Ind. Eng. Chem. Prod.
Res. Dev.
1980, 19, 175-191
175
Fischer-Tropsch Processes Investigated at the Pittsburgh Energy Technology Center since 1944 Michael J. Balrd,' Richard R. Schehl, and William P. Haynes Uniied States Department of Energy, Pinsburgh Energy Technology Center, Pittsburgh, Pennsylvania 152 13
James T. Cobb, Jr. Department of Chemical and Petroleum Engineering, Universify of Pittsburgh, Pittsburgh, Pennsylvania 15261
Fischer-Tropsch investigations conducted by the Pittsburgh Energy Technology Center of the US. Department of Energy since 1944 are discussed. This review article concentrates on iron catalysts that were used in the oil circulation, slurry, fiuidized-bed. and hot-gas-recycle reactors and most recently in the tube wall reactor. Following a brief historical background of the Fischer-Tropsch process, each reactor system is described and some of the important experimental results are presented. A thorough reference list of government publications on Fischer-Tropsch studies in pilot plant reactors is included.
Introduction With the current energy problems facing the United States, the Fiscber-Tropsch (F-T) synthesis is surely a contending process to combine with coal gasification in order to produce synthetic fuels. The process was used by the Germans during World War I1 and is currently being used in South Africa to supply that country with gaseous and liquid fuels and chemicals. There are nu-
Michael J Baird was a Si sory Chemical Engineer I Pittsburgh Energy Tech1 Center from 1974 until 1979 he was a project leader for I nation catalyst deuelopmei Fischer-Tropsch catalyst e tion. He receiued a B.S. in ical enpineerine and a Ph physicel chemTstry (19713 rrom Texas A&M Uniuersity and an M.S. in chemical enzineerinz (1979) from the Uniuersitv of Pittsburgh. He r&entlyyo'ined'Ashland Oil, Inc., ai a Senior Research Chemist in their Petroleum Catalyst Group. R i c h a r d R Schehl is Assis Manager, Process Sciences 1 sion, in the Pittsburgh E n Technology Center, Pittsbu Pa. He received the B.S. de from Ohio Uniuersity and Ph.D. degree from The Ohio S Uniuersity. For the past ten y he has been engaged in r e x inuoluing catalyst deuelopn and reactor engineering relateu bv methanation and Fischer-Tropsch synthesis.
* Ashland Oil, Inc., Petroleum Research (R&E Building), Catlettsburg, Ky. 41129.
merous review articles and books ( I , 61, 62, 70, 74) that describe various aspects of the F-T process. Recent review articles by Henrici-Olive and Olive (45), Vannice (75),Dry (19),Ponec (641, and Shah and Perrotta (72) have emphasized proposed mechanisms, kinetics, and types of catalysts and promoters used in the reaction. Fischel-Tropsch research was being conducted by the U S . government prior to World War 11. The federal F-T research program was mainly concerned with developing and testing catalysts in narrow-tube, fixed-bed reactors. However, investigations were accelerated when Congress passed the Synthetic Liquid Fuels Act on Apr 5, 1944. This legislation directed the U S . Bureau of Mines to explore and develop methods for converting to oil such materials as coal, oil shale, and agricultural and forestry products. T o obtain the operational parameters required by American industry for commercial operations, the hureau was authorized t o go beyond the existing research efforts and build and operate pilot and demonstration plants. This review article describes the different F-T William I? Haynes is currently Senior Staff Engineer of the Pittsburgh Energy Technology Center, U.S. Department of Energy. He receiued the B.S. and M.S. degrees in Chemical Engineering from the Uniuersity of Pittsburgh. He has been with the Center since 1945. During his last 3 years as a Diuision Manager he has been inuolued in rPsearch and deuelooment af processes for the conuerscon of coal to synthetic natural gas and liquid hydrocarbons. J a m e s T Cobb, Jr.,joined the Pitt faculty in 1970, following three years at the Esso Research Laboratories in Baton Rouge. He has directed research in the use of enzymes, zeolite catalysts, and coal gasifier modeling. During the summers of 1976-78 he was an Oak Ridge Associated Uniuersities Research Participant at the Pittsburgh Energy Technology Center. Actiue in numerous technical and professional societies, Dr. Cobb was president of the Pittsburgh Catalysis Society for 1978-79 and is chairman of the Professional Legislation Committee of the American Institute of Chemical Engineers.
This article not subject to U.S. Copyright. Published 1980 by the American Chemical Society
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Ind. Eng. Chem. Prod. Res. Dev., Vol. 19, No. 2, 1980
reactor systems that have been investigated a t the Pittsburgh Energy Technology Center (PETC) since 1944. Although several types of catalysts have been tested, this paper will concentrate on iron catalysts that have been used in the oil circulation, slurry, fluidized-bed, and hotgas-recycle (HGR) reactors and most recently in the tube wall reactor (TWR). Following a brief historical background of the Fischer-Tropseh process, each reactor system will be described and some of the important experimental results will be presented.
Historical Background The synthesis of liquid hydrocarbons via catalytic conversion of carbon monoxide and hydrogen was first disclosed by Badische Aniline und Soda-Fabrik (6, 57) in 1916. Research based on the reported German patents led to the development of the methanol synthesis in 1922. In 1925 Fischer and Tropsch (31) developed the normalpressure (1atm) process for the conversion of CO and H2 to liquid hydrocarbons. Following fundamental research by Fischer and co-workers (32)and by other laboratory studies in England (22,23), France ( 5 ) ,Japan (51),and the United States ( 4 4 , 731, Ruhrchemie constructed an atmospheric pressure, fixed-bed pilot plant in Germany in 1933. In the following year, four full scale plants (55) were erected. In 1936 Fischer and Pichler (29)discovered an optimum pressure range for carrying out the carbon monoxide hydrogenation in the presence of a cobaltkieselguhr catalyst. At approximately 5-20 atm (medium pressure) the yields of hydrocarbons were 10-15 % higher than with the normal pressure synthesis, and the yield of paraffin wax could be increased from 10 to 45% of the C5+ hydrocarbons. Both processes were used in nine Fischer-Tropsch plants in Germany during World War 11. The yearly capacity of the German plants in 1943 was about 6 000000 barrels of gasoline, diesel oil, lubricating oil, and paraffin wax. Fischer-Tropsch processes were also operated in France in 1937-one plant producing 1000 bbl/day, and in Japan in 1938-two plants with a total production of 2500 bbl/ day. Following World War 11, operation of the German plants was terminated, and F-T research and development was continued in the United States by petroleum companies and by the U.S. Bureau of Mines. The first commercial American synthetic fuels plant was constructed in 1951 a t Brownsville, Texas, by the Carthage Hydrocol Company, a consortium of nine U S . oil companies (62). The plant utilized a fluidized bed of iron catalyst to convert CO and H2, produced from natural gas, to approximately 7000 bbl/day of liquid hydrocarbons and chemicals. After several years of experimentation it was concluded that the process could not produce gasoline and chemicals as cheaply as competitive processes that used oil as feed, and development work was terminated in 1957. A major effort in improving the Fischer-Tropsch process was conducted by the U S . Bureau of Mines Pittsburgh Energy Research Center (PERC), now part of the U S . Department of Energy and recently renamed the Pittsburgh Energy Technology Center (PETC). At PETC emphasis was placed on reactor design where from 1944 to 1949 bench-scale and pilot plant experimentations were carried out on the oil circulation (9,10. 11, 14, 15, 16,26) and the slurry reactor (68,69) systems. In 1949, a coalto-oil demonstration plant utilizing these two reactor systems was built in Louisiana, Mo. (49). The improved American reactors were each designed to produce approximately 500 bbl/day of liquid fuel-a considerable improvement over the German fixed bed reactors that had
a capacity of 18 bbl/day per reactor (10). New discoveries of oil in the US.,Canada, and the Middle East decreased the U S . government's interest in synthetic fuels via coal gasification, and consequently operation of the demonstration plant in Missouri ended in 1953. Although large-scale Fischer-Tropsch investigations were not being conducted by the U.S. government after 1953, pilot plant research was continued at PETC with emphasis being placed on improving product distribution (67,68),catalysts (2,3, 4 , 7, 8, 50, 71),and reactor design (21,27, 36).
A. Oil Circulation Process In the oil circulation process, CO, H2,and recycle oil are passed over an iron catalyst to produce hydrocarbons ranging from methane to high molecular weight solids. The exothermic heat of reaction is removed by recycle oil which is externally cooled and then returned to the reactor. This process was first investigated by Duftschmid et al. in 1934 (20),and the U S . Bureau of Mines initiated work on the oil circulation process in 1943 (10). The U S . process used a higher temperature boiling oil than was employed by the Germans. This resulted in the removal of the heat of reaction largely as sensible heat rather than by vaporization. This mode of operation improved heat recovery and made the synthesis temperature and pressure independent of each other, thus allowing the reactor to operate over an increased temperature range. In early experiments a t the Bureau of Mines, a flow of oil was trickled through the catalyst bed, but in later experiments smoother operation and better temperature control were obtained by submerging the catalyst completely in oil (10). However, in submerged-bed operation the precipitated-iron catalysts disintegrated too rapidly and had to be replaced by a fused iron (synthetic ammonia) catalyst that had greater physical strength. Unfortunately, a gradual increase in pressure drop across the fixed bed occurred due to catalyst agglomeration. This led to a further modification in which the velocity of the cooling oil, flowing upward through the reactor, was increased sufficiently to expand the catalyst bed about 5 to 30% above its settled height-"expanded-bed operation" (16). Not only was agglomeration avoided, but the operating temperature was reduced 10 to 15 "C by using smaller catalyst particles with higher surface areas. Expanded-bed tests were also conducted using steel lathe turnings, steel shot, and nitrided iron catalysts. Operation. Three-pilot plant units were used to investigate the operability of the oil circulation process. Two were 3-gal/day units; one was a 1-bbl/day plant. The flow diagrams of the three units were basically the same (Figure 1). Fresh synthesis gas, H2 + CO, was obtained by reforming methane and was stored in large gas holders (not shown in Figure 1). The reactor for each 3 gal/day unit was a 3-in. schedule 80 carbon steel pipe 10.5 ft in length, and the reactor for the 1-bbl/day unit was 8 in. in diameter and 26.25 f t in length. The catalyst was first reduced in H2 a t 400 "C and 1 atm in a separate unit and then transferred under C 0 2 to the oil circulation reactor. The catalyst was submerged in oil, and the unit was pressurized to 300 psig. A perforated plate covered with a fine mesh screen was used to support the catalyst charge in the reactor. The importance of bringing a fresh catalyst on stream properly (induction or precarbiding) was stressed by German investigators (66). To a large extent, the catalyst activity and life are set by the method of induction. For the oil circulation tests induction was carried out in a
Ind. Eng. Chem. Prod. Res. Dev., Vol. 19. No. 2 , 1980 Water condenser
Air-cooled trap
Refrigerated condenser
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Pressure regulator
Charcoal spirits
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Charcoal-, adsorbed I aas ,
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Figure 1. Oil circulation process. Table I. Effect of Hydrogen-to-Carbon Monoxide Gas Ratio o n Conversion and Product Distribution in t h e Oil Circulation Process Using a Fused Iron Catalvst ( 1 1 1 expt no.
