Report No. 58-7 10,000 BPOD Fischer-Tropsch Synthesis Plant

		Page
Summary and Conclusions		1
Process Description		1
	Coal Preparation	2
	Gasification	2
	Dust Removal	3
	Purification	3
	Primary Synthesis	4
	Recycle CO₂ Scrubbing	4
	Secondary Synthesis	4
	Catalyst Preparation	5
	Refining	5
Capital Investment		7
Operating Costs		7
Appendix		8
	Table 1. - Total Estimated Capital Requirements	9
	Table 2. - General Plant Facilities Cost Summary	10
	Table 3. - General Utilities Cost Summary	11
	Table 4. - Estimated Working Capital	12
	Table 5. - Equipment Cost Summary - Coal Preparation	13
	Table 6. - " " " - Gasification	14
	Table 7. - " " " - Primary Waste Heat Recovery	15
	Table 8. - " " " - Dust Removal	16
	Table 9. - " " " - Purification	17
	Table 10. - " " " - Primary Synthesis	18
	Table 11. - " " " - Recycle CO₂ Scrubbing	19
	Table 12. - " " " - Secondary Synthesis	20
	Table 13. - " " " - Catalyst Preparation	21
	Table 14. - Plant Utility Requirements	22
	Table 15. - Estimated Annual Operating Cost	23
	Table 16. - Summary of Number of Employees and Payrolls by Cost Classification	24
	Table 17. - Summary of Number of Employees and Payrolls by Occupational Classification	25
	Table 18. - Detailed Equipment List - Coal Preparation	26
	Table 19. - " " " - Gasification	28
	Table 19. - " " " - Purification	29
	Table 19. - " " " - Synthesis Units	31
	Table 19. - " " " - Catalyst Preparation	32
	Figure 1. - Process Flowsheet	33
	Figure 2. - Coal Preparation Flowsheet	34
	Figure 3. - Gasification Flowsheet	35
	Figure 4. - Purification	36
	Figure 5. - Primary Synthesis Flowsheet	37
	Figure 6. - Recycle CO₂ Scrubbing	38
	Figure 7. - Secondary Synthesis Flowsheet	39
	Figure 8. - Catalyst Preparation Flowsheet	40
	Figure 9. - Steam and Power Balance	41
	Figure 10.- The Effect of Coal Cost on Production Cost	42
	Figure 11.- The Effect of Return on Investment on Gasoline Cost	43
Bibliography		44

A synthetic liquid fuels plant, designed to produce 10,000 barrels per operating day of liquid hydrocarbon product from the synthesis unit, will require an investment of approximately $114,356,500. Operating cost, with the cost of coal as a variable, are estimated to amount to the following:

Federal corporate income tax and a capital return are not included in the above costs. Incorporating these two items as a gross return on investment results in the following production costs:

1. The manufacture of purified synthesis gas with a desired 1:1 H₂:CO ratio for use in the subsequent step,

3. The refining of the hydrocarbons to produce gasoline, gas oil, LPG, and refinery off-gas.

Synthesis gas with an H₂:CO ratio of 1:1 is produced by gasifying coal with 95 percent oxygen and superheated steam. The synthesis gas is cooled in a waste heat recovery system after which the dust is removed by cyclones and electrostatic precipitators. The CO₂ content of the gas is then reduced to 0.1 percent by hot carbonate scrubbing. The hot carbonate system will also remove the greater portion of the sulfur content of the gas; however, iron oxide boxes are required for final sulfur clean up.

The conversion of synthesis gas to hydrocarbons is accomplished in fixed bed reactors which are filled with steel lathe-turnings catalyst. The synthesis gas is mixed with recycled oil which absorbs most of the heat of reaction. The oil is separated from the non-condensable gases, cooled in a waste heat recovery system, and then recycled to the synthesis reactors. The gases are cooled and further processed in the refining section. The process flowsheet is shown in figure 1, and detailed equipment lists are given in tables 18 through 22.

The following sections give a general description of the various processing operations.

