Fischer-Tropsch Oil Circulation Processes J. H. CROWELL, H. E. BENSON, J. H. FIELD,
AND H.
H. STORCH
Ofice of Synthetic Liquid Fuels, Bureau of Mines, Bruceton, Pa.
U
S D E R the provisions and that most of the heat was aurninary of Fischer-Tropsch process development is liberated at that region (34, of the Synthetic Liquid given including the essential features of the German fixed Sf?). Lurgi (37) and, later Fuels Act of April 5, 1944 bed system and the hot gas recycle and oil circulation procRuhrchemie (36)investigated (Public Law 290), the esaes; the oil slurry and fluidized bed systems are briefly the effect of recycling 1 to 5 volumes of tail gas per volume Bureau of Mines, United described. Most of the pilot plant experimentation at the of fresh gas. A more uniform States Department of the Bureau of Mines has been directed toward the development temperature gradient in the Interior, was directed to conof the oil circulation process because a simple converter catalyst bed and an increase duct laboratory research and design could be used and precise operating control atin the olefin content of the products resulted from recyto build and operate pilot tained in such a system. Several moderately successful cling of the tail gas. and demonstration plants for experiments, which employed a cooling oil circulating the production of synthetic 2. Processes using fixed through a fixed catalyst bed, were operated. However, liquid fuels from such mateshutdowns were necessitated after 3 or 4 months of synbeds of granular or pelleted rials as coal, oil shale, and thesis because the catalyst particles became cemented tocatalysts with internal coola g r i c u l t u r a l and forestry gether, and an excessive pressure drop developed across ing by direct heat exchange, products. As part of this the catalyst bed. This problem of catalyst cementing was such as the hot gas recycle program, the gas synthesis circumvented by employing a moving bed of catalyst in and oil recycle processes: section, research and dewhich lifting action of the circulating oil was used to exMichael of I. G. Farbeninvelopment branch, of the pand the catalyst bed and keep the particles in motion. dustrie developed the hot gss recycle process (6, 38, 33)usOffice of Synthetic Liquid Details of induction and synthesis operations, operating ing an iron catalyst and a conFuels, Bruceton, Pa., is indata, analyses of gas and liquid streams, calculated yields, verter which could be opervestigating the development and a typical product distribution are given for various ated without internal coolers of new F i s c h e r - T r o p s c h experiments. Calculated yields of gasoline and Diesel in the reaction space. About 100 volumes of tail gas were catalysts and is testing mateproducts resulting from thermal and catalytic cracking recycled per volume of fresh rials and procassea on both of heavy distillate and wax are given along with characterifeed gas, and the heat of reacbench scale and pilot plant zation of the gasoline and Diesel oil. tion was removed directly aa units. sensible heat in the recycled gas. A temperature rise of During the course of past 10' C. acrosa the catalyst bed wm maintained, and the heat of development of the Fischer-Tropsch synthesis in Germany and reaction was recovered in an external heat exchanger in the gas in the United States, several processes, differing chiefly in the rec clecircuit. In spite of the high gas throughput, it was not posmethod of removing the heat of reaction, have been studied. sib& to maintain uniform gas velocity throughout the bed. Hot spots formed in the bed resultingin rapid deterioration of catalyst. These processes have been classified by Storch (31) in the order The oil recycle process first investigated at I. G. Farbenindustrie of their development as follows: in Germany (5,6, 83,38) removed heat formed in the synthesis by 1. Processes using fixed beds of granular or pelleted catalysts, direct exchange with an oil circulating throu h the submer ed iron cooled by indirect heat exchange, with and without recycle of st bed, This process was investigatef in the Unite3 States e Standard Oil Development Company (W), The Texas the tail gas: (1, SO),and the Bureau of Mines. The fixed catalyst bed cooled by indirect heat exchange as develo ed by Ruhrchemie (3 7,13-15,21, .%) in Germany was used 3. Processes using suspensions of finely powdered catalysts in t i e first commercial kischer-Tropsch units. In this t e in the liquid phase; the oil slurry and fluidized bed systems: operation, synthesis gas was assed through the fixed bed% catalyst, and the heat evolve$ was transferred through metal German investigators a t Rheinpreussen (IO,l a ) , Ruhrchemie walls,to a cooling medium. Complex finned multitube reactors or ( 9 ) , and I. G. Farbenindustrie conducted laboratory and pilot relatively costly double-tube boiler-type converters were replant studies of the slurry process in which a finely divided quired to provide enough cooling surface to prevent overheating catalyst was suspended in a cooling oil, and the resultin slurry and deterioration of the catalyst. All the full scale German plants was circulated through the reactor concurrently with t i e synoperated during 1938 to 1944 employed this process of heat r e thesis gas. A bench scale study of this process is being made at moval and used a cobalkthorium dioxide-magnesium oxidethe Bureau of Mines. kieselguhr catalyst a t temperatures of 180" to 200' C.and presSeveral American companies. have investigated the Fischersures of either 1 or 10 atmospheres. Tropsch synthesis, using a flmdmed catalyst bed (18-,%'0, 86-89). Precise temperature control w y difficult to maintain in the Fresh Synthesis and recycle gas pass upward through a bed of 10Ruhrchemie fixed bed system, foy it had been found, in 1935, that to 1Wmicron catalyst particles suspended in the stream of gas, most of the reaction took place m the middle of the catalyst bed 2376
INDUSTRIAL AND ENGINEERING CHEMISTRY
November 1950
and the heat liberated during the synthesis is efficiently transferred by the fluidized catalyst to heat exchangers built inside the converter. High space-time yields have been attained in the fluidized process. This report describes the development of the oil circulation rocew, at the Bureau of Mines, usin a submerged catalyst bed. guch a system permits use of a sim$e reactor with no internal cooling surface and rovides direct control of the catalyst tempersture, eliminating t i e possibility of developing hot spots in the converter with resulting deposition of carbon and deterioration of catalyst. DEVELOPMENT OF OIL RECYCLE PROCESS