H, :CO feed ratio catalyst age, h space velocity, h" gas recycle ratio temperature, "C pressure, psig, H, + CO coriversion, % specific yield, C , + C,d product distribution, wt % gasoline (C,-204 "C) diesel (204-316 'C) heavy distillate (316-450 "C) wax ( > 450 " C )
22B0sb
22Ba
26Ab
31Cc
30BC
1:1 458 602 1:l 244 300 67 36.1
1:l 506 604 236 400 66 32.4
0.7 :1 309 59 9 1:l 242 40 6 67 24.8
1:l 624 600 1.3:l 251 400 69 29.6
0.7:l 510 600 1:l 251 450 69 20.2
58.6 12.5 19.3 9.6
57.5 13.2 15.5 13.8
43.2 15.6 19.9 21.3
64.6 14.0 11.1 10.3
41.7 17.0 15.3 26.0
1:l
Experiments 22 and 26 were conducted in the 3 gal/day unit. CO, riot removed from recycle stream. 30 and 3 1 were performed in the 1 bbl/day unit. Specific yield as g/m3 feed converted.
stepwise manner. First synthesis gas and oil entered the bottom of the reactor at 300 psig, and then reactor temperature was increased and maintained for 24 h at settings to give CO conversions of 30,45, and 60%. Following the 72-h induction period, temperature was increased until the CO conversion was increased to approximately 70%-a level that was maintained throughout the initial tests. In later experiments conditions such as temperature, pressure, synthesis gas ratio, arid gas hourly space velocity (GHSV) were varied during the run. For runs of this type, the length of time in which the operating conditions were held constant was defined as a "period". Throughout this review periods will be expressed either numerically or alphabetically, depending on how they were reported in a particular reference. During operation, cold recycle gas was mixed with fresh synthesis gas and preheated to reaction temperature. This mixture, as well as preheated recycle oil, entered the bottom of the reactor. Gaseous products were separated from the recycle oil and were passed through condensers for removal of the light hydrocarbon fraction. A bauxite reforming unit on the 3 gal/day unit was available for dehydrating oxygenated compounds. Since high molecular weight products condensed in the recycle oil, heavy oil was drained from the recycle line in order to maintain a con-
Experiments
stant level in the reactor. The effluent gas leaving the refrigerated condenser was divided into two streams. One portion was reduced to atmospheric pressure and passed to a charcoal recovery system that removed C3 and heavier gaseous hydrocarbons. The other portion, used for recycle, was passed through a monoethanolamine (MEA) system for C 0 2 removal, then mixed with fresh synthesis gas. Samples analyzed were liquid products after separation into aqueous and oil fractions and feed, recycle, and product gas streams. Fused Iron Catalyst. Once it was established that the moving-bed or expanded-bed mode of operation was superior to the fixed-bed system (161,the effect that operating parameters have on conversion and product distributions was investigated ( I I ) . The commercial catalyst used had the following composition (wt %): Fe304,94.51; MgO, 4.61; Cr203,0.65; SiO,, 0.64; KzO, 0.56; and Mn,O,, 0.03. For a CO rich feed gas (Table I) a higher yield of liquid products was formed. The raw gasoline produced with the 0.7:1 feed was more olefinic than that with the 1:l feed. But the higher partial pressure of CO accelerated disintegration of the catalyst which is in agreement with information reported by Storch et al. ( 7 4 ) and by Pichler and Merkel (63). Increasing the pressure in the %gal/day units resulted in a lower operating temperature for the
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Table 11. Effect of Temperature on Product Distribution in the Oil Circulation Processa Using a Fused Iron Catalyst ( I 1 ) temp, "C 242 C, + C, yield, g / m 3 feed converted product distribution, w t ?b gasoline (C,-204 "C) dksel (204-316°C) heavy distillate (316-450 "C) wax(>4SO°C)
248
256
24.8 28.1 26.4 43.2 15.6 19.9 21.3
48.8 14.4 18.4 18.4
57.2 15.1 15.7 12.0
Experiment 26, 0 . 7 : l H, t o CO feed gas, 400 psig, 600 h - ' GHSV.
same gas conversion (Table I). Also, at the higher pressure the production of higher boiling hydrocarbons increased. Similar effects of pressure on activity and product selectivity were observed by Anderson and co-workers (2) for fixed bed operation a t 100 and 300 psig and by Hall et al. (39) for slurry reactors at 300 and 675 psig. Increasing temperature (Table 11) caused a shift in product yield toward the production of lower boiling hydrocarbons. The greatest change was in the gasoline and wax fractions, while the production of the middle cuts (204 to 450 "C) was not appreciably affected by temperature. Increasing the temperature resulted in an increase in CO conversion up to a limit of 85% if C O Pwas not removed from the recycle gas. This conversion limit was raised above 90% by C 0 2 removal from the recycle stream, which lowered the content of oxidizing gases in contact with the catalyst, thereby decreasing the tendency of the iron to oxidize. Operation a t temperature greater than 270 "C caused spalling and disintegration of the catalyst particles. Massive Iron Catalysts. In order to increase the range of operating conditions and catalyst life, the use of massive iron catalysts was initiated (15). Iron in the form of 1 / 3 1 and 1/16-in.diameter steel shot and lathe turnings of commercially pure iron and carbon steel were used. The materials were oxidized with steam a t atmospheric pressure, 600 "C, and a GHSV of 400. The turnings were 20% oxidized, and the steel shot was 5 to 8% oxidized. The oxidized steel (Fe104)was impregnated with alkali by immersing the turnings or shot in a 2% K2C03solution for 5 min. These were then dried at 110 "C. The catalyst containing 0.02% K 2 0 was reduced with H2 at 2000 h-', 450 "C, and then stored under COz. Operation in the 3-gal/day unit was similar to the procedure described for the synthetic ammonia catalyst. Each catalyst was charged to the reactor to a height of 4 to 6 ft, submerged in oil, and pressurized to 300 psig. With the shot, it was necessary to expand the bed 5 to 10% to avoid plugging and the resulting high pressure drops across the catalyst bed. Because of the high percentage of voids in the lathe turnings. no bed expansion was necessary. There was a slight modification in the induction procedure. The oil temperature was rapidly raised to 200 "C and then it was gradually raised (2 "C/h) until a 70% CO + H2 conversion or a maximum temperature of 300 " C was reached. In experiment 32 119, using 1/16-in.steel shot as catalyst, a total life of 6 months was obtained. To achieve the desired level of H2 + CO conversion, temperatures as high as 300 "C were required. The run operated for 73 days before the catalyst had to be regenerated. Several reoxidation-reduction regenerations were required in order to obtain 6 months of operation. When employing lathe turnings catalyst life and activity were improved. In experiment 37 (15), 113 days of continuous synthesis was obtained without regeneration. By being able to operate the high void lathe turnings as a
fixed-bed, movement of the catalyst was avoided, and the active surface was subjected to less abrasive action. Operating data are presented in Table 111. Although the catalyst was impregnated with potassium prior to synthesis, a second alkali addition was given after 900 h of operation which caused significant changes to the product distribution. The trend observed after alkali addition was: (a) a decrease in C,+ C,; (b) an increase in the olefinic content of the C3 to C6hydrocarbons; (c) a decrease in the gasoline fraction, and (d) an increase in the heavy molecular weight hydrocarbons. In general, a higher yield of low-boiling hydrocarbons was produced with the lathe turnings than was produced from the fused iron catalysts. As the wax concentration was allowed to build up in the circulating oil by removing more of the reflux oil a t the overhead condenser, a greater yield of gasoline and a corresponding smaller yield of heavier hydrocarbons resulted. Apparently, the wax cracked to lighter hydrocarbons while being retained in the reactor. When the wax concentration was restored to its initial level, the gasoline and wax yields also returned to their original level. Nitrided Fused Iron. Bench-scale investigations by Anderson et al. (3,72) had demonstrated that nitrided iron was a more active F-T catalyst than conventionally reduced iron and that the yields of oxygenates were greatly increased. Since it was of interest to produce a gasoline with a high alcohol content, the performance of nitrided iron in pilot plant units was evaluated. The objectives were to study the activity and durability of the catalyst and to determine the yield of synthetic gasoline rich in alcohols. Magnetite (6-1 0 mesh) was reduced in H2as previously described and transferred under an inert atmosphere to the nitriding unit- a horizontal reactor containing iron in a shallow bed. Anhydrous ammonia gas was passed through the catalyst a t 1 atm, 500 h-', and 300 to 326 "C for approximately 30 h. A batch of 30 lb of catalyst was charged to the reactor of the 3-gal/day unit, and during the experiment operating conditions were varied to determine their effects upon yields of oxygenates and gasoline and synthesis gas conversion. The results from experiment 41 (14). which was the longest run in the oil circulation process without reactivation, are presented in Table IV. Of the variables studied, reflux had the greatest effect. When reflux of oil to the reactor was eliminated, the total yield of oxygenates increased (periods 4 vs. 6, 2 vs. 7, and 17 vs. 16). Longer residence times of alcohols in the reactor increased their probability to undergo dehydration to olefins. Oil fractions rich in alcohol had a low olefin content, while those with little alcohol had a high olefin content. In the range of 200 t o 240 "C there was a slight effect of temperature on oxygenate production. Lower temperatures favored a higher oxygenate yield (periods 7 , 8, and 101, but the yield advantage gained by low temperature operation was offset by low synthesis gas conversion. Operation at temperatures above 240 "C was also undesirable because the specific yield of oxygenates decreased considerably. Increasing reactor pressure from 300 to 400 psig resulted in increased conversion levels (periods 10 vs. 11) and increases in the production of oxygenates (periods 8 vs. 14, 6 vs. 16, and 4 vs. 17). At higher gas recycle rates, the more volatile oxygenates were stripped from the recycle oil, and the yield of oxygenates in the light-oil phase, as well as the total amount of oxygenates, increased. The effects of space velocity were similar to those for gas-recycle. Conclusions from Oil Circulation Process 1. Expanded- (or moving-) bed operation minimized the
