol(i+1) = P(i+1)/Rg/T(i+1); %mol/m^3 rho_kg(i+1) = rho_mol(i+1)*MW_avg(i+1)/1000; %kg/m^3 x(:,i+1) = Flux(:,i+1)/sum(Flux(:,i+1)); %New reaction rates T1(i+1) = exp(-4492*(1/T(i+1)-1/473)); T2(i+1) = exp(8237*(1/T(i+1)-1/473)); T3(i+1) = exp(-10000*(1/T(i+1)-1/473)); rco(i+1) = -k1.*T1(i+1).*x(3,i+1).*x(4,i+1).*P(i+1).^2./(1+k2.*T2(i+1).*x(4,i+1).*P(i+1)).^2; %gmolCO/s/m^3 rh2(i+1) = 2.*rco(i+1); rch4(i+1) = -rco(i+1).*0.03.*T3(i+1); rhc(i+1) = 0.7.*rch4(i+1); rh2o(i+1) = -rco(i+1); v(i+1) = sum(Flux(:,i+1))*Rg*T(i+1)/P(i+1); Re_p(i+1) = rho_kg(i+1)*v(i+1)*Dp/mu_avg(i+1); end Conversion = ((Flux(4,1)-Flux(4,length(Flux)))/Flux(4,1)) %Conversion based on amount leavingP_drop = -(P(length(P))-P(1))*14.7 %Total pressure drop (Psi) v_avg = (v(length(v))+v(1))/2; %Average velocityResidence = z(length(z))/v_avg %Residence time based on average velocity s m_cat = z(length(z))*Ac*Nt*100^3/1000; %Total mass of catalyst

93

T_mod = m_cat/Q %Modified residence time

T_avg = 1/sum(rco)*sum(rco.*T) %Reaction weighted average Temperature

94

Computer Process Simulations

PRO/II Keyword Input File

$ Generated by PRO/II Keyword Generation System <version 9.1>$ Generated on: Sat May 11 23:37:40 2013TITLE PRINT SPTPRINT=ON, STREAM=ALL, RATE=M,WT TOLERANCE STREAM =0.01,-0.555556,0.01,0.01, TEMPERATURE=-0.0555556, & PRESSURE=0.005, DUTY=0.001, MISCELLANEOUS=0.003, FLASH=3E-6 DIMENSION METRIC, TEMP=C, PRES=ATM, WT=KG, TIME=HR, LENGTH=M, & FLENGTH=MM, LIQVOL=M3, VAPVOL=M3, LDENSITY=KG/M3, & VDENSITY=KG/M3, XDENSITY=DENS, SPVOL=M3/KG-MOL, & SPVVOL=M3/KG-MOL, ENERGY=KJ, WORK=KW, DUTY=KJ/HR, & CONDUCT=KW/MK, HTCOEF=KW/MK, FOUL=MK/KW, VISCOSITY=CP, & KVIS=CST, SURFACE=D/CM, STDTEMP=0, STDPRES=1, & STDVAP(M3/KG-MOL)=22.414, PBASIS(KG/CM2)=1.0332 SEQUENCE SIMSCI CALCULATION TRIALS=50, RECYCLE=ALL, TVPBASIS=37.7778, & RVPBASIS=APIN, COMPCHECK=CALC, MAXOPS=1000000, CDATA=FIX, & FLASH=DEFAULT, DVARIABLE=ONCOMPONENT DATA LIBID 1,HYDROGEN/2,H2O/3,OXYGEN/4,NITROGEN/5,CO2/6,CO/7,METHANE/ & 8,ETHANE/9,PROPANE/10,BUTANE/11,PENTANE/12,HEXANE/ & 13,HEPTANE/14,OCTANE/15,NONANE/16,DECANE/17,UNDECANE/ & 18,DODECANE/19,TRIDECAN/20,TETDECAN/21,PENDECAN/22,HXDECANE/ & 23,HDECANE/24,OCTDECAN/25,NONDECAN/26,EICOSANE, BANK=PROCESS LIBID 27,PROPCARB/34,R12/35,R134A/36,R22/37,R14, BANK=SIMSCI LIBID 33,AIR, BANK=SIMSCI,PROCESS PETRO 28,c2125,322.6,800.21,379.056 PETRO 29,c2629,386.5,809.201,426.444 PETRO 30,c3035,454.9,817.193,469.278 PETRO 31,c3647,572.2,826.185,528.056 PETRO 32,c48+,861.7,838.173,624 ASSAY FIT=ALTERNATE, CHARACTERIZE=TWU, MW=TWU, CONVERSION=API94, & GRAVITY=WATSONK, TBPIP=1, TBPEP=98, NBP=LV, & CURVEFIT=CURRENT, KVRECONCILE=TAILSTHERMODYNAMIC DATA METHOD SYSTEM=SRK, SET=SRK01, DEFAULTSTREAM DATA PROPERTY STREAM=METHANE, TEMPERATURE=37.778, PRESSURE=35.023, & PHASE=M, COMPOSITION(M,GMOL/D)=7,5.99096E8, SET=DEFAULT PROPERTY STREAM=STEAM, PRESSURE=28.218, PHASE=V, & COMPOSITION(M,GMOL/D)=2,2.99548E8, SET=DEFAULT PROPERTY STREAM=CO2, TEMPERATURE=37.778, PRESSURE=28.218, PHASE=M, & COMPOSITION(M,KGM/S)=5,0.63953, SET=DEFAULT PROPERTY STREAM=O2, TEMPERATURE=23.889, PRESSURE=35.023, PHASE=M, & RATE(M)=14850, COMPOSITION(M)=3,0.99/4,0.01, NORMALIZE PROPERTY STREAM=R134A_1, TEMPERATURE=-10, PRESSURE=2, PHASE=M, & COMPOSITION(M,KGM/H)=35,6000 PROPERTY STREAM=CW_2, TEMPERATURE=15.556, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=2,1.48452E7 PROPERTY STREAM=R134A_2, TEMPERATURE=-26, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=35,2500 PROPERTY STREAM=PROPANE, TEMPERATURE=-42, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=9,2500 PROPERTY STREAM=S65, TEMPERATURE=15.556, PRESSURE=1, PHASE=M, &