Development of the oil recycle process with a fixed, submerged oatalyst bed was started in 1934 at the I. G, Farbenindustrie at Oppau by Duftschmid, Linckh, and Winkler (6, 6, 88, 88). For several years, pilot plants with converter volumes of 7 to 50 cubic feet were successfully operated by I. G. Farbenindustrie. However, because of World War 11, the process never reached a commercial scale. The initial work at the Bureau of Mines, using oil as a coolant, was begun on a bench size unit in 1944, with a flow of oil trickling downward over the catalyst. In a ahort time, it was found that smoother operation and better temperature control could be obtained by submerging the catalyst bed completely in the oil. After the submerged bed process was developed sufficiently, a 3-gallon-per-day pilot plant was constructed, and operation of this unit was started in 1947. The construction of a larger pilot plant unit with a capacity of 1 barrel per day was recently completed. Fresh feed and recycle gas enter the bottom of the converter and flow upward concurrently with the circulating oil. As the gas passes over the catalyst, reaction occurs and heat is transferred to the cooling oil. An overflow pipe, connected with the converter at a point above the catalyst bed, conducts the exit gas and oil to a separating tank, from which the gases and vapors pass overhead to a condenser and recovery system, and the cooling oil is returned to the reactor by a circulating pump. There are several fundamental differences in the fixed bed operation at the Bureau of Mines as compared to that of Duftachmid: 1. Use of a higher boiling cooling oil (300" to 450' C. boiling range at atmospheric pressure) than that recommended and emloyed in Germany resulted in removal of the heat of reaction argely as sensible heat rather than by vaporization. This allowed more efficient recovery of the heat of reaction, but, of more im ortance, it made the system pressure an independefit varia b g , whereas in Duftschmid's experiments as vaporization of the oil was inherent in the process, the o erating pressure de ended on the operating temperature and boiing ran e of the coo&g oil. Early exporimontation using evaporative cooing ws3 terminated because of this limitation. 2. BecauRe most of the heRt was removed as sensible heat in the recycle oil, a variation of the oil rmycle rate made it possible to vary the temperature differential across the catal st bed. Tem erature differentials of about 15' to 20' C. werc useiaa comare! with those of about 50' C. in Duftschmid's experiments. !hecause the entire catalyst bed is a t the most effective operating temperature, greater catalyst activity and durability are attained. 3. Use of a recycled tail gas, from which the water formed in the reaction had first been largely condensed, reduced the partiul preeaure of the water and affected the shift reaction aa shown in the equation
P
Hs
+ COP-HoO
+ CO
2317
tion they crumbled and'matted together, and inoperable conditions resulted. Fused, synthetic ammonia-type catalysts were next used in the synthesis because of their hard, glassy structure and high resistance to crushing. Several experiments were made at 300 pounds per square inch gage and various space velocities with 1: 1 hydrogen to carbon monoxide feed gas for periods of 2 to 3 months. At the end of that time, however, even though the catalyst was still active, it became necessary to terminate these experiments because the pressure drop across the catalyst bed had steadily increased until it interfered with circulation of the cooling oil. The bulk of the catalyst particles had retained their original size but were tightly cemented together and could be removed from tho converter only by chiseling and drilling. Several changes in the operation of the system were made in an attempt to prevent the catalyst particles from adhering together. A significant improvement occurred when the hydrogen to carbon monoxide ratio in the Teed gas was increased. With a 1.3:l hydrogen to carbon monoxide feed gas, it was possible to reduce the rate of catalyst cementation until a useful catalyst life of 4 to 5 months could be attained. However, it was still quite difficult to remove the catalyst at the end of the experiment. The use of a hydrogen-rich feed gas also tended to increase the methane yield, especially in a two-stage process. A new method pf operation-the moving catalyst bed-was developed to eliminate this cementation problem. Catalyst particles (10 to 40 mesh) were charged to the converter for moving bed operation, and the linear velocity of the cooling oil was increased to such a value that the catalyst bed expanded until the bed height was about 25 to 35% greater than its settled height. Considerable motion of the individual particles resulted from the high rate of oil circulation. A moving bed of synthetic ammonia-type catalyst was operated successfully in this manner for several months without any increase in the pressure drop across the catalyst and entirely without cementation of the bed. Although considerable attrition of the catalyst occurred, even after 4 months of operation this had little or no effect on the activity, and the catalyst carry-over from the converter to the oil circulating lines wasnegligible. A patent search has revealed that the general principles of the moving bed system have been previously described (8). Other advantages of the moving bed as compared to the fixed bed are: 1. Greater catalyst economy is achieved by the use of smaller catalyst particlns with greater geometric surface area per unit weight; this rusulta in greater conversion per pound of catalyst at the aame conditions or synthesis. 2. Lower operating temperatures are possible for the aame amount of conversion. 3. Charging or withdrawing catalyst i s facilitated both during operation or on termination of a run. 4. Since the preasure drop amow the moving bed is almost independent of the rate of oil circulation (within the range of flows employed), lesa energy is required for a given temperature differential in the moving bed than for the eame differential in the fixed bed. 5. A greater spacetime yield based on converter volume aa well aa catulyst volume is obtained.
On the basis of the many improvements introduced by the moving bed system, fixed bed experiments were terminated, and further studies were directed toward investigating operating and yield data for the moving bed-type operation.
(1)
Thus, b variation of the gas recycle rate, it was osuible to change t i e hydrogen to carbon monoxide usage ratio &roportion of hydrogen to carbon monoxide consumed 111 the synthesis of the synthesis gas within certain limits. The usage ratio could t en be made more nearly equal to the hydrogen to carbon monoxide ratio in the feed gas and, thus, more complete consumption of the synthesis gas was effected.
L
Both granular and pelleted precipitated iron catalyats were used in the initial experiments at the Bureau of Mines. These catalysts were active, but after relatively short periods of opera-
EQUIPMENT
The converter was a 3-inch Schedule40 iron pipe, 10 feet 6 inches long (Figure l), accommodating a catalyst settled bed height up to 8 feet. Normal bed expansion in the moving bed WM usually maintained a t 30% of the settled height, although bed expansions of 20 to 50% gave good catalyst motion with no solids carry-over a t the higher liquid rates. Figure 2 shows the relationship between superficial linear velocity of the oil and the resultant bed expansion: this relationship was obtained in a
INDUSTRIAL AND ENGINEERING CHEMISTRY
Figure 1.
Pilot Plant for Moving-Bed Experiments
Vol. 42, No. 11
November 1950 72
INDUSTRIAL AND ENGINEERING CHEMISTRY
adsorbers, each of 1-cubic-foot capacity and capable of an adsorption cycle of 8 hours. The out gas, D, was metered as it left the charcoal system and was sampled for analysis. The material retained by the charcoal was stripped by means of superheated steam. The heavier fraction was condensed by a refrigerated bath and the lighter portion was metered and collected in a gasholder.
50.0
a W
m 0 W
68
5
t
%
w 64
E X
37.5
8a
.