Ind. Eng. Chem. Prod. Res. Dev., Vol. 19, No. 2, 1980
179
Table 111. Operating Data for Steel Lathe Turning Catalyst in the Oil Circulation Process, Experiment 37 ( 1 5 ) period catalyst age, h fresh feed space velocity, h-' temperature, 'C gas recycle ratio usage ratio, H, :CO H, + CO conversion, % C, t C, yield, g / m 3 feed conv. product dlstribution, wt % gasoline (C,-204 "C) diesel (204-316 "C) heavy distillate ( 3 16-450 "C) wax ( > 4 5 0 "C)
3
5
I
8
11
14
11
327 401 210
717 601 283
927 a 5 99 210 1:l 0.88 68.5 39.0
1281 614 2 80 2:l 0.82 10.8 33.6
1521 7 00 290 2:l 0.85 70.0 37.5
2390 40 1 290 1.5:1 0.82 19.4 36.9
55.0 17.4 13.6 14.0
57.5 14.6 13.3 14.6
72.3 16.3 7.6 3.8
10.2 13.3 10.4 6.1
1:l
1:l
0.89 70.5 39.1
0.90 71.6 58.3
819 600 27 0 1:l 0.95 72.9 44.2
67.1 15.5 9.5 7.9
85.3 8.2 4.5 2.0
74.0 14.5 8.5 3.0
-
a Alkali added prior to this period. Five grams of KOH in an alcoholic solution was injected into the bottom of the reactor. The average alkali content was increased to 0.05% K,O. Table IV. Conversions and Product Yields from Oil Circulation Operation Using a Nitrided Iron Catalyst" ( 1 4 )
period ___
catalyst age, h space velocity, h-' temperature, " C 6 temperature pressure, psig reflux CO conversion, % H, + CO conversion, % material balance, w t % hydrocarbon produci; distribution, wt %
c,
i.C,b
c 3
C,=-c, hydrocarbons conciensed oxygenates in oil oxygenates in water total product distribiition, wt % hydrocarbons H 20
co,
2
4
6
I
370 200 230 2 302 yes 72.2 64.7 99.0
905 301 240 3 301 yes 75.9 68.4 99.1
1385 300 240 3 299 no 76.1 68.4 98.1
1577 200 230 3 299 no 78.5 68.1 98.5
31.6 5.0 30.2 12.2 2.1 18.8
32.0 5.6 21.3 24.1 4.5 12.1
32.3 7.1 24.3 11.6 7.1 17.0
23.3 4.5 72.2
26.5 4.7 68.8
25.8 4.2 69.9
11
14
16
17
1691 2029 200 201 220 210 2 2 299 300 no no 63.5 48.1 54.8 42.5 98.8 97.7
2125 201 210 3 396 no 60.8 52.2 95.8
2688 201 220 2 397 no 53.5 46.3 96.8
3048 300 240 2 398 no 66.1 59.7 97.5
3191 299 240 2 397 yes 69.6 62.1 95.7
29.7 6.0 23.7 16.7 3.0 20.9
24.5 3.4 20.1 28.3 7.9 15.7
22.2 4.3 20.4 26.8 10.8 15.5
22.6 5.9 17.6 27.9 7.3 18.8
22.8 5.0 16.4 24.9 6.9 24.0
27.3 4.7 19.0 21.3 9.5 18.2
28.8 5.3 20.9 25.7 3.8 15.6
26.2 2.2 71.7
27.2 2.1 10.1
21.0 3.8 69.2
24.1 2.0 73.9
25.2 2.3 72.5
26.6 4.3 69.1
25.9 4.1 69.4
Recycleifresh gas ratio was 1:l and fresh gas ratio, H,/CO, was 1:l.
pressure drop probleins associated with fixed- (or submerged-) beds of precipitated and fused iron catalysts. However. general disintegration of these materials limited catalyst life in expanded-bed operation. 2. Good catalyst lire and activity were obtained from steel lathe turnings used in the fixed-bed mode of operation. Alkali addition to the catalyst decreased the C1 + C2 yield and increased the olefinic content of C3 to C6 hydrocarbons. 3. A process variable study indicated the following: (a) A CO rich feed gas increased the yield of liquid hydrocarbons and olefins but accelerated catalyst disintegration. (b) Higher pressure favored the production of higher boiling liquids. (c) Increasing the temperature shifts the hydrocarbon product distribution toward lower molecular weight products (C, -k CJ. 4. The longest run without catalyst regeneration (3695 hi was obtained with a nitrided iron catalyst. 5. Nitriding iron catalysts promoted the production of oxygenates with the following trends: (a) Maximum production of oxygenates (31% of the hydrocarbon fraction) occurred at 220 "C, 400 psig, and a fresh feed space velocity of 200 h-' and (b) the yield of oxygenates decreased a t temperatures above 240 "C. B. Slurry Process At about the same i;ime that the oil circulation process was being developed considerable work was being con-
8
10
Includes C;.
ducted on oil slurry operations. The process was not new. In 1932, researchers a t the Kaiser Wilhelm Institute were investigating the hydrogenation of CO in a liquid medium (30). Also, studies on a liquid-phase slurry process were carried out by Kolbel and Ackerman (52, 5 3 ) ; Kolbel et al. (54); Hall et al. (39), and Farley and Ray (25). In the slurry process, a finely divided iron catalyst suspended in oil is circulated through a reactor in the presence of carbon monoxide and hydrogen. Like the oil circulation process, the exothermic heat of reaction is absorbed by the oil medium. Several advantages are gained from this process, such as: (a) local overheating of the catalyst and subsequent formation of methane and carbon deposition are avoided; (b) a lower H 2 / C 0 ratio (to 0.6) can be used without carbon deposition on the catalyst; (c) catalyst handling is easy; and (d) reactor design is simple. Operation. A flow diagram of the slurry process is shown in Figure 2. Synthesis gas a t 350 psig was passed through a charcoal scrubber for removal of traces of sulfur, then through a flow controller, a high-pressure gas meter, a surge tank, and a preheater prior to entering the bottom of the reactor. The reactor was a 3 in. 0.d. pipe 10 ft in length. A 1.5-in. recycle line transferred oil and catalyst to the bottom of the reactor where the slurry came into contact with preheated synthesis gas. During reaction a major portion of the exothermic heat was absorbed by the oil medium, while a small amount was removed by evaporative cooling induced by reflux from the cooling coil in
180
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Ind. Eng. Chern. Prod. Res. Dev., Vol. 19,
Llquid level sight gloss
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600
800
1000
1200
1400
CATALYST AGE, HOURS
Figure 3. Temperature and conversion measurements as a function of catalyst age for promoted iron in the slurry process.
the expanded section in the top of the reactor. Product gases passed through a series of traps to remove entrained slurry and to recover condensable oil and water. The gaseous products passed through a back-pressure regulator, a low-pressure wet test meter, a continuous carbon dioxide analyzer, and then was vented from the system. Gaseous samples were analyzed daily while liquid samples were analyzed following their collection from the traps a t the end of a period. Promoted Iron Catalyst (69). The catalyst was prepared by precipitating the metal nitrate with sodium carbonate. After washing and calcination, the oxides were impregnated with copper and potassium, dried, and ground to less than 61 pm (250 mesh). The catalyst, 100 Fe:lO Cu:l KzO was suspended in a synthetic diesel oil, added to the reactor, and then precarbided by passing 1:l Hz/CO through the slurry a t 130 h-l and 100 psig. Initially, the temperature was 200 "C, but it was gradually increased to 243 "C during the 72 h induction period. During 7.5 weeks of operation, temperature was varied from 242 to
276 "C, pressure from 100 to 250 psig, and space velocity from 130 to 300 h-l. Temperature and conversion measurements as a function of catalyst age are illustrated in Figure 3. A t most conditions there was a steady decline in activity, thus requiring a continuous increase in temperature or pressure to compensate for the deactivation. The decrease in conversion after 280 h resulted from circulation difficulties, and the increase in conversion at 425 and 1000 h occurred after pressure and temperature changes. However, the continual loss in conversion following temperature and pressure increases suggested catalyst settling as a cause of low activity. Analysis of catalyst samples taken during 837 h of synthesis revealed a continual buildup of carbon. Increasing the pressure from 100 to 200 psig, while maintaining the space velocity constant, resulted in raising CO converstion from 56 to 78%. Changes in pressure had little effect on the production of light hydrocarbons. The hydrocarbon product distributions for several periods of operation are summarized in Table V. For initial operation (C740 h), a relatively constant product distribution was obtained. But when temperature was increased in order to maintain adequate CO conversion, a significant shift to lighter products occurred. Product water was found to contain almost 25% by weight of chemicals boiling below 100 "C, and the light oil product contained 44 wt '30of oxygenated materials-the predominate oxygenated compounds being ethyl, propyl, and butyl alcohols. Nitrided Iron Catalyst. Since oxygen-containing compounds can be used as solvents and as raw material for plasticizers and detergents, newly developed nitrided iron catalysts were of commercial interest. To investigate the possibility of using nitrided iron in commercial operation, the Pittsburgh Coke and Chemical Co. and PETC cooperated in a study to determine the operability of the catalyst in the slurry system (68). Fused and precipitated iron catalysts were reduced and nitrided as previously described for oil circulation experiments. The catalysts were induced using almost similar operating parameters as were used for the Cu-K promoted iron catalyst. The only exception to this was that over the 72 h precarbiding period, pressure was increased from 100 to 200 psig (i.e., stepwise activation). During synthesis, instead of recycling unreacted gas and products from the
Ind. Eng. Chem. Prod. Res. Dev., Vol. 19, No. 2, 1980
Table V. Hydrocarbon Distribution for Slurry Process Using a Copper-Potassium Promoted Iron Catalyst (69 )
__
-
period catalyst age, h operating conditions temperature, "C pressure, psig space velocitya product distribution, ,wt % C,-C, in gas gasoline ( < 2 0 4 "C) diesel oil (204-343 "C) heavy oil (> 343 "C)
" SCFH of fresh gaslft'
7
8
569 737
10
25 31 10 34
32 30 18 20
Table VII. Effect of Throughput o n Product Distribution from a Nitrided, Fused Iron Catalyst in Slurry Operation (68) space velocity, h-'
12
929 1001 1 2 4 1
255 256 263 200 250 250 240 300 300 26 31 11 32
11
215 250 300
276 250 300
34 30 14 22
33 35 24 8
of slurry in reactor.
