Hydrogenation of Sha istillat e J
W
er
E. L. CLARK, R. W. MITESHUE, H. J. MANDINER U . S. Bureau of Mines, Bruceton, Pa.
BOYD MOKRIS W . S. Bureau of Mines, Rifle, Colo. T h i s work was undertaken to prepare a large quantity of jet and Diesel fuel from crude shale oil by recycle coking followed by hydrogenation of the coker distillate a t 1500 pounds per square inch gage pressure and 835" F. over a fixed-bed cobalt molybdate catalyst. In two hydrogenation runs, a t liquid hourly space velocities of 0.9 and 1.1, some 1500 gallons of product were obtained, corresponding to 96 to 989'0 volumetric liquid recovery, 1200 to 1400 cubic feet hydrogen absorption, and 300 to 400 cubic feet of hydrocarbon gases (Cl to C,) per barrel of feed coker distillate. Catalyst activity as evidenced by the degree of nitrogen and sulfur elimination as well as the general reduction in boiling range was highly satisfacEory, and runs in excess of 300 hours without regeneration were made. A simple batch distillation separated
the hydrogenate into two fractions-the 6Oq0 overhead material was the jet fuel made, and the bottoms were the Diesel fuel. The product jet and Diesel fuels met the salient requirements of the applicable specifications MIL-F-5624 and Navy 7-0-2e, respectively. The combination process explored here produced these usable products in a 75 to 85% over-all liquid yield, based on crude shale oil, and hence represents a possible technical solution to the shale oil refining problem. An integrated full scale plant for this operation may realize a by-product credit on the ammonia and hydrogen sulfide produced in the hydrogenation step (5.7 and 1.9 pounds per barrel of eoker distillate, respectively) as well a5 on the coke and gas of the recycle coking step. PREPARATION OF COKER DISTILLATE
HE methods of refining the oil obtained by retorting Colorado oil shale are being investigated on both a laboratory a n d pilot plant scale. Unique problems are encountered in the refining of shale oil. Chief among these is the presence of almost 2 weight % of nitrogen in the shale oil. This amount of nitrogen is sufficient t o impair the activity of commercial oil-cracking catalysts ( 7 ) . I n addition, the high pour point, high viscosity, a n d large proportion of high boiling fractions of this shale oil indicate that considerable molecular weight reduction (cracking) is necessary t o convert i t t o more valuable, lower boiling liquid fuel. One proposed solution of this refining problem was suggested by Reed and Berg (9) and mas further discussed by Thorne et al. (11); this method consists in the recycle coking of crude shale oil t o yield a distillate of specified end point, catalytic hydrogenation of this distillate, and fractionation of the hydrogenated Oil t o produce jet and Diesel fuel. Successful reforming of Santa Maria petroleum crude in a Pilot Plant in which this Procedure was used has been reported by Berg et al.( 1 ) . The Catalyst used in the hydrogenation was cobalt molybdate on alumina, of the general type reported by Byrnes, Bradley, and Lee ( 2 ) . This catalyst is operated a t 500 t o 1500 pounds per square inch gage with air regeneration after each process period ( 3 , 4). To test this catalyst and the procedure outlined, the Bureau of Mines coked a batch of shale oil in a continuous recycle coking unit and hydrogenated the resulting coker distillate in a continuously operating pilot plant.
The crude shale oil fed t o the coking unit was obtained by retorting Co]orado oil shale in a n T\T-T-T; retort. These retorts and their operation have been described (5,10). The facilities of the B~~~~ of ~i~~~~Demonstration plant Shale Oil &finery (6,8 ) were used for the coking operation. A diagram of the equipment for this operation is given in Figure 1. The total feed rate was 300 barrels per day with a heater-outlet temperature of 9400 F. and a 2 . 5 ~ 1recycle t o fresh feed ratio. The product distribution and operating conditions are given in ~ ~ 1. bThe l naphtha and light-oil fractions were combined t o give a distillate having an end point of 6900 to 7000 F. The charge stock and coker distillate are compared in Table 11. Inspectjon of the two materials indicates a considerable decrease in molecular weight due t o the thermal treatment, as evidenced by the lower pour point, viscosity, and boiling range and the higher API gravity of the coker distillate. This rather extensive cracking reduced the nitrogen content by almost one third but had little effect on the sulfur corltent. HYDROGENATION OF COKER DISTILLATE
Plant Equipment. At Rruceton, Pa., the high pressure, liquidphase, coal hydrogenation pilot plant was modified t o serve as a fixed catalyst bed shale-oil hydrogenation plant. A flow diagram of the modified plant is shown in Figure 2. The blended
2173
~
INDUSTRIAL AND ENGINEERING CHEMISTRY
2174
YIELDSA X V OPERATING CONDITIONS FROV TABLE I. PRODUCT DELAYED COKING OF N-T-U SHALEOIL Charge Rate,
Operating Condition; Shale oil Recycle stock
Bhl /Stream Day
86.3
212.8
Temp., 437 940 a2 5 780
Heater inlet Heater outlet Coke chamber Flash chamber, top Fractionator, top Fractionator. bottom Product Yields
Naphtha 40.6
34.8 1.2 45,fi
F.
360 GOO
Light Gas Oil :39,1
38.3 2.3 32.8
Coke
Gas
Loss
... 1.5.8 ...