60
w
Z 25.0
RAW MATERIALS
dz
I
-s
,-- 56 row $W
12.5 52
W 0
m
48
2379
0 0
0.06
0.12
OB
OIL LINEAR VELOCITY, FEET PER SECOND
Figure 2. Extrapolated Plot of Catalyst Bed Height us, Oil Linear Velocity i n Converter No. 2
small glass tube operated to simulate conditions in the converter. The results were subsequently verified in the operation of the converter itself. During the synthesis runs, an indication of the height of the catalyst was obtained by the use of a carefully weighted ball float of a specific gravity such that it would sink in the clear oil but float on top of the moving catalyst. A light rod attached to the top of the ball extended to the upper portion of the converter into a transparent sight glass. The rate of oil circulation was adjusted so that the top of the extension rod was maintained a t the proper height, indicating the position of the catalyst-oil intwface. Control of the desired bed expansion was simple and offered no operational difficulty. A schematic flow sheet of the system is given in Figure 3; in order to simplify tracing the streams, symbols have been assigned to them. Cold recycle gas Ct was mixed with fresh synthesis gas A and preheated to reaction temperature by means of an electrical resistance heater. The mixture was fed into the bottom of the converter along with the stream.of cooling oil. Because of the heat losses inherent in small pilot plants, it was necessary to preheat the recycle oil to converter temperature by means of a small Dowtherm heater. In large scale operation, cooling of the oil stream would be required, as approximately 90% of the heat of reaction is removed by the circulating oil. This heat could be recovered by generating high pressure steam, which could then be used for process heating and other energy requirements. The products of the reaction, unreacted synthesis gas, and the circulating oil passed from the converter to the gas separator (a 6-inch, Schedule40 pipe, 24 inches long), where the gaseous products and circulating oil were separated. The gas separator acted also as a sump, where the excess circulating oil was drained off periodically as heavy-oil product to maintain a constant level in the converter. The circulating oil flowed through two parallel l w m e s h strainers, which were examined for catalyst carry-over a t regular intervals. The circulating oil was then returned to the converter by the pump. The lighter reaction products in the form of gases and vapors, some vaporized circulating oil, and the unreacted synthesis gas passed upward through the condenser where the stream was cooled; the uncondensed portions were passed through a backpressure regulator to the charcoal recovery system. Simultaneously, the heavier hydrocarbons and water were condensed, the water was decanted, and the hydrocarbons were refluxed to the converter to aid in maintaining the proper level. The activated charcoal system removed the Ca and heavier gaseous hydrocarbons. This system consisted of two parallel
Catalyst Properties and Reduction. Iron catalysts of the synthetic ammonia-type were used in the submerged bed system. However, the use of other less expensive catalysts in this system is being investigated. In the fixed bed, particle sizes of 2 to 4 and 4 to 6 mesh were suitable, whereas in the moving bed 8- to 16- and 20- to 42-mesh particles were satisfactory. It is possible that finer mesh sizes may be more advantageous. Because catalyst activity depends essentially on the amount of available surface, other factors being equal, a calculation of the relative geometric surfaces of the various catalyst fractions shows the advantage of the finer mesh sizes, such a8 those used in the moving bed (Table I). The bulk density of the catalyst was about 150 to 160 pounds per cubic foot. Considerable reduction in the size of the particles undoubtedly occurred by attrition during the synthesis, but catalyst losses remained small. For example, a charge of 8- to 16-mesh catalyst, after 4 months of operation, showed the sieve analysis given in Table 11. I t was estimated that the catalyst lost by carry-over during this period did not exceed 1%. Before the catalyst was charged to the converter, it was reduced by substantially dry hydrogen in a separate unit. Operating conditions for the reduction were: 460 Atmospheric 48-72 2000 1 8
0 3-0.4
TABLE r. RELATIVE SURFACE AREASOF SYNTHETIC AMMONIATYPECATALYST
TABLE11.
Mesh
Relative Surface Area
2-4 4-6 8-16 20-42
1.0 2.0 5.9 14.5
hEVE
ANALYSES O F RAW AND USED SYNTHETIC AMMONIACATALYST
Original Unreduced Catalyst Weight' Mesh %
Same Catalyst after 4 Months of Synthesis Weight Mesh %
8-10 10-14 14-16 > 16
8-10 10-16 16-24 24-32 32-42 >42
... ...
40.9 42.0 14.5 2.6
... ...
13.5 23.8 18.7 32.9 9.0 2.1
_ I
100.0
100.0
TABLE111. RAW AND REDUCED CATALYST ANALYSES Before Reduotion, % FeaO4 93.51
% i004 ... 066 351 MnsOa KZ0
Si02
0.56 0.64
After Reduction, % ' Total iron 89.2 Reduotion of the iron 95.5 Loss of weight on reduction 2 4 . 2
INDUSTRIAL AND ENGINEERING CHEMISTRY
2380 Water condenser
Vol. 42, No. 11
Refrigerated condenser
A cooling-oil flow rate of 30 to 50 gallons per hour was required with fixed bed operation to maintain e temperature differentialof 15' to 20°C. Reflux oil regulator across an 8-foot bed at 70% conversion of the synthesis gas and a fresh-gas space velocity per hour of 300. The ratio of the rate of oil circulation to the rate of oil production was about 240 through 400: 1. With the moving bed, the super:(D) ficial linear velocity of the oil had to be about 0.10 to 0.15 foot per second to keep the catalyst particles moving in a satisfactory manner. This velocity was equivalent to a rate of about 120 to 150 gallons per hour and, as a result, the temperature differential across a 4-foot settled catalyst bed was decreased to 4" to 6" C. for condiGas holder tions and results corresponding to those given for an 8-foot fixed catalyst bed. The ratio of the rate of oil circulation to the rate of oil production in this operation was about 960 t o 1200:l. However, this ratio could be preheater preheater greatly decreased by operating with a greater depth of catalyst and larger Figure 3. Bureau of Mines Fisoher-Tropsch Oil Circulation Process temperature differential, because the required circulating-oil velocity would The reduction was continued until the water concentration of the remain constant while the amount of oil production would increase. exit gas approached that of the inlet gas. Analyses of a typical OPERATION catalyst are given in Table 111. Following the reduction treatment, the catalyst was cooled to Induction of Catalyst. The importance of bringing a fresh room temperature in a hydrogen atmosphere. It was discharged catalyst on stream properly was stressed by many German inand stored under carbon dioxide until ready for use in the convestigators (84). To a large extent, the catalyst activity and verter. ultimate catalyst life are determined by the method of inducting Synthesis Gas. Synthesis gas for the experiments was prothe catalyst. Conditions for the optimum induction procedure duced and furnished by the Bureau of Mines gas plant a t Brucefor the