Table VI. Effect of Temperature o n Product Distribution from Nitrided, Fused Iron Catalyst in Slurry Operation (68 )
181
H, + CO conversion: wt % product distribution, wt % CI + c, C, + hydrocarbons H,Ob oxygenates in H, 0 oxygenates in oil
200
300
400
500
62.0
46.6
37.9
33.8
26.4 39.1 11.0 14.7 8.8
19.1 39.5 14.2 13.8 13.4
20.6 36.3 16.6 13.2 13.2
19.6 34.9 20.0 15.2 10.3
a Pressure = 300 psig; temperature = 250 "C; feed gas = 1:l H,/CO; single pass operation. Exclusive of oxygenates.
Table VIII. Effect of Recycle Ratio o n Product Distribution from Nitrided, Fused Iron Catalyst in Slurry Operation (68)
_______
_ _ _ . . _ I . _ _
recycle ratio
temperature, "C ___.___
H, + CO conversion," % product distribution, wt % c, + c, C, + hydrocarbons H,Ob oxygenates in H,O oxygenates in oil
220
230
240
250
258
22.0
29.0
44.9 62.0 71.4
13.9 29.1 32.6 9.7 14.7
17.4 32.4 25.2 9.7 15.3
20.4 34.9 16.4 12.6 15.7
26.4 35.9 39.1 44.1 11.0 8.3 14.7 8.1 8.8 3.6
a Pressure = 300 psig; SV = 200 h..'; feed gas = 1 : l H,/ CO; single pass operation. Exclusive of oxygenates.
top of the reactor as was done when testing the Cu-K promoted iron, the product gas was passed through a series of cold traps to condense the middle oils, light oil, and aqueous product, then part of the dry exit gas was recycled back to the reactor (not shown in Figure 2). Experiments were continuously rim for 900 to 3200 h and were either terminated voluntarily or operated purposely under extreme conditions that caused a loss of nitrogen from the catalyst and reduced the yield of oxygenates. The influences of temperature, space velocity, gas recycling, and catalyst type (fused vs. precipitated) were studied a t 300 psig. The effects that temperature had on product distribution are shown in Table VI. The selectivity to CI + C2 and C,+ hydrocarbons increased with temperature. A maximum yield of oxygenates in water occurred at 250 "C, and the concentration of oxygenates in the oil phase reached a maximum a t 240 "C. The decrease in oxygenate yield a t the higher temperatures was caused by the chemical instability of the oxygen-containing compounds in the reaction environment. At synthesis conditions of 250 "C and 300 psig, as space velocity was increased from 200 to 500 h-' (Table VII), the concentration of oxygenates in the water remained essentially constant. In the oil phase, maximum yields of oxygenates were obtained at GHSV's of 300 and 400 h-l. Recycling moduct gas through the reactor decreased C, + C2 yieldand increased the iota1 oxygenate selectivit; (Table VIII). Results from experiments made with fused iron and precipitated iron nitrides are compared in Table IX. The percentage of total oxygenates was greater in the product obtained from the precipitated catalyst. Also, yields of oil-soluble oxygenates were higher with the precipitated catalyst, indicating a higher average molecular weight for the total oxygenates, but the precipit,ated catalyst appeared to be less active than the fused iron.
H, + CO conversion," % product distribution, wt %
c, + c, C, + hydrocarbons
H,Ob oxygenates in H, 0 oxygenates in oil
0
0.5
47.3
56.3
26.7 39.1 14.0 16.2 4.0
22.2 31.8 17.4 15.4 13.2
Pressure = 300 psig; SV = 300 h - ' ; feed gas = 1:1 H,/ CO; temperature = 250 "C. Exclusive of oxygenates.
SYN
sas
'I ,
F E E D LINES
(
M U L T I P L E THERMOCOUPLE WELL 8 BAFFLE TUBE
i
4
i --BOTTOM 5"NTHESlS GASFEFDLNE
Figure 4. Fluidized-hed reactor.
Conclusions 1. Two catalysts were evaluated in the slurry reactor: copper -potassium promoted iron and nitrided iron. 2. The following characteristics were observed for the Cu-K-Fe catalyst: (a) There was a steady decline in activity with time, thus requiring increases in temperature (242 to 276 "C) to compensate for deactivation. (b) Loss in conversion was related to catalyst settling and carbon buildup problems. (c) Large amounts of oxygenated compounds were found in the light oil and water fractions. 3. Nitrided iron (precipitated and fused) was more stable than the Cu-K-Fe catalyst. The precipitated nitrided iron catalyst produced a greater percentage of oxygenates, especially oil-soluble oxygenated hydrocarbons, than did nitrided fused iron.
Ind. Eng. Chem. Prod. Res. Dev., Vol. 19, No. 2, 1980
182
Table IX. Operating Conditions and hroduct Distribution from Fused and Precipitated Nitrided Iron Catalysts in t h e Slurry Process (68) ~
fused catalyst age, h 7 61 -9 05
1457-15 5 3
248-368
392-824
848-920
2 50 200 62.0
250 300 46.6
250 300 38.4
250 20 0 43.7
250 20 0 41.0
26.4 27.2 11.9 11.0 14.6
2.3 35.1 12.7 17.1 16.7 16.1
17.3 10.0 23.6 18.7 5.3 25.1
temperature, "C space velocity, h-' H, + CO conversion," % product distribution, wt % CI + c, C,-C, hydrocarbons oil hydrocarbons H, 0 oxygenates in H,O oxygenates in oil a
Pressure = 300 psig.
Table X.
~~~~~~~~~
precipitated catalyst age, h
8.8
21.3 12.4 21.9 11.1 4.8 28.4
24.3 14.1 18.0 5.1 12.8 25.7
Exclusive of oxygenates.
Operating Conditions and Yields from Nitrided Fused Iron Catalyst in a Fluidized-Bed Reactor (17) period
catalyst age, h pressure, psig average temperature, "C fresh feed space velocity,a h-' linear velocity, ft/s recycle ratio CO conversion, % H, + CO conversion, % material balance, wt % hydrocarbon yield, wt % C.
c, C,'
c 3
C,'
c4 C,' c,
1
2
3
4
5
7
8
10
30 100 238 1000 0.9 3:l 21.4 24.4 103.4
54 150 2 39 1500 0.9 3:l 15.2 18.3 100.2
78 200 2 39 2000 0.9 3:l 15.1 17.7
101.0
168 300 238 3000 0.9 3:l 21.2 24.3 93.5
240 300 240 1500 0.8 6: 1 28.1 34.6 100.6
411 300 251 1000 0.8 9:l 62.7 68.0 103.2
576 300 251 750 0.7 12:l 74.7 78.3 100.6
21.3 9.0 4.3 3.2 16.0 9.0 8.5
21.8 9.5 2.9 4.7 13.4 6.2 16.5
21.9 10.6 3.2 5.2 14.9 6.9 6.4
23.6 6.1 2.7 8.7 16.8 11.6 5.3
24.0 10.2 1.4 4.3 14.4 5.8 8.0
21.5 9.5 0.4 6.4 6.1 4.3 4.8
3.4
22.5 10.1 0.3 6.8 5.1 4.4 4.7 1.6 1.6 8.5 24.0 0.9 9.6 34.9 1l.2 53.9
-_
C,' 5.3 oxygenates in oil oxygenates in water 23.4 unknown oxygenates hydrocarbons condensed total product yield, w t % hydrocarbons 30.3 10.3 H2 0 59.4 CO, a Feed gas ratio (H,:CO) was 1:l.
__ __
__ __7 . 3 22.5 __ --
36.0 18.0 46.0
--
__ __ __
24.7
22.8
-_
__
__ __
2.0 12.8 26.1 1.3 4.6
36.3 23.6 40.1
3 5.9 23.3 40.8
35.4 13.8 50.9
_-
__
-_
32.7 22.4 44.8
__
-_
28.5
__
11
12
742 300 252 150 0.5 8:l 76.2 77.7 98.1
814 300 252 500 0.5 12:l 83.5 84.9 98.6
837 300 253
24.7 10.4
28.8 12.0
23.9 9.0
7.7 3.1 3.6 4.1
9.2 4.3 5.3 4.3 0.9 1.8
5.1 12.7 3.8 4.1 1.5 3.0
33.4
35.4
_-
__
1.7 4.5 30.8
__
1000
_-
--
9.3
_--
33.0 9.3 57.7
28.7 10.6 60.7
period
a
In the oil phase, 8 5 t o 90% were alcohols.
7
8
9
10
11
1000 0.8 251 9:l 68
150 0.7 25 2 12:l 78
750 0.6 252
750 0.5 252 8:l 78
500 0.5 252 12:l 85
15.9 44.8 19.4 5.7 4.9 4.1 3.2 1.5 0.5
15.3 40.7 26.7 5.7 4.4 3.4 2.7 0.7 0.4
14.2 42.7 20.4 5.8 6.3 5.8 3.0 1.0 0.7
12.8 41.3 20.9 6.5 7.6 6.5 2.8 0.9 0.6
13.0 42.4 22.1 6.8 6.3 5.5 2.4 0.8 0.7
45.4 28.5 18.5 7.6
48.0 36.8 13.2 2.0
35.7 36.4 20.1 7.7
39.8 35.7 16.3 6.1
1O:l 80
--
__ __
-_
Table XI. Selectivity of Oxygenates from Fluidized-Bed Operation ( 17)
fresh feed space velocity, h-' linear velocity, f t / s av temperature, "C recycle ratio H, t CO conversion, % oxygenated products in aqueous phase, wt % CH,OH C,H,OH C,H,OH C,H,OH C5HIlOH C,HI,OH acetone methyl ethyl ketone methyl propyl ketone oxygenated productsa in oil phase, wt % carbon no. 1-4 carbon no. 5-7 carbon no. 8-9 carbon no. 10-13 carbon no. 1 3 +
0.8 9:l 54.0 60.8 101.6
32.4 17.8 50.0
Ind. Eng. Chem. Prod. Res. Dev., Vol. 19, No. 2, 1980
183
Q EXCHANGER
LIGHTPRODUCT
PRODUCT GAS-
VENT 1
L
ANALYZER
MANOMETER
ICALORIMETER
I SYNTHESIS GAS
Figure 5. Fluidized-bed unit.