COMPOSITION(M,KGM/H)=2,3E6 PROPERTY STREAM=S68, TEMPERATURE=15.556, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=2,150000 PROPERTY STREAM=S72, TEMPERATURE=15.556, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=2,6600 PROPERTY STREAM=CW_1_HP, TEMPERATURE=15.556, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=2,1.6E7 PROPERTY STREAM=S81, TEMPERATURE=43.333, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=2,7000 PROPERTY STREAM=CW_3, TEMPERATURE=15.556, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=2,1000 PROPERTY STREAM=WATER_MAKEUP, TEMPERATURE=15.556, PRESSURE=1, & PHASE=M, COMPOSITION(M,KGM/H)=2,1000 PROPERTY STREAM=S83, TEMPERATURE=15.556, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=2,131313 PROPERTY STREAM=S87, PRESSURE=1, PHASE=L, COMPOSITION(M,KGM/H)=35, & 131313 PROPERTY STREAM=CW_HP, TEMPERATURE=25, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=2,1000 PROPERTY STREAM=S97, TEMPERATURE=15.556, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=2,40000 PROPERTY STREAM=CW_1, TEMPERATURE=15, PRESSURE=1, PHASE=M, & COMPOSITION(M,KGM/H)=2,1000 OUTPUT FORMAT=STREAMSUMM, STREAMS=HEAVIES,LIGHTS,NAPHTHA_2, & LIGHTS_1,HI_FEED,NAPHTHA_1, SPTPRINT=ON OUTPUT FORMAT=STREAMSUMM, STREAMS=LPG_PROD,NAPHTHA_PROD, & DIESEL_PROD, SPTPRINT=ON OUTPUT FORMAT=Fuel Output, STREAMS=LPG_PROD,NAPHTHA_PROD, & DIESEL_PROD, SPTPRINT=ON OUTPUT FORMAT=Combustion, STREAMS=S16, SPTPRINT=ON OUTPUT FORMAT=eric, STREAMS=WATER_MAKEUP,CW_3,HI_FEED, SPTPRINT=ON OUTPUT FORMAT=STREAMSUMM, STREAMS=NAPHTHA,DIESEL, SPTPRINT=ON OUTPUT FORMAT=STREAMSUMM, STREAMS=R134A_1_VAP,R134A_VAP_P1, & R134A_2_VAP,S37,S76,S79,PROPANE_VAP,S19,S40, SPTPRINT=ON FORMAT IDNO=Fuel Output,SID,TEMP,PRESSURE,RATE(WT),LFRAC,LIQUID, & ARATE(BBL/DAY),TOTAL,CFRAC(WT) FORMAT IDNO=Combustion,SID,NHV,ILHV,RATE(GV),VAPOR,ARATE FORMAT IDNO=eric,SID,NAME,DOUBLINE,PHASE,LINE,TEXT,LINE,RATE(M), & RATE(WT),RATE(LV),TEMP,PRESSURE,MW,HTOTAL,ENTHALPY,LFRAC,TR, & PR,ACENTRIC,WATSON,SDENSITY,SSPGR,SAPI,LINE,TEXT,LINE,VAPOR, & RATE(M),RATE(WT),ARATE,SRATE(GV),MW,ZKVA,ENTHALPY,CP,DENSITY, & TCOND,VISCOSITY,LINE,TEXT,LINE,LIQUID,RATE(M),RATE(WT), & SRATE(LV),MW,ZKVA,ENTHALPY,CP,DENSITY,SURFACE,TCOND, & VISCOSITY,TVPRXDATA RXSET ID=P_OXIDATION, NAME=Partial oxidation of methane REACTION ID=OXIDATION STOICHIOMETRY 2,2/3,-1.5/6,1/7,-1 RXSET ID=FTR, NAME=Fischer-Tropsch Reaction, & KINETICS(PROCEDURE)=FTR_KINE REACTION ID=R1