submerged bed system are still being investigated at the ton, Pa,, using the Girdler hydrocarbon reforming process. This Bureau of Mines. The method of induction employed in the process employs natural gas, steam, and carbon dioxide as raw experiments reported here was as follows: materials and produces gases with hydrogen to carbon monoxide The coolin oil was charged to the proper level in the converter ratios ranging from 1O:l to 1:4. The hydrogen plus carbon through w h i z a nitrogen flow was being maintained. Then the monoxide content of the synthesis gases usually exceeded 99%, catalyst, which had been stored under carbon dioxide, was added, with fractions of 1% of methane, carbon dioxide, and nitrogen the converter was sealed and the oil circulation rate was adjusted present. The sulfur content was less than 0.1 grain per 100 cubic to the required value. $he nitrogen flow was then interrupted, and the normal flow rate of synthesis gas was set at a space v e feet, The synthesis gases used during these experiments had locity per hour of GOO for moving bed experiments and 300 for fixed hydrogen to carbon monoxide ratios of 1:1, 1.3:1, and 2:1, bed experiments (based on the volume of the settled bed). After respectively. Special tests also were run with a feed gas consistthe system was pur ed of nitrogen by the incoming synthesis gas, ing of 50% hydrogen, 25% carbon monoxide, and 25% carbon the unit was brougtt up to full operating pressure (usually 300 ounds per square inch gage), and recycling of tail gas was started. dioxide in order to simulate second-stage operation. %eat was then applied to the circulating oil. A stepwiseinduction Cooling Oil. A cut of Diesel oil from the synthetic fuels pIant procedure was pursued in which the converter tcmperaturc was a t Harnes, France, was used as coolant in the initial experiments. steadily raised from an initial conversion temperature of about Subsequently, in starting a run, oils of the desired boiling range 150" C. until a 30% conversion of the synthesis gas was reached (around 210° C,); this conversion was held for 24 hours. In a (300"to 450" C.) were taken from the products of the preceding like manner the conversion was further raised to and maintained experiments. However, after a few days of synthesis, most of for 24 hours at 45 and 60% levels, respectively, by increasing the the original charge of oil appeared to have been replaced, and operating temperature in the required steps. Continuing this steady state was reached in which the composition of the coolant schedule, there wag attained on the fourth day, (b normal synthesis gas conversion of SO%, which was held for the duration of the depended on the prevailing operating conditions and the distributest. During experiment 19, with a moving bed, the induction tion of the products formed in the converter. Analysis of a typiwas carried out as follows: cal '(equilibrium" recycle oil (heavy oil) as found in experiment 19 HsfCO Average Temp. Conversion, Tynz., Differenshowed the following characteristics : % tial, C. Date
-
Gravity. OA.P.I., 80' F. 41.3 A.B.T.M, Distillation, ' C. l e t drop
6%
206 23 1 253 286
A.B.T.M. Distillation, 320 353
ass
460
' c.
May May May May
18. 19, 20, 21,
1949 1949 1949 1949
30.8 48.3 59.8 87.9
213 226 235 239
4 5 6 7
Synthesis. Normal operation followed the induction, and only minor changes in the operating variables were required from time to time to maintain the desired conversion. Generally, as the
INDUSTRIAL AND ENGINEERING CHEMISTRY
November 1950
TABLEIV.
CO&fPARISON
When it appeared that the catalyst activity was declining rapidly, the flow of synthesis gas was shut off, and hydrogen waa passed through the converter in an attempt to reactivate the catalyst. The first hydrogen treatment was effective in increasing the activity, but subsequent ones were of decreasing value unless higher treatment temperatures were employed. Ordinarily, a hydrogen treatment was required after about 1 or 2 months of synthesis. These treatments were generally carried out with the catalyst submerged in the circulating oil at 250' to 300' C., 300 pounds per square inch gage, and total hydrogen space velocities (fresh and recycle hydrogen) of ahout 600 to 800 over a 24hour period. Water and methane were evolved, the former from the reduction of the iron oxide and the latter from the reaction between the hydrogen and the iron carbide. In a few caws the cooling oil was drained and the catalyst was subjected to a more severe hydrogen treatment, at 450" C., which brought about a virtually complete rereduction.
OF O P E R O I N J ~ CoNDlTloNs I N OIL-
SUBMERGED FIXED AND MOVING BEDS
Experiment No.
(All operation at 300 lb./aq. inoh gage) 14 17 19 21 ~ o v i n ~ ,Moving,
Catalyst bed
Weightof ofHI: iron eddgae lb. Ratio COohar in & Total aatalyst age, hr. Temperature, 'pax.&a. C. Temperature did.. C. Pressure Start ofdrop, expenment lb./aq. inah E n d of experiment Feed gas rate, stand. ou. ft./
hr.
Fixed. bFt 4-6" Mesh 42.8 &O;
264-266 I6 4 49 124
StfAkiO 4-Ft. s4-F4. tat10 Fixed, Hei ht Rei ht 8-Ft 6.2.h: 6.25t: 4-6" Expanded, Expanded Mesh s-16 20-42 147.6 .3:l 124.7 .3:l 21.4 1:l 2300 2670 760 2.56-266 244-249 249-264 16 6 6 4 6 6 16 6 6 123.6 323.2 123
Spaoe and velocity pres&) (stand. vol./vol./ temp.
hr.
Baaed on settled aatalyat Based on expanded oatalyst Space wei ht velooit H*+Cbfeed ly.&.?ift* Reoyale ratio h a d gas: fresh
sad
Conversions
co % &%o
%
Ueage rdtio (Hr:CO He+CO: oonverted)hr./lb. Fe ou ft. Ci+bn, &ou. meter of feed
300
.. .
302
...
601 482
600
2.88
2.66
4.97
6.74
2.0:l
1.O:l
1.O:l
69.9 69.9 69.9 1.0 2.01 24.6
gas
676.9 4.2 69.2 1.1 1.76 26.3
461
PRODUCT YIELDS
63.0 79.8 70.3 1.04
6716 .O .8 68.4 0.93
3.49
3.98 21.6
catalyst aged, the hydrogen to carbon monoxide usage ratio dropped slowly, with a simultaneous decline in the catalyst activity, so that it became necessary to increase the operating temperature graduauy to maintain the conversion. Comp&on of operating data for some of the fixed and moving bed experimenta is given in Table IV. Experiments 17, 19, and 21 were terminated voluntarily, but experiment 14 had to be discontinued because of the excessive pressure drop. Table I V shows the advantage of the moving bed over the k e d bed in regard to catalyst economy. Thus, although 3.49 cubic feet of hydrogen plus carbon monoxide were converted per pound of iron in experiment 19, compared to 1.76 cubic feet in experiment 17, a higher operating temperature was not needed for experiment 19 a~ might have been expected, and it was possible to operate at 8 temperature which was actually IO' to 15" C. lower. In addition, a t the given conditions, less CI and Cz were produced in the moving bed experiment. It usually became necessary, as the run progressed normally, to increase the operating temperature about 1" to 2" C. per week. TABLE v. A Fresh Feed)
compo. 29iO Cu. Ft./day nent COY CO H~
Vol. % 0.10 48.03 66.46
CI
0.28
cs cr
... ...
Cr
ci c 4
Ci Cb Ci
CC
Ci
Nt
ct
...
... ... ... ... I . .
I . .
&is . I *
Lb./day 0.37 99.6 9.88
0.36
,..
... ... ... ,..
... ...
... ... 6:ba 6..
&,
am
U o m t r d for rmounb In ford gam.
0 .
e..