C. Fluidized-Bed System Nitrided fused iron catalysts were also evaluated in a fluidized-bed reactor ( 27). In earlier fluid-bed experiments using reduced iron, where monoxygenated products were the goal, temperatures greater than 300 "C were required to prevent agglomeration of the product on the catalyst. However, such high operating temperatures caused carbon deposition that resulted in rapid catalyst deactivation. A nitrided fused iron catalyst could be used a t lower temperatures and would produce alcohols that could be advantageously blended with gasoline (65). The purpose of testing nitrided fused iron in a fluidized-bed reactor was to determine if the shorter gas-catalyst contact time would change the yield of oxygenated compounds from that already observed in the oil circulation and slurry processes. Operation. The fluidized-catalyst reactor (Figure 4) consisted of a 1-in. schedule 30 stainless steel pipe 6 f t in length and was enclosed in a 3-in. schedule 40 heat-exchanger jacket. A 3/s-in. o.d. baffle tube with external fins was inserted in the center of the reactor to help disperse the catalyst and to aid in fluidization. It also served as a thermowell for a five-junction thermocouple. Two porous stainless steel filters a t the top of the reactor were used to prevent carryover of small catalyst particles with the product gas. A ball valve a t the bottom of the reactor prevented the catalyst from surging back into the inlet section of the unit. The flow diagram Sor the fluid-bed system is shown in Figure 5 . Synthesis gas a t 500 psig was passed through an activated carbon unit. The feed stream could enter the reactor a t three inlet ]points-located a t the bottom of the reactor and at points 18 and 36 in. from the bottom. Steam could be added to the synthesis gas by passing the bottom feed stream through an electrically heated saturator placed upstream from the preheater. Hydrogen and ammonia could also be added to the bottom of the reactor. Exit gases leaving the top of the reactor passed through a series of traps used to collect waxes, liquid hydrocarbons, oxygenates, and water. The gas stream from the liquidproduct receiver was split; one part went to a recycle compressor and the other passed through a pressure-reducer, a gas meter, and a COSanalyzer before being vented. Nitrided Fused Iron Catalyst. A fused iron catalyst, 120 to 230 mesh, wa:i reduced with H2 a t 400 "C, 1 atm,
Table XII. Product Yields Obtained with Nitrided Fused Iron Catalysts in Various Types of Reactors ( 17 ) fluidizedbed slurry
oil circulation
pressure, psig 30 0 300 300 temperature, 'C 251 247 230 space velocity, h - ' 1000= l O O b 200c H, + CO conversion, % 69.2 68.7 68.0 recycle ratio 2:l 1:l 9:l 1.2 usage ratio, H,:CO 0.95 0.75 hydrocarbon yield, wt 7'6 31.5 28.9 29.3 c, + c2 C, + gases 28.4 29.3 23.6 oxygenates in oil phase 2.9 12.9 14.7 oxygenates in water phase 18.1 20.6 26.1 condensed hydrocarbons 9.9 17.9 5.9 total yield, wt % hydrocarbons 35.4 29.9 26.4 13.7 9.3 2.2 H2 0 71.4 50.9 60.8 CO, a Based on settled-bed volume. Based o n slurry volume. Based o n fixed-bed volume.
and 3000 h-l GHSV based on the volume of the settled bed. The catalyst was nitrided with ammonia a t 325 to 375 "C, 1 atm, and a space velocity of 310 h-' to produce t-Fe2N. Because the stepwise activation of the nitrided catalyst used in the slurry operation had resulted in higher yields of oxygenates and less C1 + C2, a similar activation procedure was used. Synthesis was conducted at the following conditions: 238 to 253 "C, 100 to 300 psig, and fresh feed space velocities ranging from 500 to 3000 h-l. Recycle ratios were adjusted with each change in synthesis gas rate in order to maintain a superficial linear velocity of 0.5 to 0.9 ft/s through the catalyst bed. In the experiments using a nitrided fused iron catalyst it was not necessary to distribute the fresh feed for temperature control; therefore the fresh feed entered the bottom inlet and the upper inlets were used for pressure-drop measurements. Data pertaining to gaseous products were calculated from mass-spectrometer analyses of 3 to 4 spot samples taken during each steady-state period, and liquid products were accumulated during steady-state operation. Operating conditions and experimental results for run F-54 are summarized in Table X. The data indicates that a fluidized bed is well suited for synthesis of highly oxy-
184
Ind. Eng. Chem. Prod. Res. Dev., Vol. 19, No. 2, 1980 Pressure
Steam 1
Dowtherm
1
Reactor
3 I
I
'
400 p s i g ~
2200c~l&I
;T! F
Synthesis gas
compressor
Hot- gas
compressor Water
?
Water
$
I
C O z scrubber
Refrigerated water
@LB-*
Tail g a s
Figure 6. Hot-gas-recycle system.
genated products. The maximum operating temperature was limited to 255 "C. At higher temperatures nitrided iron catalyst decomposes and produces products similar in composition to those obtained from a reduced iron catalyst. The high linear velocity of gas through the reactor and the consequently short residence time favored the production of high yields of oxygenates. The selectivity of oxygenates in the aqueous and oil phases for fluidized-bed operation is given in Table XI. In Table XII, the product yields from a fluidized-bed, a slurry, and an oil circulation reactor are compared.
Conclusions 1. Nitrided iron performed well in a 837-h run a t low operating temperatures (238-253 "C). Use of non-nitrided iron required higher temperatures in order to prevent "waxing up" of the catalyst by the product. 2. Use of a fluidized bed is well suited for the synthesis of highly oxygenated products using nitrided iron catalysts. Of the total hydrocarbons produced, 25 to 36% by weight were water-soluble alcohols. 3. In comparison to the oil circulation and slurry processes using Fe2N, a larger weight percent of oxygenates was produced in the fluidized bed reactor. The high linear velocity of gas through the reactor (shorter residence time) favored production of oxygenated compounds.
D. Hot-Gas Recycle System In a continuing effort to evaluate various reactors for F-T synthesis, PETC directed its program next to the hot-gas recycle process. The process uses a fixed-bed catalyst through which large volumes of recycle gas are circulated to remove the heat of reaction as sensible heat. Investigations of the HGR process were first carried out in Germany by W. Michael and reported by Faragher and Foucher (24) in 1939. In this work an ammonia-type catalyst was compressed into 1-cm cubes and used in a fixed-bed reactor. Catalyst attrition, carbon deposition, and a high resistance to gas flow were major operating problems in the German process. The development by PETC of an active lathe turning catalyst having low resistance to gas flow, good heat transfer characteristics, and superior physical and chemical stability had the potential of overcoming many of the difficulties experienced by the
Germans (74). Experiments utilizing lathe turnings (12, 13, 27, 33) as well as fixed beds of steel wool (27) and parallel-plate assemblies containing active catalyst coatings (34, 36) will be discussed. Operation. A flowsheet of the hot-gas-recycle system is shown in Figure 6. The total feed gas, consisting of fresh and recycle gases, entered the top of the reactor and passed downward through the catalyst bed a t a superficial linear velocity of about 4 to 5 ft/s. The gas leaving the reactor entered a gas-to-gas heat exchanger and was cooled to 200 "C. Part of the product stream was cooled, metered, and sampled; the remainder was returned to the system via the hot-gas-recycle compressor. Five to 15% of the hot recycled gas was cooled to condense water vapor and oil and was then treated with a 20% monoethanolamine solution to absorb COz. The COz and water vapor content of the recycle stream were maintained below 10%. Before entering the reactor the mixed feed gas was heated by a gas-to-gas heat exchanger and a Dowtherm jacketed gas heater. The reactor was constructed from a 3-in. schedule 80 carbon steel pipe. 12 ft in length, and was electrically heated in order to maintain adiabatic conditions during operation. Thermocouples were placed in wells 1 ft apart along the length of the reactor. The catalyst was supported on a stainless steel screen located a few inches above the gas outlet. Bed heights up to 10 ft could be obtained. The top of the reactor was equipped with a removable head through which a magnetic sampler could be lowered to withdraw samples of catalyst from the top of the bed during operation. High Voidage Catalysts. The first catalysts tested were materials with high void space: steel lathe turnings and steel wool. These catalysts were activated by the following procedure: (a) oxidation with steam at 600 "C, 1 atm, until 20% of the iron was converted to Fe301; (b) impregnation with an aqueous solution of K2C03followed by; (c) H2 reduction a t 400 "C and 100 psig. In most experiments an induction period was used to bring the system on stream. Induction was accomplished by passing 1:l H 2 / C 0 through the reactor a t a GHSV of 200 h-' a t 260 "C, 400 psig, and a recycle rate sufficient to maintain the temperature differential over the catalyst bed at 10 "C. Synthesis was initiated by changing the fresh feed Hz/CO ratio to 1.3-1.5 and increasing the reaction temperature to 300 "C. The operating variables investigated were those
Ind. Eng. Chem. Prod. Res. Dev., Vol. 19, No. 2, 1980
185
Table XIII. Cold-Gas Recycle Operation Using Steel Lathe Turnings, Experiment 1 ( 1 2 ) period cat a1y st age , h fresh feed H, :CO ratio space velocity, h - ' recycle gas total recycle/fresh feed C0,-scrubbed recycle/fresh feed H, :CO ratio reactor conditions av temp, " C av pressure, psig gas velocity, Pt/s conversions CO conversion, % H, + CO conversion, % hydrocarbon dickribution, wt %