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STOICHIOMETRY 1,-2/2,1/6,-1/7,0.083273/8,0.003331/9,0.003331/ & 10,0.003331/11,0.005407/12,0.004925/13,0.00448/ & 14,0.004072/15,0.003698/16,0.003358/17,0.003046/ & 18,0.002766/19,0.00251/20,0.002277/21,0.002066/ & 22,0.001874/23,0.0017/24,0.001542/25,0.001398/ & 26,0.001268/28,0.004763/29,0.002447/30,0.002265/ & 31,0.001947/32,0.000824 RXSET ID=HI, NAME=Hydroisomerization Cracking REACTION ID=C2125 STOICHIOMETRY 7,0.20109/8,0.053643/9,0.25606/10,0.19426/ & 11,0.18631/12,0.15598/13,0.13415/14,0.11767/ & 15,0.1048/16,0.094472/17,0.13725/18,0.12595/ & 19,0.11636/20,0.10814/21,0.101/22,0.094739/ & 23,0.089213/24,0.084296/25,0.079893/26,0.075927/ & 28,-1 REACTION ID=C2629 STOICHIOMETRY 7,0.24092/8,0.064269/9,0.30678/10,0.23274/ & 11,0.22321/12,0.18688/13,0.16072/14,0.14098/ & 15,0.12556/16,0.11318/17,0.16443/18,0.15089/ & 19,0.13941/20,0.12956/21,0.121/22,0.1135/23,0.10688/ & 24,0.10099/25,0.095718/26,0.090966/29,-1 REACTION ID=C3035 STOICHIOMETRY 7,0.28356/8,0.075643/9,0.36107/10,0.27393/ & 11,0.26271/12,0.21995/13,0.18916/14,0.16593/ & 15,0.14779/16,0.13322/17,0.19353/18,0.1776/ & 19,0.16408/20,0.15248/21,0.14241/22,0.13359/ & 23,0.1258/24,0.11887/25,0.11266/26,0.10706/30,-1 REACTION ID=C3647 STOICHIOMETRY 7,0.35668/8,0.095146/9,0.45417/10,0.34457/ & 11,0.33045/12,0.27666/13,0.23794/14,0.20872/ & 15,0.18589/16,0.16757/17,0.24344/18,0.22339/ & 19,0.2064/20,0.1918/21,0.17914/22,0.16804/23,0.15824/ & 24,0.14952/25,0.14171/26,0.13467/31,-1 REACTION ID=C48 STOICHIOMETRY 7,0.53714/8,0.14329/9,0.68396/10,0.5189/ & 11,0.49764/12,0.41664/13,0.35832/14,0.31432/ & 15,0.27994/16,0.25235/17,0.3666/18,0.33641/ & 19,0.31082/20,0.28884/21,0.26977/22,0.25306/ & 23,0.2382/24,0.22516/25,0.2134/26,0.20281/32,-1PROCEDURE DATA PROCEDURE(KINETIC) ID=FTR_KINE, NAME=Kinetics of FTR CODE PCO = XVAP(4)*RPRES PH2 = XVAP(1)*RPRES T1 = EXP(-4492*(1/RTABS-1/473)) T2 = EXP(8237*(1/RTABS-1/473)) K1 = 0.0173*0.6*100**3/1000 K2 = 4.512 RRATES(1) = K1*T1*PH2*PCO/(1+K2*T2*PCO)**2 ISOLVE=1 RETURNUNIT OPERATIONS HX UID=DC_H_C COLD FEED=S83, M=S86 OPER DUTY=12.2 CONTROLLER UID=CN12 SPEC STREAM=S86, TEMPERATURE(F), VALUE=70 VARY STREAM=S83, RATE(KGM/H) CPARAMETER IPRINT, SOLVE HX UID=DC_L_C COLD FEED=S87, M=S88 OPER CDTADEW=1 CONTROLLER UID=CN13 SPEC HX=DC_L_C, DUTY(KJ/HR), VALUE=9.6 VARY STREAM=S87, RATE(KGM/H) CPARAMETER IPRINT, SOLVE PUMP UID=REF_1 FEED CW_1 PRODUCT M=CW_1_HP OPERATION EFF=77, PRESSURE=41.828 PUMP UID=REF_3 FEED CW_3 PRODUCT M=CW_3_LP OPERATION EFF=77, PRESSURE=2.3609 MIXER UID=M1 FEED METHANE,STEAM,CO2 PRODUCT M=S1 CONTROLLER UID=CN1 SPEC STREAM=S1,FRACTION, COMP=2,WET, DIVIDE, STREAM=S1, & FRACTION, COMP=5,7,WET, VALUE=0.5 VARY STREAM=STEAM, RATE(KGM/H) CPARAMETER IPRINT, SOLVE HX UID=SU_PREHEAT COLD FEED=S1, M=SYNGAS_FEED OPER CTEMP=537.78