TABLE VI. PRODUCT YIELDSOF VARIOUSCOMPONIDNTS (ExptR. 14, 17, and 19, 70% conversion; expt. 21, 138.4%oonversion) Fixed Bed Expts. Moving Bad Expts. (Settled Bed Space (Settled Bed Spaae V+city. 300) Velooity, 600) Component No 14 No 17 . No 19 No. 21 (From S t r e p a and 1: l&:dO 1.3:1 H Y : ~ O 1.3: 1 ",:bo 1: 1 H*:Cb Charcoal Spirits) -8peoifio yields, g./ou. m. of feed gas13.3 13.9 10.3 10.7 c& C,Hs 9 2 10.2 7 6 8.1 2.0 2.2 4.0 2.7 6.3 6.4 3 7 4.1 10.0 13.2 13.1 12.7 3.1 3.6 3 4 3.6 CIHa 7.4 9.1 9 3 9.2 1.6 1.6 2.0 1.4 4 4.2 3 . 4 5 . 9 3. .. 7 0.2 0.7 1.7 ... 2.6 3.6 2.7 Cl+ in aharooal spirits 2.4 Light oil 30.9 30.8 31.8 39.2 Heavy oil 37.4 36.3 42.6 38.4 Total 124.4 134.1 139.1 186~5 117.2 116.0 99.9 107.8 C1-k Data not available on oharooal recoveries.
2g %$o
23:
e..
... ...
;:2
COMPOSITION OF G A 8
B CI (Rea ole Gas) (Inlet Gas) 6920 Cu. Ft./Day 2960 Ft./Dby Vol. % Lb./day' Vol. % Lb./day 12.66 92.6 26.22 92.1 31.23 144.6 19.43 44.9 61.06 18.97 46.66 7.59 1.97 6.24 3.68 4.89 0.60 3.01 1.20 8-01 0.34 1.58 0.68 1.68 0.20 1.49 0.40 1.49 0.76 6.27 1.49 6.27 0.14 1.36 0.28 1.86 0.41 3.64 0.81 3.04 0.06 0.78 0.18 0.78 0.21 2-37 o,41 0.005 0.07 0.01 0.07 0.06 0.69 0.10 0.60 0.83 1.01 0.60 0.66 ..I
Product streams collected (Figure 3) and analyzed were the charcoal spirits, light oil, heavy oil, and the water layer, which contained most of the water-soluble oxygenated compounds. The charcoal spirits consisted of the portion of hydrocarbons from the tail gas adsorbed on activated charcoal; this was later steam-stripped, and the hydrocarbons refrigerated to about 5 O to 10" C. The light oil and water layers were collected in the refrigerated trap which was maintained a t about 5' to 10" C. The measured gas streams (Figure 3) were the fresh Synthesis
1.O:l
21.9
2381
(1
STREAMS ca
IN
EXPIBRIM~ENT 19
(Tail gas), 1362 Cu. Ft./Day G./ Vol. % 26.22 19.43 46.66 8.68 1.20 0.68 0.40 1.49 0.18 0.81 0.13 0.41
Lb./day 42.3O 20.66 3.48 2.240 1.39 0.73 0.69 2.42 0.62 1.72
0.01
o,oa
0.10
0.32 0.81
o m .I.
ocrUie3 gas 227.0 ,
..
0.88
i6:3 7.6 4.0 3.7 13.1 3.4 9.3 2.0
1.08
6.9
...
0.2 1.7
... ..a
D (Out Gas), 1329 Cu. Ft./Day G./ au. m. of feed Vol. % Lb./day gaa 26.45 41.7 226.6 19.90 20.66 ,,, 46.81 3.48 3.77 2.24 i0:3 1.23 1.39 7.6 0.70 0.73 4.0 0.31 0.62 2.8 1.19 1.89 10.2
...
...
...
0.13
0.27
...1 . 4 ...
... ,,. ,.. $81 ...
...
... 0 .
6:ii
.*.
. . I
::: .I.
Charcoal Spirits Plus Adsorbod Gases 84.9 Cu. Ft./Da; G./ 011. m. of feed Vol. % Lb./day gna 14.30 0.62 ,
...
... ... ...
..
..... ..
. . I .
3:i3 12.70 10.92 26.66 6.07 16.00 0.40 8.90 ..I
6.28
o:i7 0.63 0.62 1.46 0.36 1.09 0.03 0.82
2.9 3.4 7.9 2.0 6.9 0.2 1.7
0.67
3.6
*.
0.9
I . ,
2382
INDUSTRIAL AND ENGINEERING CHEMISTRY
TABLSVII. PRODUCT STREAMRECOVERIESAND INFRARED ANALYSES FOR OXYGENATED GROUPS AND OLEFIN CONTENT (Experiment 19) Weirrht
Charcoal spirits Light oil Heavy oil Molecular weight of functional groupa
5.07 6.08 7.90
0.65 0.75 0.85
*.
11.2 3.8 0.34
3.7 3.9 3.0
0.30 0.58 0.37
0.05 0.14 0.34
0.11 2.76 0.00
99 52 22
..