c, + c, cs
gasoline ( C,=-204 "C) diesel oil (204-31 6 "C) fuel oil (316-450 'C) wax (> 450 " C ) alcohols (C,, C , + C,) total product distribution, wt % hydrocarbons H, 0
co,
1
2
3
4
5
418-442
490-514
617 -641
6 6 5-6 89
761-785
1.28 1002
1.48 107 1
1.46 878
1.48 867
1.51 1099
60 1.9 1.01
60 1.8 1.56
50 2.3 1.82
40 2.2 1.91
40 1.7 2.25
302 406 7.5
303 404 8.1
29 8 400 5.5
300 398 4.4
310 40 2 5.7
87.6 89.0
90.1 89.7
89.2 87.5
88.8 86.8
92.7 90.5
27.6 10.7 44.0 4.2 2.3 0.7 10.4
33.3 13.0 37.8 3.2 1.7 0.3 10.6
31.3 12.0 38.9 4.7 2.6 0.5 9.9
31.6 12.0 38.3 5.1 2.8 1.2 9.2
32.4 11.7 40.0 5.5 1.7 0.5 8.3
30.3 24.5 45.2
32.6 26.6 40.8
33.0 27.0 40.0
33.8 26.0 40.2
32.7 27.2 40.1
Table XIV. HGR Experiments Using Steel Lathe Turnings and Steel Wool as Catalysts ( 2 7 ) catalyst expt no. 7
wt, lb
K,O,
%
90
14.2
0.05
voids, purpose OS expt first experiment with hot-gas com pressor
t Y Pe steellathe turning
%
induction none; temperature raised gradually t o 306 "C in 76h
t o obtain better activity by induction techniques to decrease methane yield by increasing K, 0 content of catalyst test activity oSstee1 wool
steellathe turnings
89
15.4
0.04
148 h at 270 " C with 1:l H,/CO a t 200 h - '
steellathe turnings
89
15.4
0.16
280 h at 270 "C with 1:1 H, /CO at 200 h - '
steel wool
97
4.3
0.12
none
16
induction studies & lower recycle ratios
steel lathe turnings
88
15.9
0.17
240 h at 270 "C with 1.4:1 H,/CO at 200-500 h - ' GHSV
17
test multiple feed to reactor
steel lathe turnings
88
15.9
0.18
none; temp raised gradually t o 320 "C in 43 h
8
9
11
comments and length of r u n in h activitv-lower than experiments with cold recycle compressor ; 682 h good activity; methane yield t o o high; 979 h
GHSV lowered methane yield; 2 265 h
GHSV
affecting product distribution, pressure drop, and catalyst activity. During the period in which a hot-gas-recycle compressor was being designed and constructed, the process was simulated using a cold-gas compressor. In the simulated, or cold-gas-recycle operation, the recycle gas was cooled, thus condensing out C5+ hydrocarbons and water, then compressed and reheated to reaction temperature. Operation using steel lathe turnings in the cold-gas-recycle mode was successful. Results are shown Table XIII. Good temperature control of the catalyst was obtained using gas recycle ratios of 40:l to 60:l. The observed pressure drop of 1.5 psi per foot of bed height was low enough to make the recycle method of heat removal very attractive. The purpose and performance of selected experiments conducted in the hot-gas-recycle mode are summarized in Table XIV. The pressure drop through the catalyst bed was affected by the dhape of the catalyst and the linear
low activity; shutdown due to carbon buildup; 366 h satisfactorv - oDeration: 920 h satisfactory operation; 711 h
velocity of the gas. A comparison of steel turnings with 88% percent voids to steel wool with 97% voids showed that the pressure drop was 50% less for steel wool. However, the steel turnings were more active. A pressure drop of less than 0.5 psi per foot of catalyst bed of steel lathe turnings was obtained by operating with a recycle to fresh feed ratio of 20 and by splitting the total gas flow and injecting portions into the reactor a t different entry ports. The effect which operating variables had on product distribution is given in Table XV. The product distribution was shifted toward heavier hydrocarbons by decreasing either the reaction temperature, H 2 / C 0 ratio, space velocity, HzO and COz content in recycle stream, or by increasing the potassium loading in the catalyst. The yield of gaseous olefins could be increased by operating a t a higher space velocity. Water and COPin the recycle stream inhibited the formation of oxygenates. The only operating difficulty encountered was the tendency of the
186
Ind. Eng. Chem. Prod. Res. Dev., Vol. 19, No. 2, 1980
Table XV. Effect of Variablesa o n Product Distribution for Hot-Gas-Recycle Experiments (27) reactor temp change experiment & period: recycle t o fresh feed ratio total cold average temp, "C CO, in recycle gas, % H,0 in recycle gas, % K, 0 content on catalyst, % space velocity, h-' H, :CO ratio fresh gas recycle CO conversion % hydrocarbon distribution, wt % c, + C,b c 3
fresh gas Ha :CO ratio
11-A
11-B
16-D
16-F
162 2.6
81 2.8
17 2.6 305 5.0 4.3
1000
353
390
4.8
5.0
265
244
hourly space velocity 8-H
8-K
16 4.5 313 2.4 4.9
60 2.8 300 10.1 4.1
60 2.9 329 10.1 4.0
1000
1003
1400
% H,O in recycle gas
% CO, in recycle gas
7-D
7-5
7-C
67 67 305 6.4 0.2
65 2.5 310 4.9 2.5
60 2.5 319 5.3 5.0
400
400
600
7-K 60 0 319 29.8
7.3
-
600
potassium impregnation 8-C 105 3.0 305 4.6 3.4 0.04 400
9-C 108 2.4 315 5.3 4.8 0.16
399
1.41 1.23 1.36 2.91 4.32 3.58 10.1 22.4 87.3 92.0 98.5 98.4
1.29 1.30 1.28 1.31 1.29 1.29 1.28 1.33 6.51 4.03 1.50 3.21 8.04 7.68 4.33 3.87 97.5 95.6 77.9 85.1 97.1 91.8 96.4 94.1
30.1 4.2 46.4 7,l 12.2
gasoline (C,'-204 o C ) c diesel oil (204-316 "C) fuel oil + wax (316 " C + ) total product distribution, wt % hydrocarbons 30.5 14.6 H, 0 54.9 co, oxygenates and olefins,e wt % C,-C, alcohols 3.0 other oxygenates 1.2 C,-C, olefins 19.8
48.3 6.8 42.1 2.5 0.3
47.3 16.5 35.5 0.7 0
59.5 17.5 23.0 0 0
33.9 13.0 45.8 6.8 0.4
32.1 11.2 50.6 5.8 0.3
17.3 0.5 48.9 13.0 20.3
25.7 1.3 49.5 10.2 13.3
19.5 3.1 58.1 11.3 7.4
24.9 3.7 59.8 6.7 4.9
38.0 15.2 39.4 3.6 3.8
21.4 3.1 60.3 10.4 4.8
26.5 10.9 62.6
30.1 17.6 52.3
44.9 29.0 52.5 17.0 2.f?id 54.0
29.7 17.4 52.9
31.3 28.7 40.0
27.2 11.6 61.2
30.4 19.4 50.2
30.5 5.5 64.0
29.2 15.5 55.3
28.1 17.1 54.8
3.3 1.4 23.4
1.2 0.2 6.0
2.4 0.5 16.8
4.5 3.8 21.9
3.5 1.3 20.3
4.5 0.9 22.2
2.2 0.6 34.0
1.2 0.6 9.4
4.1 1.4 27.5
1.9 0.3 4.3
3.0 0.3 10.1
' Italicized figures indicate change in variable. Includes paraffins and olefins. data (see ref 27, p 30). e Weight % of hydrocarbon fraction.
Includes oxygenates.
Questionable
Table XVI. Operating Conditions and Results from Testing Alkalized Lathe Turnings in a 12-in. Diameter Using t h e Hot-Gas-Recycle Process, Experiment 26 ( 3 3 ) period catalyst age, h bed temperature, "C average A temperature pressure, psig H,/CO in fresh feed fresh feed space velocity, h - ' total recycle in fresh feed, vol/vol CO conversion, % H, + CO conversion, % material balance, wt % hydrocarbon product distribution, wt %
c,
+
c,
c 3
gasoline (C,'-204 "C) diesel oil (204-316 "C) fuel oil (316-450 "C) wax (> 450 C) oxygenates in hydrocarbons, wt % C,-C, alcohols other oxygenates total product distribution, wt % hydrocarbons H, 0
co,
A
C
E
F
G
H
176
488
912
97 2
1032
1096
309 36 403 1.43 7 05 22 83.4 75.5 94.2
308 32 40 5 1.45 607 27 88.7 79.8 94.2
27 5 23 402 1.40 504 31 67.5 57.4 104.0
29 1 25 40 2 1.38 50 5 31 77.9 67.7 105.0
325 34 404 1.40 802 23 87.9 79.4 95.0
338 26 406 1.39 799 24 94.0 86.6 97.5
16.8 1.0 60.9 10.7 7.4 3.2
19.4 3.5 59.0 9.1 6.2 2.8
18.4 0 49.0 10.2 9.8 12.6
20.2 3.7 46.7 5.6 8.3 15.5
18.7 2.0 62.4 9.3 5.0 2.6
23.9 3.0 59.7 8.8 3.3 1.3
3.7 1.6
4.2 1.6
2.4 0.5
3.2 1.0
3.6 1.6
2.6 1.0
30.5 18.0 51.5
31.0 19.3 49.7
31.5 15.9 52.6
33.3 16.9 49.8
31.6 20.2 48.2
32.3 20.9 46.8
catalyst to spa11 after prolonged synthesis. The fines produced by spalling tended to collect in the lower section of the bed together with carbon formed during the reaction. These deposits decreased the void volume and caused an increase in pressure drop. After 2 to 3 months of operation, synthesis was usually terminated due to a decline in catalyst activity. The highest gasoline fraction, 62% of total hydrocarbons, was obtained using alkalized lathe turnings in a 12-in. diameter reactor (33). Product distribution data (Table XVI) showed maximum gasoline selectivity oc-
curring a t 325 "C. Below 325 "C, as temperature was decreased the percentage of gasoline decreased, while the heavier hydrocarbons increased. Above 325 "C, gasoline yields decreased as temperature was increased possibly as a result of thermal cracking. Parallel-Plate Assemblies. With the objective of minimizing pressure drop through the catalyst bed the performance of parallel-plate assemblies coated with catalyst was evaluated. Carbon steel plates, 6 in. in length but of varying widths, were flame-sprayed with either fused iron (34)or magnetite ore (21). Fifteen sprayed plates were
Ind. Eng. Chem. Prod. Res. Dev., Vol. 19. No. 2, 1980
187
Table XVII. Testing o f Flame-Sprayed Fused Iron o n Parallel Plate Assemblies in HGR Process, HGR-33 ( 3 4 ) period bed temperature, "C average A temperature pressure, psig Ha/CO in fresh feed fresh feed space velocity, h-' total recycle in fresh feed, vol/vol CO conversion, % H, + CO conversion, % material balance, w t % hydrocarbon product distribution, wt % Ci + Ca c 3
A
Bb
C
D
E
FC
G
H
269 20 400 1.4:l 600 52 80.6 76.4 93.6
265 19 4 00 1.4:1 600 51 82.0 74.7 96.3
27 1 20 400 1.4:l 600 50 82.1 75.0 98.0
27 8 20 40 0 1.4:l 600 51 81.8 74.6 95.7
2 80 20 400 600 52 76.7 73.1 95.6
278 20 400 1:l 600 51 75.6 72.9 95.8
279 41 400 1:l 600 24 76.8 72.6 97.6
288 20 400 1:l 600 54 89.8 85.2 93.6
59.7 6.6 31.8 1.9 0 0
49.1 10.6 33.6 2.5 3.0 1.2
46.1 10.4 41.5 2.0 0 0
47.5 10.3 35.9 2.5 2.1 1.7
36.1 6.1 47.7 2.1 5.3 2.7
31.0 6.7 50.4 3.5 3.4 5.0
30.7 6.5 51.8 2.1 5.0 3.9
37.7 7.8 45.5 3.8 4.1 1.1
1:l
gasoline (C,'-204 "C) diesel (204-316 "C) fuel oil (316-450 "C) wax(>45O0C) oxygenates in hydrocarbons, wt % C,-C, alcohols 7.6 6.1 5.9 5.2 4.0 3.4 3.4 4.3 other oxygenates 0.4 0.8 0.7 0.9 1.1 1.1 1.4 0.8 total product distribution, wt % hydrocarbons 28.0 28.4 29.4 27.3 26.4 26.9 27.7 25.8 16.0 16.1 16.3 16.8 10.6 10.4 9.8 9.4 HlO 55.6 54.5 55.9 55.7 63.0 62.7 62.5 64.8 COl a The catalyst age was not given in ref 3 4 . The catalyst was immersed in an alcoholic solution of 3% KOH prior t o these periods. The Bauxite reactor was in service during these periods.