CONREACTOR UID=OXIDATION FEED O2,SYNGAS_FEED PRODUCT M=S2 OPERATION ADIABATIC RXCALCULATION MODEL=STOIC, XOPTION=FAIL, REFS=IDEA RXSTOIC RXSET=P_OXIDATION REACTION OXIDATION BASE COMPONENT=3 CONVERSION 1 EQUREACTOR UID=SYNGAS FEED S2 PRODUCT M=SYNGAS OPERATION ADIABATIC, DP RXCALCULATION PARTIALPRESSURE, MODEL=METHANATION REACTION METHANATION REACTION SHIFT CONTROLLER UID=O2 SPEC STREAM=SYNGAS, TEMPERATURE(F), VALUE=1950 VARY STREAM=O2, RATE(KGM/H) CPARAMETER IPRINT, SOLVE HX UID=DC_H_REBOIL HOT FEED=SYNGAS, M=S4 DEFINE DUTY(KJ/HR) AS COLUMN=DC_HEAVY, DNAME(KJ/HR)=REBOILER HX UID=AFTERCOOL1 HOT FEED=S44, M=S17 UTILITY WATER, TIN=15.556, TEMPERATURE=37.778 CONFIGURE COUNTER OPER HTEMP=43.333 FLASH UID=F1 FEED S17 PRODUCT W=WATER_4, L=NAPHTHA_3, V=S76 ADIABATIC COMPRESSOR UID=LIGHT_2 FEED S76 PRODUCT V=S50 OPERATION CALCULATION=ASME, PRATIO=3, EFF=77 HX UID=AFTERCOOL2 HOT FEED=S50, M=S77 UTILITY WATER, TIN=15.556, TEMPERATURE=37.778 CONFIGURE COUNTER OPER HTEMP=43.333 FLASH UID=F2 FEED S77 PRODUCT L=S78, V=S79, W=WATER_5 ADIABATIC COMPRESSOR UID=LIGHT_3 FEED S79 PRODUCT V=S35 OPERATION CALCULATION=ASME, PRATIO=3, EFF=77 MIXER UID=M4 FEED LIGHTS_1,S35 PRODUCT M=S14 HX UID=PROP_VAP HOT FEED=S14, M=S15 COLD FEED=PROPANE, M=PROPANE_VAP CONFIGURE COUNTER OPER CDTADEW=1 CONTROLLER UID=CN14 SPEC HX=PROP_VAP, HTEM(C), VALUE=-30 VARY STREAM=PROPANE, RATE(KGM/H) CPARAMETER IPRINT, SOLVE FLASH UID=LIGHT_SEP FEED S15 PRODUCT V=TAIL_2, W=WATER_6, L=LIGHT_FEED ADIABATIC COLUMN UID=DC_LIGHT PARAMETER TRAY=15,IO FEED LIGHT_FEED,7, SEPARATE PRODUCT OVHD(M)=TAIL_1,100, WATER(M)=WATER_DIST_1,1, & BTMS(M)=S95, SUPERSEDE=ON CONDENSER TYPE=PART DUTY 1,1,,CONDENSER DUTY 2,15,,REBOILER PSPEC PTOP=23, DPCOLUMN=0.68046 PRINT PROPTABLE=PART ESTIMATE MODEL=SIMPLE, RRATIO=3, BTEMP=100 SPEC ID=COL4SPEC1, STREAM=TAIL_1, RATE(KGM/H), COMP=9,WET, & DIVIDE, STREAM=LIGHT_FEED, RATE(KGM/H), COMP=9,WET, & VALUE=0.05, ATOLER=0.05 SPEC ID=COL4SPEC2, STREAM=S95,FRACTION(WT), COMP=9,26,WET, & VALUE=0.99, ATOLER=0.05 VARY DNAME=CONDENSER,REBOILER REBOILER TYPE=KETTLE HX UID=DC_L_REBOIL HOT FEED=S4, M=S5 DEFINE DUTY(KJ/HR) AS COLUMN=DC_LIGHT, DNAME(KJ/HR)=REBOILER