45 COOH 44 17 24 (28 co a Aldehyde contents are uncertain and are partly included under COOH and CO. b Results are given in weight of functional y u ~ for : actual percentages, these values should be multiplied by factor of assume?molecular weig t o compound. molecular weight of functional group ,,
24
gas feed; the inlet gas; the recycle gas and tail gas which were
identical in composition; the out gas; and the charcoal adsorbed gases. Table V shows the composition of the gas streams of experiment 19 in which a moving bed with 1.3:l hydrogen to carbon monoxide feed gas ratio, 70% gas conversion, 1: 1 gas recycle ratio, and an expanded bed space velocity per hour of 460 were used. Derived specific yields of the various hydrocarbon components for experiments 14, 17, 19, and 21 are given in Table VI so that fixed and moving bed operation can be compared. The yields for experiment 19 were computed from data on the gas streams given in Table V and from the weights and analyses of the recovered liquid-product streams (Table VII). The percentages of a- and &olefins and of the various oxygenated groups (alcohols OH, esters COO, acids COOH, and ketones CO) in the oil streams were determined by infrared analysis; the oil yields given in Table VI were corrected for thebremoval of these oxygenated groups assuming that there was no loss of hydrocarbon. The individual Cq, C6, and C6 yields given in Table VI are only for the portions of these components in the gas stream and the rharcoal spirits and do not include the amounts in the light oil. The yield data show that somewhat lower C, and C2 yields (21.9 and 21.5 grams per cubic meter of feed gas) were obtained in the moving bed experiments as compared to values of 24.5 and 26.3 grams per cubic meter in the fixed bed experiments, 14 and 17. The higher C$ yield which resulted in the moving bed experiments is a decided economic advantage. In all experiments the olefinic content of the C( to c6 components exceeded that of the corresponding paraffins; conversely the ethylene yield was lower than that of the ethane. C1 and 12%Yields at Varying Recycle Ratios. A series of tests (Table VIII) was carried out in the moving bed system to determine the effect of the gas recycle on the Ct, Ct, and C, production and the hydrogen to carbon monoxide usage ratio. In these tests the fresh synthesis gas Row rate was fixed a t a settled bed space velocity per hour of 300, conversion was held a t 70% by adjusting the operating temperature, and the gas recycle ratio was varied. The gas recycle ratio was increased in steps from
Vol. 42, No. 11
1:1 to 6: 1 and then returned to 1:1 so that a check could be made. The results given in Table VI11 show that as the gas recycle ratio was increased, the CI and C, production decreased and the hydrogen to carbon monoxide consumption rose. Both these factors are desirable for efficient conversion of the synthesis gas. However, economic analysis is necessary to determine whether these advantages are offset by the increased cost of the higher recycle. SIMULATED SECOND-STAGE OPERATION
Because experiments 14. 17, 19. and 21 were operated a t a 70% conversion of the synthesis gas, a value considered a t that time a8 optimum for the first stage, investigation of a second synthesis stage was necessary to obtain operating and yield data on the conversion of the unreacted hydrogen and carbon monoxide in the tail gas. A 70% conversion in the second stage would give an over-all usage of 91% of the original synthesis gas. When a 1: 1 hydrogen to carbon monoxide feed gas was used in the synthesis with a proper recycle ratio, a tail gas with hydrogen t o carbon monoxide ratio of substantially 1: 1 was obtained. Since after removal of carbon dioxide, the resulting gas would be similar t o the original feed, the second stage would be expected to be the same as the first. However, with the use of a 1.3: 1hydrogen to carbon monoxide fresh feed gas, since only a 1.1:1 hydrogen to carbon monoxide usage was obtained a t 70% conversion, the hydrogen t o carbon monoxide ratio of the resulting tail gas was about 2: 1. In order to simulate a second stage in this case, a 2: 1 hydrogen to carbon monoxide feed gas was used in part of experiment 19. In another part of this experiment the same gas with 26% carbon dioxide added was fed, so that operation of a second stage was simulated in which the carbon dioxide formed in the first stage was not removed. A comparison of the results from these tests showing the effect on the C1 and CZ yields, conversion, and usage ratio of the 2:1 hydrogen to carbon monoxide gas, with and without carbon dioxide added, is given in Table IX. The data presented in Table I X show that: Using a constant 1:1 gas recycle ratio: 1. The hydrogen t o carbon monoxide usage ratio increased from 1.04: 1 to 1.45: 1 as the feed gas ratio waa increased from 1.3:1 to 2 : l (with no carbon dioxide added). 2. When the 2: 1 feed gas contained 25% carbon dioxide, the usage ratio increased to 1.63 (267' C. maximum temperature) ELO compared to 1.45: 1 for the 2: 1 feed gas without carbon dioxide added (251' C., maximum temperature). This indicated that the
TABLEIX. COMPARISON BETWEEN FIRSFSTAGEAND SIMULATED SECOND-STbGE OPERATION Simulated Second
AND USAGERATIOWITH TABLEVIII, C,, Cz PRODUCTION VARIANTRECYCLERATIO Av. CI. cz,,c;
Recycle Ratioa, Tail Gas/Fresh Gas
a
Temp., Reqd. for 70% Conversion,
C,Oz in Tail Gas,
Usage Ratio, c. % Hz:CO 27.4 0.989 1.03 25.1 1.16 23.4 22.9 1.13 21.5 1.15 20.3 1.22 26.6 1.02 Feed gas flow rate 61.5 cu. ft./hr.; 300 settled bed space
S ecific eield G . / c ~'M. . Feed Gas (Hz CO) 21.5 22.1 20.9 18.2 15.7 14.2 21.4 velocity.
+
Simulated" Second
(C8ze (257 Stage Cot Sta First Stage Removed) Preient) Gas feed (Hz:CO) ratio 1.3:l 2:l 2:l COa in feed, % 0.1 0.1 25.0 Usage (Hz: CO) ratio 1.04:l 1.45:l 1.63:l Gas feed (Hs+CO), CU. ft./hr. (NTP b ) 123.2 123.0 123.0 COz in feed gas, cu. ft./hr. (NTPb) .... .... 41.0 Recycle ratio (tail gas: fresh gas) 1:l 1:l 1:l H2 -I-CO conversion, % 70.3 70.8 66.4 HZ-I-CO aonverted, cu. ft./hr. 86.6 87.0 81.6 CI, Cz, C- specific yield, g./ CO converted 32.1 49.0 50.1 cu. m. 7% Temperature (max.), C. 249 251 267 Feed gas contained approximately 50% Hz, 25% CO,and 25% COa b Normal temperature and pressure.
+
November 1950
INDUSTRIAL A N D ENGINEERING CHEMISTRY
shift of carbon monoxide to dioxide was retarded by the presence of the carbon dioxide in the feed gas. For approximately the same hydrogen plus carbon monoxide conversion (70%): 1. The tem erature required for 1.3:l and 2:l hydrogen to carbon monoxije feed gas without carbon dioxide was essentially the same (249' and 251' C.). 2. When 25% of carbon dioxide was added to the 2: 1 feed gas, the re uired temperature for 70% conversion of the same amount of hy%ogen plus carbon monoxide, as in the other two caaes, incremed-to 267' C. ' was essen3. The combined specific ield of CI,CZ,and C tially the same for 2 :1 gas witxout and with carbon &oxide added (49.0 and 60.1 grams per cubic meter of hydrogen plus carbon monoxide converted, respectively) ; i t was appreciably higher, however, than the value of 32.1 grams per cubic meter of hydrogen plus carbon monoxide converted obtained with 1.3:1h dro. an over-all fimre of36.1 pen to carbon monoxide feed p ~ Thus, grams per cubic meter of hydrogen plus carbon mo6oxide converted may be calculated for 91% conversion of 1.3: 1 hydrogen to carbon monoxide feed gas in two stages as described.