assembled into a bundle with approximately 0.1 in. between each plate. For testing, 12 bundles were transferred to the 3-in. diameter reactor. The recycle system and operation were the same as those used for testing the fixed bed of lathe turnings (27). T h e flame-sprayed assemblies of fused iron (D-3001) were not treated with alkali before synthesis. Operating conditions and results for run HGR-33 are presented Table XVII. Attempts to iincrease gasoline yield by addition of alkali to the catalyst after periods A and B and by using A1,0, to upgrade the product steam during periods D and F were unsuccessful. A significant increase in the selectivity to gasoline occurred when the H2/C0 ratio in the feed was decreased from 1.4:l to 1.1. Decreasing the recycle ratio (period Cr) had a slight beneficial effect on gasoline production. The performance of magnetite ore was tested in a similar manner as was used for HGR-33. The assembled catalyst was impregnated with 0.13 wt % K 2 0 by soaking the bundles in a 6% solution of K2C03before being placed in the reactor. Product yields for experiment HGR-34 at space velocity of 1000 and 2000 are presented in Table XVIII. The make of gasoline was about 48% of the total hydrocarbon production. The high CO conversion a t 2000 space velocity operation indicated that magnetite was a very active catalyst. The experiment was terminated voluntarily after 540 h of synthesis.
Conclusions 1. Use of an inexpensive lathe-turning catalyst with a high void content pe:rmitted successful operation of the simulated hot-gas-recycle process. Excellent temperature control of the catalyst was maintained using recycle-gas ratios of 40:l to 60:l. 2. In the hot-gas-recycle operation carbon steel lathe turnings were satisfactory catalysts, whereas steel wool catalysts with greater void volume were not as active. Good temperature control was achieved by operating with a recycle-to-fresh feeld ratio as low as 20. An alkali promoter, K 2 0 , on the catalyst caused a shift in product
Table XVIII. Testing of Flamesprayed Magnetite on Parallel Plate Assemblies in the HGR Process, HGR-34 (21) period catalyst age, h bed temperature, "C average A temperature pressure, p i g &/CO in fresh feed fresh feed space velocity, h-l total recycle in fresh feed, vol/vol CO conversion, % H, + CO conversion, 75 hydrocarbon product distribution,
C
D
H
260
308
476
325 50 400 1:4:1 1000 15.9
320 40 400 1:4:1 1000 20.4
325 50 400 1:4:1 2000 14.4
98.8 94.4
98.2 93.4
94.4 87.5
36.4 13.2 44.1 5.0 0.4 0.3
38.4 13.9 42.5 4.5 0.2 0.1
33.8 12.5 48.3 4.4 0.5 0.4
29.3 19.0 51.7
30.3 14.0 55.7
33.3 21.1 45.6
wt %
Cl + c, C, gasoline (C3'-204 C) diesel (204-316 "C) fuel oil (316-415 "C) wax ( > 4 5 0 "C) total product distribution, wt % hydrocarbons H2 0 CO,
distribution, making more gasoline and less gaseous products. The recommended quantity of K 2 0 to produce the maximum gasoline yield was from 0.13 to 0.18 wt 70. 3. A 12-in. diameter reactor using lathe turnings was operated successfully for over 1000 h. 4. A major operating difficulty encountered when using steel lathe turnings was spalling of the catalyst after prolonged synthesis. The collection of fines a t the bottom of the reactor caused an increase in pressure drop. 5. To overcome problems of catalyst spalling, the performance of thermally sprayed parallel plates was evaluated. In this mode of operation flame-sprayed fused iron and magnetite proved to be durable and very active catalysts.
188
Ind. Eng. Chem. Prod Res. Dev., Vol. 19, No. 2, 1980
E. Tube Wall Reactor The tube wall reactor was developed at PETC for the synthesis of methane (18,28,35, 41,42, 56, 60) and now utilization of TWR technology for the F-T process is under investigation. The catalyst is applied to the outer wall of a support pipe, and a coolant medium within the support pipe is used to effectively remove the heat of reaction. In addition to low pressure drop, the tube wall reactor has demonstrated the advantage of good temperature control. Operation. A flow diagram of the TWR bench-scale unit is shown in Figure 7 . Synthesis gas was compressed, dried, and desulfurized before entering the reactor. During F-T operation. reaction temperature was varied from 255 to 342 "C and pressure from 300 to 1000 psig. Heavy hydrocarbon products (waxes) were removed in a steam trap, and the remaining liquid products were passed through a water-cooled condenser and collected. The gaseous products, reduced to atmospheric pressure, were metered and the CO content was continuously monitored with an infrared analyzer. On a daily basis product gas samples were collected for gas chromatography analysis, and temperature profiles along the length of the catalyst bed were taken. At the termination of a period the condensed products were collected and separated into aqueous and oil fractions and each analyzed. The reactor, shown in Figure 8, consisted of two carbon steel pipes, the outer one 1.5 in. 0.d. X 35 in. in length and the inner one 0.75 in. 0.d. X 24 in. in length. The inner pipe, usually referred to as the catalyst support tube, was welded to a support flange that could be positioned in the center of the reactor shell. In the reactor, the top of the catalyst bed was 3 in. from the feed gas inlet, and the width of the annulus section where gas flows was 0.20 in. Thermocouple wells, parallel to and butting the support tube, permitted temperature profiles to be measured over the length of the reactor zone. During synthesis the exothermic heat of reaction was removed by boiling Dowtherm contained within the support tube. Dowtherm vapor passed to the top section of the reactor where it was condensed, and the hot liquid coolant was transferred back to the bottom of the support tube via a dip tube. The dip tube extended from a reservoir in the top section of the reactor condenser down through the center of the support tube. The temperature of the catalyst bed, or rate of heat removal, was maintained by controlling the Dowtherm vapor pressure. The catalysts employed during the initial series of tests (40)were magnetite (Fe304),a beneficiated taconite ore (Fe203+ Fe,04), and an iron-aluminum alloy (Raney iron). The materials were applied to the outside of the support tube by either a flame-spraying or plasma-spraying operation (48). The flame-spraying technique utilized an oxygen-hydrogen flame to melt the powdered iron oxide feed. The plasma process does not require a combustible gas mixture, but instead excites a plasma gas (a mixture of H2 and N2) to a flame temperature of approximately 30 000 O F . This temperature is about ten times that obtained by the O2-Hzflame. When the iron catalyst was applied via the flame-spraying technique, the support surface was first grit blasted and then coated with a nickel aluminide bond coat. The bond coat (95% Ni and 5% All, after leaching and reduction, has been shown to have negligible synthesis activity. The bond coat was not needed when the catalyst was applied by the plasma technique. In both spray procedures, the catalyst was applied to a thickness of approximately 0.023 in. over a 6-in. section of the support tube. Following catalyst application, the support tube was
--iiRecctor -r
z
a
C
O infrared monllor
U 8
Figure 7. T u b e wall reactor bench scale unit.
n,
A
F f 0 I
0
support flange
&-Gas
inlet
01
c
N 0
Ri I / / outlet-~
I
I
3 Figure 8. Tube wall reactor.
assembled into the reactor, and the catalyst was activated. For the iron oxide materials, an active surface was prepared by hydrogen reduction at 400 "C and atmospheric pressure. The degree of reduction was determined by monitoring the water content in the exit stream. The reduction period for these experiments was continued for 24 h after the formation of water (collected in the cold trap) ceased. When Raney iron was used, a 5% solution of NaOH was used to digest 70% of the aluminum from the alloy. The extent of leaching was determined by monitoring the volume of hydrogen evolved using a wet test meter. The remaining catalyst material, consisting of unreacted alloy and elemental iron, was washed with water and then was reduced in HP. Following reduction, and in some cases induction, synthesis gas was introduced
Ind. Eng. Chem. Prod. Res. Dev., Vol. 19, No. 2, 1980 Table XIX.