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HX UID=FTR_HPSTM HOT FEED=S5, M=S6 COLD FEED=CW_1_HP, M=HPSTEAM_1 CONFIGURE COUNTER OPER CTEMP=254.44 CONTROLLER UID=CO2 SPEC STREAM=S6,FRACTION, COMP=1,WET, DIVIDE, STREAM=S6, & FRACTION, COMP=6,WET, VALUE=2 VARY STREAM=CO2, RATE(KGM/H) CPARAMETER IPRINT, SOLVE CONTROLLER UID=CN7 SPEC HX=FTR_HPSTM, HTEM(F), VALUE=510 VARY STREAM=CW_1, RATE(KGM/H) CPARAMETER IPRINT, SOLVE, ITER=100 HX UID=E1 HOT FEED=S6, M=S7 UTILITY WATER, TIN=15.556, TEMPERATURE=37.778 CONFIGURE COUNTER OPER HTEMP=43.333 FLASH UID=LPG_SEP FEED S95 PRODUCT V=S19, L=NAPHTHA_4 ADIABATIC PRESSURE=1 COMPRESSOR UID=PROP_C1 FEED PROPANE_VAP PRODUCT V=PROP_VAP_P1 OPERATION CALCULATION=ASME, PRATIO=3, EFF=77 COMPRESSOR UID=PROP_C2 FEED PROP_VAP_P1 PRODUCT V=PROP_VAP_P2 OPERATION CALCULATION=ASME, EFF=77 DEFINE PRAT AS COMPRESSOR=PROP_C1, PRAT HX UID=PROP_COND HOT FEED=PROP_VAP_P2, M=PROPANE_COND UTILITY WATER, TIN=15.556, TEMPERATURE=21.111 CONFIGURE COUNTER OPER HLFRAC=1 CONTROLLER UID=CN19 SPEC STREAM=PROPANE_COND, TEMPERATURE(F), VALUE=80 VARY COMPRESSOR=PROP_C1, PRAT CPARAMETER IPRINT, SOLVE VALVE UID=V4 FEED PROPANE_COND PRODUCT M=PROPANE OPERATION PRESSURE=1 HX UID=R134A_1_VAP HOT FEED=S7, M=S8 COLD FEED=R134A_1, M=R134A_1_VAP CONFIGURE COUNTER OPER CDTADEW=1 CONTROLLER UID=CN6 SPEC STREAM=S8, TEMPERATURE(C), VALUE=1 VARY STREAM=R134A_1, RATE(KGM/H) CPARAMETER IPRINT, SOLVE FLASH UID=H20_RMVL FEED S8 PRODUCT V=S9, W=WATER_1 ADIABATIC DP=0.068046 HX UID=FTR_PREHEAT HOT FEED=HPSTEAM_1, M=HPSTEAM1_SAT COLD FEED=S9, M=FTR_FEED CONFIGURE COUNTER OPER CTEMP=225.85 PLUG UID=FTR FEED FTR_FEED PRODUCT M=FTR_PROD OPERATION LENGTH=18.288, DIAMETER=31.75, TUBES=3.2253E5, & POINTS=50, THERMAL RXCALCULATION PARTIALPRES, RUNGEKUTTA, NSTEPS=50 PACK DPCORR=ERGUN CATALYST PDIAM=2.032, PORO=0.4 RXSTOIC RXSET=FTR REACTION R1 CONTROLLER UID=CN4 SPEC PLUGFLOW=FTR, CONVERSION(1), VALUE=0.95 VARY PLUGFLOW=FTR, TUBES CPARAMETER IPRINT, SOLVE HX UID=E2 HOT FEED=FTR_PROD, M=S10 COLD FEED=CW_2, M=S64 CONFIGURE COUNTER OPER HTEMP=200 CONTROLLER UID=CN10 SPEC STREAM=S64, TEMPERATURE(F), VALUE=110 VARY STREAM=CW_2, RATE(KGM/H) CPARAMETER IPRINT, SOLVE, ITER=50 FLASH UID=3P_1 FEED S10 PRODUCT L=HEAVIES, W=WATER_2, V=LIGHTS