The coal preparation facilities are designed to handle approximately 5,300 tons per day of run-of-mine coal. The coal as received has the following analysis:

The coal is conveyed from the track hopper to the hammer mills where the size is reduced to minus 3/4 inch. After leaving the hammer mills the coal passes over magnetic separators and then to storage hoppers or to the outdoor storage pile. The coal feeds from the storage hoppers to the ball mills where the size is further reduced to 70 percent through 200 mesh. The pulverized coal is conveyed by air to the storage bins and from there it is fed through a lock hopper and a Bailey feeder to the gasifier.

The gasification process is designed to produce synthesis gas, a mixture of carbon monoxide and hydrogen, with an H₂ : CO ratio of 1 : 1. The gasification reaction is carried out at approximately 435 p.s.i.g. so that no recompression is necessary prior to the synthesis reaction.

The feed to the three operating gasifiers consists of 190 tons of coal per hour, 332,000 pounds of superheated steam per hour, 135 tons of 95 percent purity oxygen per hour, and 10,000 pounds of oxygenates and water per hour. The exit gases leave the gasifier at 2,200º F. Approximately 50 percent of the ash and unreacted carbon is removed as slag through the slag accumulating system to a settling basin while the remainder is entrained in the raw synthesis gas and subsequently removed in a dust removal system.

A waste heat recovery system consisting of economizers, vaporizer, superheaters, and boiler feed water pumps is designed to cool the exit gases to 300º F. The heat recovery unit will produce 740,000 pounds per hour of steam at 460 p.s.i.g. and 1,000º F. A portion of this steam is recycled to the gasifier and the surplus is used in other sections of the plant.

The oxygen supply for gasification is provided by three 1100 ton-per-day oxygen plants. Steam driven compressors are used to compress the oxygen to 460 p.s.i.g.

The entrained ash and carbon must be removed from the gas stream prior to the synthesis stage. This is accomplished by cyclone separators and electrostatic precipitators operating in series.

The cyclones are designed on the basis of an allowable entrance velocity of 50 feet per second. The electrostatic precipitators are sized for a 5 second retention time with a linear velocity of 4 feet per second.

After the entrained ash has been removed from the gas stream it is necessary to remove the CO₂ and sulfur. Specifications for the synthesis reaction necessitate the reduction of the CO₂ content of the gas feed to 0.1 percent. Further, since the synthesis catalyst is rapidly poisoned by sulfur, it is essential that the sulfur be reduced to 0.1 grain per 100 cubic feet.

The CO₂ content of the gas stream is reduced to the required concentration by use of hot K₂CO₃ solution. The hot solution, approximately 194º F., is circulated countercurrently to the gas stream in packed towers. The fouled hot carbonate solution is regenarated with low pressure steam in packed towers and the sour gas is sent to the sulfur recovery unit.

Sulfur removal is accomplished primarily in the hot carbonate towers but the residual must be removed by passing the gas through iron oxide boxes. Facilities have been provided to revivify the iron oxide without removing it from the catch boxes.

The purified synthesis gas if fed to the fixed bed catalytic reactors where it is reacted over an oxidized iron catalyst. Catalyst preparation is outlined in a later section. By varying the temperature and pressure of operation it is possible to control, within limits, the type of product which will be formed. By operating the reactors at about 400 p.s.i.g. and an average temperature of about 525º F, the formation of gasoline range hydrocarbons is favored at the expense of oxygenated chemical synthesis.

The synthesis gas is fed at a space velocity of 500 V/V/H to the 9 foot diameter by 30 feet high reactors where 90 percent of the feed is gas is reacted. The pressure drop across the catalyst bed is 30 p.s.i. and an oil recirculating pump is required for each reactor.

The overhead stream from the reactors is separated into three streams, water and water soluble oxygenates, the light hydrocarbons and CO₂, and the heavy oil which is recirculated to the reactors. The water and oxygenates are recycled to the gasifier and replace part of the steam required for gasification. The hydrocarbon stream is scrubbed to remove CO₂ and then the liquid products are recovered as outlined in a later section on refining.