Although the use of 2: 1 hydrogen to carbon monoxide feed gas with 25% carbon dioxide increased the usage ratio to 1.63: 1, an appreciable decline in the catalyst activity was noticed after only 8 days of such operation. Before and after the carbon dioxide addition, similar tests with 2: 1 hydrogen to carbon monoxide feed gas showed that an increase in the maximum operating temperature from 252" to 269" C.was necessary to maintain 70% conversion. Since such high carbon dioxide concentrations in the feed gas produced permanent deterioration in the catalyst activity, operation with this gas is probably undesirable even though it depresses carbon dioxide formation in the synthesis. Because of the high methane production in the second stage of a two-stage synthesis using a 1.3 :1 hydrogen to carbon monoxide feed, use of a 1:1hydrogen to carbon monoxide feed gas is more desirable. Operation with 1:1 hydrogen to carbon monoxide gas in the first stage has been found satisfactory in the moving bed system, and in a two-stage process only about 30 grama of CI, CZ,C; per cubic meter of hydrogen plus carbon monoxide converted should be produced. Also it is less expensive to produce synthesis gas of about 1:1 hydrogen to carbon monoxide ratio than a hydrogen-rich gas when coal is used as a raw material. CHARACTERISTICS OF PRODUCTS
Properties of Crude Streams. The product streams recovered were the charcoal spirits, light oil, heavy oil, and the water layer. Some of the physical properties of the oil s t r e a m found in experiment 19, using a 1.3:1 feed gas and 1:1 gas recycle, are shown in Table X. There was no appreciable difference in the produats when 1 :1 hydrogen to carbon monoxide feed gas was used. Properties of Product Water Layer. The water layer decanted from the product oil streams contained several oxygenated compounds, chiefly alcohols. A typical analysis of the product water from experiment 19 is given in Table XI. The weight of oxy-
TABLE X. PROPERTIES OF OIL STREAMS AS RECOVERED Stream No. (Figure 3) Gravity "A.P.I., 60° F. A.S.T.M. distillation, C. 1st drop
40
Viscosity, centistokes at 120' C. Bromine Ne., y"00 8. oil Acid No., mg. OH/g.
Charcoal Spirits (1) 70.6
Li&t (2) 62.4
Heavy Oil (3) 41.3
... ...
TABLE XI. MASSSPECTROMETER ANALYSIS OF PRODUCT WATER LlYER
(Experiment 19) Component Weight % CHIOH 0.7 CiHsOH 2.5 0.8 n-CaHvOH n-C4HoOH 0.5 Acetic acid 0.9 Acetone 0.1 94.5 Water 100.0
-
TABLE XIT. LIQUIDPRODUCT DISTRIBUTION FOUND IN PILOT PLANT OPERATIONS (600 apace velocity)
Values. Weight % Expt 19 Ex t 21 (1.3:l H2:CO) ( 1 : l %:CO)
Product
genated compounds produced in the synthesis in all streams is about 10 to 12% of the total C$ hydrocarbons plus oxygenated compound6 production. Product Distribution. Typical product distributions, including total oxygenated compounds obtained in experiments 19 and 21, are given in Table XII. Apparently the differencein feed gas within the limits shown did not affect the product distribution. In a brief test the gasoline yield was increased by 8 to 10% without materially affecting the methane production by increasing the hourly space velocity from 600 to 1OOO. The heavy distillate and wax cuts can readily be cracked to the Diesel or gasoline ranges. Table XI11 shows the calculated yield of hydrocarbons and oxygenated organic compounds per lo00 cubic feet of synthesis gas. A conversion of 96% of the synthesis gas in two stages is assumed. Column 2 shows the results of thermal cracking of the heavy ends mainly to Diesel oil, and column 3 gives the yield on a total cracking tb gasoline. The yields on cracking were computed from published data (8,11). Product Ratings. Octane ratings and Reid vapor pressures of the following gasoline samples are reported in Table XIV: raw gasoline from the synthesis (C~-40Oo F.); this same material after reforming over bauxite catalyst at 350" C. and 25 pounds per square inch gage to remove the oxygenated compounds and
TABLE XIII. PRODUCT YIELD PER 1000 CUBIC FEETOF HYDROGEN PLUSCARBON MONOXIDE (Experiment 19, 95% conversion) Gasoline Primary and Diesel Products Oil Only CI,C,, C', lb. 1.86 2.14 Gasoline a Ci-4OO0 F 1, gal 1.012 1.069 Diesel oii ~40O-6OO0F ) ai. 0.187 0.515 . Heavy distillate (600-g2% F.) gal. 0.211 Wax (>842O F.) gal. 0.189 ... Oxygenated compounds, Ib. 1.31 ...
...
Gaaoline only 2.61 1.525
... ... ... ...
TABLE XIV. OCTANE RATINGSAND REIDVAPORPRESSURES OF GASOLINE SAMPLES
29 33 40 52 67 82 97 160 191
99 0.8
2383
Description Raw gasoline (Cc-40O0 F.) Reformed gaaolme Clear Plus 1 ml. TEL Plua 3 ml. TEL
bi:8 2.1
2.84 22.8 0.41
R.V.P., Lb. 8.1 9.3
...
... ... 9.9
... ... ...
A.S.T.M. Motor Octane No. 50.6 67:2 78.8 81.6 77 1 84 8 research
83:6 [91:2 research]
86.8
2384
INDUSTRIAL AND ENGINEERING CHEMISTRY
to shift the double bond from the
a to an internal position; a final blend (9.9 pounds, Reid vapor pressure) consisting of 65.6% by weight of reformed gasoline, 31.9y0of polymer from the unsaturated Ca and C, components in the out gases, and 3.5’5 of butane. Table XIV shows the improvement in octane rating effected both by the reforming treatment and by the addition of tetraethyllead. The octane rating of the final leaded gwoline blend is about the same as that reported by Bruner ( 4 ) for Hydrocol gaaoline from the fluidized process. The character of the Diesel oil as produced is indicated by the following data: Cetane rating 6 clfio gravity
Goisity Pour point Flrrah point
( 0 en oup) Conredson ,argon on 10% bottoms
82 .O $blending value) 47.6 A.P.I. at 60’ F. 3 4 . 8 S.U.9. et 100’ F. +20° F. 2 0 8 O F.