ADIABATIC DP=1 FLASH UID=L_H_SPLT FEED HEAVIES PRODUCT V=NAPHTHA_1, L=HI_FEED ADIABATIC PRESSURE=1.2 CONTROLLER UID=CN3 SPEC STREAM=HI_FEED,PCT, COMP=5,WET, VALUE=0.1 VARY FLASH=L_H_SPLT, PRES(ATM), MINI=1 CPARAMETER IPRINT, SOLVE PUMP UID=HI_1 FEED HI_FEED PRODUCT M=HI_FEED_P OPERATION EFF=77, PRESSURE=27.5 CONREACTOR UID=HI FEED HI_FEED_P PRODUCT M=HI_PROD OPERATION ISOTHERMAL RXCALCULATION MODEL=STOIC, XOPTION=FAIL, REFS=IDEA RXSTOIC RXSET=HI REACTION C2125 BASE COMPONENT=28 CONVERSION 1 REACTION C2629 BASE COMPONENT=29 CONVERSION 1 REACTION C3035 BASE COMPONENT=30 CONVERSION 1 REACTION C3647 BASE COMPONENT=31 CONVERSION 1 REACTION C48 BASE COMPONENT=32 CONVERSION 1 HX UID=FTR_LPSTM HOT FEED=LIGHTS, M=S11 COLD FEED=CW_3_LP, M=LP_STEAM CONFIGURE COUNTER OPER CTEMP=126.67 CONTROLLER UID=CN9 SPEC HX=FTR_LPSTM, HTEM(F), VALUE=270 VARY STREAM=CW_3, RATE(KGM/H) CPARAMETER IPRINT, SOLVE HX UID=E3 HOT FEED=S11, M=S12 UTILITY WATER, TIN=15.556, TEMPERATURE=37.778 CONFIGURE COUNTER OPER HTEMP=43.333 HX UID=R134A_2_VAP HOT FEED=S12, M=S13 COLD FEED=R134A_2, M=R134A_2_VAP CONFIGURE COUNTER OPER CDTADEW=1 CONTROLLER UID=CN8 SPEC STREAM=S13, TEMPERATURE(C), VALUE=0 VARY STREAM=R134A_2, RATE(KGM/H) CPARAMETER IPRINT, SOLVE FLASH UID=3P_2 FEED S13 PRODUCT W=WATER_3, V=LIGHTS_1, L=NAPHTHA_2 ADIABATIC MIXER UID=M3 FEED NAPHTHA_2,HI_PROD,NAPHTHA_1 PRODUCT M=S26 VALVE UID=V1 FEED S26 PRODUCT M=HEAVY_FEED OPERATION PRESSURE=2 COLUMN UID=DC_HEAVY PARAMETER TRAY=25,IO FEED HEAVY_FEED,11 PRODUCT OVHD(M)=NAPHTHA,1000, WATER(M)=WATER_DIST_2,1, & BTMS(M)=DIESEL,800 CONDENSER TYPE=PART DUTY 1,1,,CONDENSER DUTY 2,25,,REBOILER PSPEC PTOP=1.75, DPCOLUMN=0.25 PRINT PROPTABLE=PART ESTIMATE MODEL=CONVENTIONAL, RRATIO=3 SPEC ID=COL2SPEC1, STREAM=NAPHTHA,PCT(WT), COMP=1,16,WET, & VALUE=98 SPEC ID=COL2SPEC2, STREAM=DIESEL,PCT(WT), COMP=17,32,WET, & VALUE=98 VARY DNAME=CONDENSER,REBOILER REBOILER TYPE=KETTLE HX UID=NAPHTHA_COOL HOT FEED=NAPHTHA, M=S34 COLD FEED=S68, M=S69 CONFIGURE COUNTER