The hydrocarbon stream which contains approximately 11 percent CO₂ is scrubbed in hot carbonate scrubbers to reduce the CO₂ content to approximately 0.5 percent. The design and operation of this section is similar to the initial purification except that no sulfur removal equipment is required.

In order to obtain the maximum gasoline yield from coal a secondary synthesis unit has been designed. The reactors and auxiliary equipment of this unit are the same as the primary synthesis unit and may be used as primary synthesis reactors if required.

In the second-stage synthesis 50 percent conversion of the feed gas is achieved at a space velocity of 300 V/V/H. No recycle, and therefore, no intermediate CO₂ scrubbing is required. The product streams from this unit are added to the product streams from the first-stage synthesis unit.

The conversion of synthesis gas to hydrocarbons is greatly accelerated by the use of oxidized iron catalysts. An oxidized iron lathe-turnings catalyst has been developed by the Bruceton laboratory of the Bureau of Mines. This catalyst presents high active surface with fairly low pressure drop. Careful control of cutting speed and depth of cut in the lath operation results in a catalyst with about 83 percent to 85 percent void space. The turnings will be cut from S.A.E. 1018 steel on 40 standard lathes. The operation of these lathes would be prorammed so that an operator would not be required for each lathe.

A batch oxidation process is then carried out in a stainless steel lined reactor where 20 percent of the iron is oxidized by superheated steam. The steam is fed to the reactor at 1,110º F. and a space velocity of 400 V/V/H. After an oxidation period of approximately 40 hours the batch is cooled and impregnated with a solution of potassium carbonate. After drying to remove the water the oxidized catalyst is transferred to the synthesis reactors.

The final reduction step, with hydrogen, is done in the spare synthesis reactors. The hydrogen is fed to the reactors at 840º F. and a space velocity of 2,000 V/V/H. The iron oxide is reduced to iron and water is produced. Since it is necessary to remove this water, the hydrogen recycle stream is cooled, dehydrated with silica gel, and reheated before re-entering the reactor. The reduction cycle is approximately 48 hours.

The refinery is designed to refine a high gravity crude for maximum gasoline yield.

The purified synthesis gas is fed to the fixed bed catalytic reactors where it is reacted over an oxidized iron catalyst. Catalyst preparation is outlined in a later section. By varying the temperature and pressure of operation it is possible to control, within limits, the type of product which will be formed. By operating the reactors at about 400 p.s.i.g. and an average temperature of about 525º F., the formation of gasoline range hydrocarbons is favored at the expense of oxygenated chemical synthesis.

The synthesis gas is fed at a space velocity of 500 V/V/H to the 9 foot diameter by 30 feet high reactors where 90 percent of the feed gas is reacted. The pressure drop across the catalyst bed is 30 p.s.i. and an oil recirculating pump is required for each reactor.

In the second-stage synthesis 50 percent conversion of the feed gas is achieved at a space velocity of 300 V/V/H. No recycle, duct streams from this unit are added to the product streams from the first-stage synthesis unit.

The conversion of synthesis gas to hydrocarbons is greatly accelerated by the use of oxidized iron catalysts. An oxidized iron lathe-turnings catalyst has been developed by the Bruceton laboratory of the Bureau of Mines. This catalyst presents high active surface with fairly low pressure drop. Careful control of cutting speed and depth of cut in the lathe operation results in a catalyst with about 83 percent to 85 percent void space. The turnings will be cut from S.A.E. 1018 steel on 40 standard lathes. The operation of these lathes would be programmed so that an operator would not be required for each lathe.

The refinery is designed to refine a high gravity crude for maximum gasoline yield.

The light ends recovery unit is designed to accumulate and compress the light overhead streams and to separate and recycle them to the polymerization or reforming units. The off-gas, principally methane, is distributed to the plant for process heating.