General Performance of Catalyst in TWR Operation ( 4 0 )
FT-STW no. 1 2 3
4 6
189
catalyst (I
synthesis time, co h conversionb
magnetite (PL) Raney iron ( F L ) (with induction) magnetite ( F L ) Raney iron (PL) taconite ( F L ) (with induction)
262 357 261 3 50 349
78.3 93.0 89.9' 72.0 77.2d
C, t C,
gasoline
diesel
total synthesis, h
71.2 71.1 70.8 68.3 69.7
9.3 7.9 16.9 19.5 14.2
1.9 0.8 2.0 1.4 3.4
1080 1529 1005 854 1833
wt % of HC product
At 300 psig, 3 H,/CO, 305, and 325-328 "C. 5 is the exposure velocit PL = plasma sprayed; F L = flame sprayed. Feed rate for FT-STW-3 was 20J. A 't for fresh synthesis gas feed expressed as SCFH/ftZ of geometric catalyst surface. FT-STW-6, period F, feed rate was 23.6Jwith H,CO = 2.1; temperature was 335 'C, pressure was 650 psi. Table XX. Conversionis and Product Yields from Flamesprayed Taconite in Tube-Wall Reactor ( 4 0 ) period length of period, h time o n stream, h temperature,(I "C pressure, psig ' CO conversion, % H2 + CO conversion, % hydrocarbon product distribution, wt % Ci + Ca c 3
gasoline (C,=-204 "C) diesel (204-316 "C) fuel oil (316-450 "C) wax (> 450 "C) oxygenates in hydrocarbons, wt % C,-C, alcohols other oxygenates total product distribution, wt % hydrocarbons H2 0
co2
a
D
E
F
G
I
M
0
19 181 27 5 650 30.7 28.4
96 277 310 650 54.8 51.1
72 349 335 650 77.2 66.1
168 517 340 650 93.0 75.5
144 829 340 650 94.6 76.6
168 1499 341 1000 95.9 79.1
168 1833 343 1000 88.7 65.2
60.0 9.9 24.1 4.3 1.7
67.0 10.3 20.6 1.1 1.0
69.7 10.5 14.2 3.4 2.2
69.1 11.5 17.5 1.5 0.4
70.7 11.0 16.7 1.2 0.4
72.1 9.6 17.1 0.9 0.3
73.4 11.0 14.1 1.2 0.3
15.2 0.8
11.8 0.3
4.0 0.3
6.0 0.3
2.4
_-
5.2 2.6
0.3
46.7 40.6 12.7
39.0 31.5 29.5
35.3 23.8 40.9
33.3 23.0 43.7
33.5 22.6 44.0
34.5 22.4 43.1
32.6 11.0 56.5
__
_-
-_
-
Hot spot temperatwe, usually temperature a t bed inlet.
Table XXI.
Summary (of t h e F-T Processes Investigated at PETC major process variables
process I. oil circulation
mode of operation
(a) CO. H, and recvcle oil are Dassed through a bed i f catalyst. ( b ) Catalyst is maintained as a fixed bed or a moving bed. ( c ) Exothermic heat is removed by recycle oil which is externally cooled and returned t o reactor. 11. slurry ( a ) Synthesis gas is bubbled through a slurry or suspension of catalyst in cooling oil. (b) Heat of reaction is adsorbed by the oil medium. 111. fluidized-bed ( a ) Synthesis gas is used t o fluidize a bed of 120-230 mesh iron. IV. hot-gas-recycle ( a ) Synthesis gas is passed through a fixed bed of iron. ( b ) Large volumes of recycle gas are used t o remove the heat of reaction as sensible heat V. tube-wall reactor (a) Catalyst is flame-sprayed t o outer wall of support pipe or heat exchanger tube (I)) Synthesis gas is passed over the catalyst coated pipe and a coolant medium within the pipe absorbs the heat of reaction.
into the reactor a t 300 psig and a t the lowest temperature to be evaluated. Results. The objective of the initial evaluation of the TWR for the synthesis, of gaseous and liquid fuels was to
catalysts tested ( a ) precipitated and fused iron ( b ) steel lathe turnings ( c ) steel shot ( d ) nitrided fused iron
temp, "C
pressure, psig
( a ) 236-251 ( a ) 300, 4 0 0 , 4 5 0 ( b ) 270-290 ( c ) 300 ( d ) 210-240
( b ) 400 ( c ) 400 (d) 300,400
( a ) copper, potassium (a) 242-276 promoted iron ( b ) nitrided iron ( b ) 220-258 (precipitated and fused)
( a ) 100, 200, 250
( a ) nitrided fused iron
( a ) 238-253
(a) 100-300
( a ) steel lathe turnings ( b ) Frallel plates coated with fused iron ( c ) parallel plates coated with magnetite ( a ) magnetite ( b ) Raney iron ( c ) taconite
( a ) 300-390 ( a ) 400 ( b ) 265-288 ( b ) 400
( b ) 250, 300
(c) 320-325 ( c ) 400 ( a ) 325-340 ( a ) 300, 650 ( b ) 325-340 ( b ) 300, 650 (c) 325-340 (c) 300, 650, 1000
determine the effects of temperature and pressure on catalyst activity and product distribution and also to compare the performance of plasma- vs. flame-sprayed catalysts. The general performance of the catalysts is
190
Ind. Eng. Chem. Prod. Res.
Dev., Vol. 19, No. 2,
1980
tabulated in Table XIX. These preliminary data indicate that the flame-spraying operation produced a more active catalyst than was produced by the plasma technique. However, much more work needs to be done to establish a data base before definite conclusions concerning the catalyst application technique can be made. Of the catalysts tested (40), taconite iron ore demonstrated the best activity and life. The material is easy to flame-spray and is very inexpensive. The catalyst was operated for 2.6 months before the run was voluntarily terminated. During the first two weeks of operation the beneficial effect of induction was demonstrated. During the initial activation procedure the catalyst was not precarbided, and the average CO conversion during period A (94 h) was 24% a t 275 "C and 300 psig. Following period A the catalyst was subjected to a 71-h induction treatment-2/1 synthesis gas over the catalyst a t 250 "C and 80 psig. During period B, CO conversion increased to 45% a t 255 "C and 300 psig. The effects of temperature and pressure on conversion and product distribution for several periods during run FT-STW-6 are given in Table XX. The majority of the hydrocarbon product was light gases (C, to C3). In general, increasing temperature a t constant pressure shifted the yield of hydrocarbon products to lighter molecular weight compounds. A t approximately 340 "C (periods G, I, M, and 0),changes in system pressure did not alter the hydrocarbon product distribution, and except for period I the overall product distribution was similar. In a recent test flame-sprayed taconite operated for 8 months in the bench-scale unit (43). After 6 months of synthesis the catalyst started to deactivate and was subjected to its first regeneration treatment. Catalyst activity was restored by reducing the deactivated iron in H2a t 400 "C followed by induction at 240 "C, but catalyst performance following a second regeneration after 8 months of operation was only satisfactory for one week.
Conclusions 1. Satisfactory CO conversions were obtained with flame-sprayed magnetite and Raney iron. 2. Flame-sprayed taconite demonstrated the longest catalyst life. The catalyst was successfully regenerated after 6 months of operation, and the total life was approximately 8.5 months. The majority of hydrocarbon products were light gases, and the maximum selectivity to raw gasoline was 24% of all hydrocarbons produced. 3. Current investigations are being directed at improving the liquid yield by addition of promoters to taconite. Current Fischer-Tropsch Processes Today the only commercial operating Fischer-Tropsch plant is SASOL in South Africa (47). Synthesis gas derived from coal is reacted to liquid products via two catalytic processes. Five fixed-bed reactors (ARGE process) containing precipitated iron are used to produce predominantly straight-chain, high-boiling hydrocarbons, and three fluid-bed units (Synthol system) employing an ammonia synthesis catalyst are used to produce mainly low-boiling products, that is C1-C4 and gasoline (37). The daily capacity of liquid products is about 2000 bbl for each fluidized-bed reactor and 550 bbl for each fixed-bed unit. Although the SASOL plant has a low thermal efficiency (38),South Africa's lack of oil resources, the country's desire for energy independence, and their possession of a sufficient supply of economically minable coal provides a suitable environment for the operation of the plant. In December 1974, SASOL announced it will build a new
oil-from-coal plant (SASOL-11),producing 50 000 bbl/day of products, approximately 7 times the output of SASOL I (46). Construction of SASOL-I1 is expected to be completed by 1981. Because of the decreasing U S . reserves of gaseous and liquid fuels, the 1973 oil embargo, and the increases in the price of gasoline, government interest has been renewed in F-T as an attractive process for making very clean co-product fuels from coal, namely substitute natural gas (SNG) and petroleum-like liquids. This renewed interest by the US.Department of Energy resulted in a contracted study to the Ralph M. Parsons Co. (58,59) for F-T economics and conceptual plant designs. Results of the study indicated that it would be possible to increase overall plant thermal efficiency and that the use of a tube wall reactor would be a major factor in achieving this goal. The TWR was developed at PETC for methanation of synthesis gas (18, 28, 35, 41, 42, 56, 60), and now utilization of TWR technology for the F-T process is under investigation.
F. Summary A summary of the Fischer-Tropsch processes investigated at the Pittsburgh Energy Technology Center is given in Table XXI.
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CATALYST SECTION
Strong Acid Chemistry. 8. Alkane Alkylation Reactions in HF-TaF, Michael Siskin* and Ivan Mayer Corporai'e Research Science Laboratories, Exxon Research and Engineering Company, Linden, New Jersey 07036
The direct alkylation of n-butane by larger alkanes in HF-TaF5 has been demonstrated. Evidence for the analogous direct alkylation of propane has been presented but the data are not conclusive. A general mechanism has been elucidated for this alkane alkylation reaction.
Higher normal alkanes (>C,) undergo rapid ionization via hydride abstraction in strong acids. The reactive secondary ions so formed rearrange to the more stable tertiary ions. These tertiary ions can then undergo cleavage by a @-scissionmechanism to form the most stable tertiary cation and the complementary lower alkene (Condon, 1958). If the same reaction is carried out in an excess of a solvent, such as isobutane or isopentane, the ionized higher normal alkane preferentially loses a proton and the corresponding branched alkene formed is immediately alkylated by the ionized solvent. The larger cation thus formed is still more reactive and cracks to a smaller carbocation and a lower alkene (Condon, 1958). In the case 0 196-432 1/80/ 1219-0191$01 .OO/O
of n-heptane in isobutane solvent the following sequence of reactions takes place (eq 1-4). hydride abstraction n-C;H,, + H' i-C;HlS+ + H, (1) deprotonation (alkene formation) T
h
-H+
i-CjHl>+ i-CiH,, alkylation of Ci alkene i-C7H14+ t-C,HS+ CllH2,3+ @-scission
-
C,lH,,+ C
H+
C,Hl,+ + CsH,2
1980 American Chemical Society
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