97

OPER HTEMP=43.333 CONTROLLER UID=CN16 SPEC STREAM=S69, TEMPERATURE(F), VALUE=100 VARY STREAM=S68, RATE(KGM/H) CPARAMETER IPRINT, SOLVE MIXER UID=M5 FEED S34,NAPHTHA_3,S78 PRODUCT M=S56 FLASH UID=NAPH_SEP FEED S56,NAPHTHA_4 PRODUCT L=NAPHTHA_PROD, W=WATER_7, V=S37 ADIABATIC PRESSURE=1 COMPRESSOR UID=LIGHT_1 FEED S37 PRODUCT V=S44 OPERATION CALCULATION=ASME, PRATIO=3, EFF=77 COMPRESSOR UID=3PH_1 FEED R134A_2_VAP PRODUCT V=R134A_2_P OPERATION CALCULATION=ASME, PRES=7, EFF=77 HX UID=R134A_2_COND HOT FEED=R134A_2_P, M=R134A_2_COND UTILITY WATER, TIN=15.556, TEMPERATURE=21.111 CONFIGURE COUNTER OPER HLFRAC=1 CONTROLLER UID=CN5 SPEC STREAM=R134A_2_COND, TEMPERATURE(F), VALUE=80 VARY COMPRESSOR=3PH_1, PRES(ATM) CPARAMETER IPRINT, SOLVE VALVE UID=V3 FEED R134A_2_COND PRODUCT M=R134A_2 OPERATION PRESSURE=1 COMPRESSOR UID=H2O_1 FEED R134A_1_VAP PRODUCT V=R134A_VAP_P1 OPERATION CALCULATION=ASME, PRATIO=1.8, EFF=77 COMPRESSOR UID=H2O_2 FEED R134A_VAP_P1 PRODUCT V=R134A_VAP_P2 OPERATION CALCULATION=ASME, PRATIO=1.8, EFF=77 HX UID=R134A_1_COND HOT FEED=R134A_VAP_P2, M=R134A_1_COND UTILITY WATER, TIN=15.556, TEMPERATURE=21.111 CONFIGURE COUNTER OPER HLFRAC=1 VALVE UID=V2 FEED R134A_1_COND PRODUCT M=R134A_1 OPERATION PRESSURE=2 COMPRESSOR UID=LPG_1 FEED S19 PRODUCT V=S40 OPERATION CALCULATION=ASME, PRATIO=3.46, EFF=77 COMPRESSOR UID=LPG_2 FEED S40 PRODUCT V=S38 OPERATION CALCULATION=ASME, PRATIO=3.46, EFF=77 HX UID=E4 HOT FEED=S81, M=S84 COLD FEED=TAIL_2, M=S16 CONFIGURE COUNTER OPER CTEMP=10 CONTROLLER UID=CN2 SPEC STREAM=S84, TEMPERATURE(F), VALUE=60 VARY STREAM=S81, RATE(KGM/H) CPARAMETER IPRINT, SOLVE HX UID=LPG_COOL HOT FEED=S38, M=LPG_PROD COLD FEED=S72, M=S73 CONFIGURE COUNTER OPER HTEMP=43.333 CONTROLLER UID=CN18 SPEC STREAM=S73, TEMPERATURE(F), VALUE=100, ATOLER=1 VARY STREAM=S72, RATE(KGM/H) CPARAMETER IPRINT, SOLVE HX UID=DIESEL_COOL HOT FEED=DIESEL, M=DIESEL_PROD COLD FEED=S65, M=S67 CONFIGURE COUNTER OPER HTEMP=43.333 CONTROLLER UID=CN15 SPEC STREAM=S67, TEMPERATURE(F), VALUE=100 VARY STREAM=S65, RATE(KGM/H) CPARAMETER IPRINT, SOLVE CALCULATOR UID=CA1 RESULT 1,Production Rate DEFINE P(1) AS STREAM=LPG_PROD, RATE(LV,BBL/D),TOTAL,WET DEFINE P(2) AS STREAM=NAPHTHA_PROD, RATE(LV,BBL/D),TOTAL,WET