The catalytic cracking unit is designed to process 3,835 barrels per operating day of heavy oil and wax product from the synthesis unit and heavy polymers from the polymerization and reforming units. With reactor operating conditions of 900º F. and 12 p.s.i.g. the following yield structure has been estimated:

The catalytic polymerization plant will use the U. O. P. solid phosphoric acid catalyst type process. Olefines in the C₃ - C₄ range are converted to gasoline range material by contacting them with the catalyst at 425º F. and 1,100 p.s.i.g. At these operating conditions 90 percent olefin conversion can be accomplished if the olefin concentration in the feed stream is maintained at 40 percent (by volume). This is done by recycling a portion of the propane from the top of the depropanizer back to the polymerization unit.

The straight-run gasoline boiling range material will be formed over a platinum catalyst using the "Ultra-forming" process. Use of this process is well suited to a low sulfur feed and will give a maximum yield of high octane reformate. The following yield pattern had been estimated:

The refinery product has high lead susceptibility and the addition of 3 cc. of TEL will give a product with an octane rating comparable to premium gasoline. Sufficient butane will be recovered for blending to 10 pounds Reid vapor pressure.

Tankage facilities have been designed to insure adequate storage for the finished product and all intermediate streams.

As shown in table 1 the total capital investment, including working capital, is $114,356,500. The capital investment costs include allowances for spare parts. The capital investment costs include allowances for spare parts. The capital requirements for each operating unit are given in tables 5 through 13. The cost of the oxygen plant and the sulfur recovery plant were estimated from available quotations.

The power plant cost was estimated on the basis of the steam and power balance, figure 9. A spare boiler and turbogenerator were provided. The cost of individual items for plant facilities and plant utilities are given in tables 2 and 3.

Table 1 shows a total construction cost of $101,006,500. The capital investment for initial catalyst and chemical requirements is $1,604,900. Interest during construction, calculated on the basis of a two year construction period, is $2,692,900, is estimated as shown in table 4.

Table 15 shows the estimated annual operating cost summary. It has been assumed that this plant will have a 90 percent operating factor. This 35 day down time allowance assumes two 10 day shutdowns for equipment overhaul and 15 days for unscheduled interruptions.

Personnel requirements have been estimated using as a base the synthetic fuel plant operation at Brownsville, Texas. They are summarized in the two manning tables, table 16 and 17.

The operating cost with coal at $4 per ton is shown in table 15 as $26,656,700. The unit operating cost is calculated by considering the gas oil to have 70 percent of the value of gasoline. Credit is taken for propane at the current market price. The resulting operating cost is $0.230 per gallon of premium gasoline. Gross return on the investment is an important variable which exerts a large effect on the cost of the finished product. The effect of this variable is shown in figure 11.

The principal item of operating cost is the cost of coal. Coal at $3 per ton represents 20 percent of the annual operating cost; at $4 per ton coal is 26 percent of the operating cost, and at $5 per ton coal is 30 percent of the operating cost. Figure 10 shows the effect of coal price on the operating cost.

Coal cost/ton	Operating cost/gallon
$3.00	$0.215
4.00	.230
5.00	.244

	Production cost/gallon
Gross return on investment, %	Coal at $3/ ton	Coal at $4/ton	Coal at $5/ton
6	$0.274	$0.289	$0.303
12	.333	.348	.362
20	.412	.426	.441
30	.510	.525	.540

H₂	4.9%
C	70.1
S	1.4
N₂	1.4
O₂	6.6
Ash	14.3
Moisture

H₂O	29.6%
H₂	29.7
CO	29.7
CO₂	9.4
N₂	1.3
S	0.3

H₂O	0.2%
H₂	48.8
CO	48.8
CO₂	0.1
N₂	2.1

C₁	1.0%	C₄	3.1%
C₂	1.1	C₅⁺	37.1
C₃	3.2	Gas oil	40.1
C₃	1.6	Coke	6.0
C₄	4.8	Water	2.0

H₂	2.0%	C₄	5.4
C₁	1.5	C₅	79.1
C₂	3.5	Heavy
C₃	5.3	Polymer	3.2