Trace
Although the pour point is high, it can be lowered by blending with the cracked Diesel fraction obtained from the cracking of the hertvy distillate and wax. SUMMARY A N D CONCLUSIONS
I
Operation of the Fischer-Tropsch synthesis, using a bed of iron catalyst granules through which cooling oil was circulated at a rate sufficientto expand the bed and impart motion to the catalyst particles, successfully overcame the obstacles encountered when a stationary or fixed bed was used. Data are given for experiments in which a 91% over-all synthesis gas conversion, using 1.3: 1 hydrogen to carbon monoxide feedgas, was attained by operating a two-stage process with 70% conversion in each stage. Second&age operation was simulated by w e of the proper feed gas. Both stages were operated at a settled bed space velocity per hour of 600, 240’ to 260’ C., 300-pound-per-square-inch-gage pressure, and with a 1:1 recycle ratio of tail gas to fresh synthesis gas. A CI,Cz,and Ca specific yield of 36.1 grams per cubic meter of hydrogen plus carbon monoxide converted was obtained a t these conditions. The C1,Cp,and Cy yields were reduced by increasing the tail gas to fresh gas recycle ratio. An unexpected, rapid decline in catalyst activity occurred when 25% of carbon dioxide was added to the 2: 1 hydrogen to carbon monoxidc feed gas to simulate second-stage operation without carbon dioxide scrubbing. Data are also given for operation of this system with 1: 1 hydrogen to carbon monoxide feed gas. In this case it was not necessary to simulate a second stage of synthesis, for after removal of carbon dioxide from the tail gas, essentially a 1:1 hydrogen to carbon monoxide gas is obtained. No attempt was made to determine the maximum catalyst life which could be expected. However, one batch of catalyst was used for about 2700 houra at various operating conditions and still possessed considerable activity a t the end of that time. The chief advantages of the moving bed as compared to the fixed bed are as follows: greater catalyst economy, lower operating temperature for equal amounts of conversion, ease of charging or withdrawing the catalyst, reduced temperature differential across the catalyst bed without increased pressure drop, and greater space-time yield based on converter volume as well as catalyst volume. Detailed analyses and measurements of the gas and liquid streams, as well as the yield of various components calculated from the data, are tabulated. Other process information is presented concerning the physical properties of the product streams as recovered, typical product distributions, and evaIuation of the gasoline and Diesel oil produced. Calculated yields of products a t 95% synthesis gas conversion are tabulated. Some important advantages in operation of the Fiacher-Tropsch synthesis are offered by the moving bed oil circulation system. The prooess seems desirable beoause feed gases of epproximrttely I :I hydrogen to carbon monoxide ratio oan be used; these are readily and less expansively produced from coal than are eynthe-
Vol. 42, No. 11
sis gases high in hydrogen. Other advantages are the operability of the system and the relatively low quantities of oxygenated compounds and C1,Ca, and C; produced. When coal is used as a raw material, the cost of the purified synthesis gas has been estimated to be about 60to 70% of the direct cost of the synthesis. On this basis any increased efficiency in the synthesis gas utilization is of primary importance. Flexibility of the operation of the moving bed oil circulation system may offer the best possibility of achieving more economic synthesis gas usage. ACKNOWLEDGMENT
The authors wish to express their appreciation to the control laboratory, catalyst analysis group, and mass spectrometer section of the Bureau of Mines for the analytical data, and to the pilot plant group for the operating data. LITERATURE CITED
(1)Atwell, H.V.,U. 5. Patent 2,433,225(Dec. 23,1947). (2)Ibid., 2,438,029(March 16,1948) (3) Atwell, H.V.. and Schroeder, W. C., U. 8. Dept. Comm., OTS Rept., P B 367 or 412 (May 16, 1945). (4) Bruner, I“. H.,IND. ENQ.CHEM.,41,2511-15 (1949). (6) Duftschmid, F., Linckh, E., and Winkler, F., U. 8. Patent 2,159,077(May 23, 1939); 2,207,581(July 9, 1940); 2,287,092 (June 23, 1942); and 2,318.602 (May 11, 1943). (6) Faragher, W., and Foucher, J., Syntheses from CO Ha at I. G. Farbenindustrie, 1,part C, prepared at Ludwignhaven (1947). (7)Faragher, W. F., Horne, W. A., Howes, D. A., Schindler, H., Chaffce, C. C., Weet, H. L., and Rosenfeld, L.; U. S. Dept. Comm., 02’8 Rept., PB 1366 (1945). (8) FJAT reel No. K-23,L-70214. (9) FIAT reel No. K-31,frames 1108-16. (10) FIAT reel No. X-115,frames 1388-1447, 1693-1672. (11) Greensfelder, B. S.,Voge, H. H., and Good, G. M.. IND.ENO. CHEM.,41,2583(1949). (12)Hall, C. C.. Craxford, S. R., Gall, D., and Smith, 8. L., H.M. Stationery Office, London, BIOS Final Rept. 1712. (13) Hall, C. C., and Haenael, V., U. 9. Dept. Comm., OTS Rept., P B 415 (1945); Rept. 30,Filo 27-69 (January 1940). (14) Hall, C.C.. and Powell, A. R., Ibid., P B 286 (1945). (15) Horne, W. A. and Jones, J. P., Ibid., PB 294 (1946)[UrnAge, 97, No.3 (1946)l. (16)Johnson, E.A., U. S. Patent 2,393,909(Jan. 29, 1948). (17) Keith, P.C . , Chem. Eng. News,25, 1044 (1947); Petroleum Refiner. 26,171-2 (1947). (18) Keith, P. C., Natl. Am. Urn J . , 164, No. 6, 11-16 (1946); Petroleum News, 38,No.27,R506-11 (1946); Oil Gas J . , 45, No. G; 102, 105, 107, 108, 111, 112 (1946). (19) Latta, J. E.,and Walker, 8.W., Chem. Eno. Progress, 44, 173-6 (1 948). (20) Murphrec, E. V., Tyson, C. W., Campbell, D. H., and Martin, H. 2. (to Standard Catalytic Company), U. 9.Patent 2,360,787 (Oct. 17,1944). (21) Neumann, R., rind Sohroeder, W. C., U. S. Dept. Comm., OT5 Rept., P B 1279 (1945). (22) Peck, E. B.,U. S. Patent 2,161,974(June la, 1939). (23)Pier, M., and Michael, W., U. 8. Patcnt 2,167,004(July 26, 1939) (24)Reichl, E.R.,U.9.Dept. Comm., OTISRept., PB 22841 (1945). (25) Roberts, G.,Jr., and Phinney, J. A., Oil Uus J., 45, No. 46,41 72,73,139 (1947). (26) Rubin, L. C., U. S. Patent 2,448,279(Aug. 31, 1948). (27) Ruesell, R. P.,Chem. Eng. News, 25,86-7 (1947). (Mar. 31,1945). (28) Ryan, P.,Oil Uaa J., 43,264,267,268 (29) Scharmann, W.G.,U. 8. Patent, 2,451,879(Oct. 19,1948). (30) Sensel, E.E., Ibid., 2,411,760(Nov. 26, 19461. (31) Storch, H. H.,Golumbic, N., and Anderson R. B., “Fisoher-
+
I
Tropsch and Related Synthesis,” New York, John Wiley B Son, in press; Storch, H. H., “Advancer in Catalysis,” Vol. 1, pp. 115-53,New York,Academic Press Inc., 194% (32) Storch, H. H.,Powell, A. R.,and Atwell, H. V., U. 8. Dept. Comm., OT5 Rspt., P B 2051 (1946). (33) Tech. Oil Mission, films, Library of Congress, Washington, D. C., reel 13,bag 8043,item 2,frame 150. (34)Ibid., reo136,bap 3453,item 4-13, (36)IMd., reel 87,bag 3461,item 21. (36)I b M , reel 48, bag 8441,items 86 and 88. (37)Ibid., reel 47,bag 3446,item 98. (38)Ibid., reel 184,item I-a, No,8, item I-b, No. 11, 12,and 18. RBCBrV8D
May 4, 1960.