DEFINE P(3) AS STREAM=DIESEL_PROD, RATE(LV,BBL/D),TOTAL,WET PROCEDURE R(1) = P(1)+P(2)+P(3) RETURN SPLITTER UID=SP1 FEED S16 PRODUCT M=HP_STM_FUEL, M=SYN_FUEL OPERATION OPTION=FILL SPEC STREAM=SYN_FUEL, RATE(KGM/H),TOTAL,WET, VALUE=5000 CALCULATOR UID=CA2 RESULT 1,Preheater/2,Steam Generation DEFINE P(1) AS HX=SU_PREHEAT, DUTY(KJ/HR) DEFINE P(2) AS STREAM=SYN_FUEL, RATE(GV,M3/H), PHASE=V,TOTAL, & WET DEFINE P(3) AS STREAM=SYN_FUEL, LHV(KJ/M3) DEFINE P(4) AS STREAM=HP_STM_FUEL, RATE(GV,M3/H), PHASE=V, & TOTAL,WET DEFINE P(5) AS STREAM=S16, LHV(KJ/M3) PROCEDURE R(1) = P(1)*1E6/P(2)/P(3) R(2) = P(4)*P(5)*1E-6*0.85 RETURN CONTROLLER UID=CN17 SPEC CALCULATOR=CA2, R(1), VALUE=1 VARY SPLITTER=SP1, SPEC(1) CPARAMETER IPRINT, SOLVE MIXER UID=M2 FEED WATER_1,WATER_2 PRODUCT M=S74 PUMP UID=REF_2 FEED WATER_MAKEUP PRODUCT M=S92 OPERATION EFF=77, PRESSURE=20.34 MIXER UID=M6 FEED S74,S92 PRODUCT M=S82 HX UID=FTR_COOLING HOT FEED=S82, M=FTR_WATER OPER HDTBBUB=5 CONTROLLER UID=CN11 SPEC STREAM=FTR_WATER, RATE(WT,KG/H),TOTAL,WET, VALUE=1.9E6 VARY STREAM=WATER_MAKEUP, RATE(KGM/H) CPARAMETER IPRINT, SOLVE PUMP UID=P1 FEED S97 PRODUCT M=CW_HP OPERATION EFF=77, PRESSURE=41.828 HX UID=HP_STM_BOIL COLD FEED=CW_HP, M=HP_STEAM DEFINE DUTY(KJ/HR) AS CALCULATOR=CA2, R(2) CONTROLLER UID=CN20 SPEC HX=HP_STM_BOIL, CTEM(F), VALUE=490 VARY STREAM=S97, RATE(KGM/H) CPARAMETER IPRINT, SOLVE, ITER=100EN

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TABLE 35- CUMULATIVE CASH FLOW

Yearly Cash Flow ($1,000’s) DCFR Analysis

End of Year

Investment

Sales Revenue

Taxable Profit in Previous

YearTax Payable

Cash Flow After Taxes


Cash Flow Discounte

d


Discounted

Fixed Capital

Working Capital Total

Fixed Operating

Costs

Variable Operating

Costs

Total Operating Costs

1 354,182 - 354,182 - - - - - - (354,182) (354,182) (354,182) (354,182)

2 1,062,546

- 1,062,546 - - - - (354,182) - (1,062,546) (1,416,728) (985,590) (1,339,772)

3 2,125,091

219,654 2,344,745 - - - - (1,062,546) - (2,344,745) (3,761,473) (2,017,406)

(3,357,179)

4 - - - 477,561 595,939 1,073,499 895,114 (2,344,745) - 72,379 (3,689,094) 57,764 (3,299,414)

5 - - - 491,888 1,047,298 1,539,175 1,843,934 (178,386) - 555,524 (3,133,570) 411,243 (2,888,171)

6 - - - 506,644 1,078,706 1,585,350 1,899,252 304,759 100,571 464,096 (2,669,473) 318,678 (2,569,493)

7 - - - 521,843 1,111,067 1,632,910 1,956,230 313,902 103,588 470,496 (2,198,977) 299,674 (2,269,819)

8 - - - 537,499 1,144,399 1,681,898 2,014,916 323,319 106,695 477,088 (1,721,889) 281,865 (1,987,955)

9 - - - 553,624 1,178,731 1,732,355 2,075,364 333,019 109,896 483,878 (1,238,011) 265,171 (1,722,783)

10 - - - 570,232 1,214,093 1,784,325 2,137,625 343,009 113,193 490,871 (747,140) 249,521 (1,473,262)

11 - - - 587,339 1,250,516 1,837,855 2,201,754 353,300 116,589 498,075 (249,065) 234,846 (1,238,417)

12 - - - 604,959 1,288,031 1,892,991 2,267,806 363,899 120,087 505,464 256,429 221,082 (1,017,335)

13 - - - 623,108 1,326,672 1,949,781 2,335,840 374,815 123,689 513,136 769,564 208,170 (809,165)14 - - - 641,802 1,366,472 2,008,274 2,405,916 386,060 127,400 521,007 1,290,571 196,055 (613,110)15 - - - 661,055 1,407,466 2,068,522 2,478,093 397,642 131,222 529,114 1,819,685 184,686 (428,424)16 - - - 680,887 1,449,690 2,130,578 2,552,436 409,571 135,158 537,465 2,357,150 174,013 (254,411)17 - - - 701,314 1,493,181 2,194,495 2,629,009 421,858 139,213 546,066 2,903,216 163,993 (90,418)18 354,182 (219,654) 134,528 722,353 1,537,977 2,260,330 2,707,879 434,514 143,390 420,397 3,323,612 117,109 26,69119 - - - - - - - 313,021 103,297 (103,297) 3,220,315 (26,691) 0

Discount Rate: 7.8%

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Senior Design Project - Gas-to-Liquids Plant Design