... 7.5 ... ...
3.6
, . .
Recovery, %
97
Experimental Results. Two pilot plant runs were nixtic. witli a single pass of shale-oil cokw-distillate a t 1500 pouritlr per square inch gage total prcssure and at a n average rc,ac.tortemperature of' 835" F. Xlthough a large incrcase in tcnii)c.rat urc' was observed in the first catalyst section, only small vat,iatioiis O C C U I T ~thtwafter ~ as a result of thc addition of cooling h y t l r o g e ~ ~ . Thr, following tc:mlwr:rtuw prmofil(hd:it:t :iw tyl)ic-:ll o t t l i t , ( ' o i i trol obtainablv when opc9r:tting conditions ~ L I Y , go(~t1; t 1ic.y in^ based on ii ntatist.icit1 a~i:rlysisof a l l iritc%rnalthc,i,iiiocoul)l(,I , I , : L & ings duviiig :in ;irbitr:iry 60-liour pvi,iod ut'ar thv c S I i ( 1 of' t l i c . first run : Catalyst Zone 1 2 3 4
... ... ...
ASTM distillation,
99
aiid mixed coker distillate was drawn from a tared vessel by a variable-stroke reciprocating pump and with hydrogen supplied by a compressor was passed through a preheater consisting of a 190-foot coil of E/,e-inch outside diameter by 6/32-inch inside diameter stainless steel tubing in a n electrically heated molten lead bath. From the prcheat,er, the oil-hydrogen mixture was passed down through the catalyst-filled reactor and thence t o a condenser and a light-oil trap, where final cooling was accomplished by direct injection of water. This water-injection procedure permitted the uncondensed gases, after pressure reduction and metering, t o be sent directly t o the compressor for recycling. The condensed oil and water passed t o a float-operated discharge trap and thence t o a receiver at atmospheric pressure, where the water was decanted. The gases t h a t were dissolved in the product oil and released when the pressure was decreased were metered and collected for sampling. The arrangement of the catalyst in the reactor is shown in Figure 3. Four catalyst sections separated by sections of inert packing were provided. A hydrogen inlet line was connected through a manually operated control valve t o each inert-packing section. Thermocouples were inst.alled a t the t o p and bottom of each section of catalyst. The inert sections thus served t o mix the reactants with cool, fresh hydrogen and t o provide a means of controlling the highly exothermic hydrogenation reaction. The reactor was heated externally by six electrical furnaces, which, in turn, were controlled by individual thermocouples peened t o the outer reactor wall. The catalyst capacity of the reactor, as shown, was approximately '/a cubic foot. Operation of this hydrogenation pilot plant includes the following steps:
Metering the flow of fresh hydrogen gas and recycle gas t o the compressor suction; measuring the oil feed and scrubbing water flow rates by periodic gage-glass readings on the charge tanks; and weighing the total oil and water discharged from the plant both before and after decantation. A metered gas-purge stream, regulated to maintain the hydrocarbon content of the recycle gas at 10 t o 15 volume yo, was combined with the metered gases flashed from the products, and the mixture was sampled. The reactor temperatures were controlled by use of cooling hydrogen to prevent excessive inlet temperatures to any individual catalyst section and by adjustment of the preheater temperakure and the heat supplied by the external reactor heaters. By adding steam to the cooling water, the condenser was maintained a t 190" t o 200' F. to prevcnt deposition of ammonium salts.
Vol. 43, No. 9
Temperature, Inlet 6.51 f 7 840 f 7 840 f 3 836 =t3
O
14'.
Outlet 824 f 13 838 f 7 836 f 3 833 f 6
The plant was not provided with faci1itic:s for catalyst regc*lic~xtion, and therefore a nebv chargt. of cat,alyst was used for each ruii. The first run (SO-1) continued for more than 300 hours. It was discontinued at that time bec:tuk of thc slow but progicwivc increase in the nitrogen content of the -400" F. overhead fraction o f the product oil. It was anticipatcd t h a t continuing increase of nitrogen content would makc i t difficult t o rncet the limiting accelerated gum requircment (see Evaluation section) of' the applicable specification; this limiting condition was not achieved in the period of the run, however. The average: input of 19.6 pounds per hour corresponds t o a liquid hourly spacc: velocity of 1.13 volumes of feed stock per volume of catalyst. h sample of the product oil was taken every 6 hours and subjected t o a n ASTM distillation; t'he fraction boiling below 400" E'. was collected separately. This fraction was analyzed for nitrogen content, and the residue (400" F. to end point) was ex-itmined for aniline point. These data served as a n index of the activity of the catalyst. A daily composite sample also was taken for ultimate analysis. From these analyses and similar ones of the liquid feed and the gaseous product, a daily elemental balance was calculated t o determine hydrogen consumption, feed rate, and hydrocarbon gas production. Plotted against time in Figure 4 are control analyses of the distilled fractions of the samples taken every 6 hours, the daily determinations of the hydrogen absorbed, and hydrocarbon gases produced. After an initial period in which the data were erratic as a result of poor temperature control, the nitrogen content of the product oil fraction boiling below 400"F. was below 0.08 weight
TABLE 11. RECYCLE COKINGOF N-T-U SHALE-OILCHARGE STOCK AND PRODUCT PROPERTIES Gravity! O API a t 60' F. Pour point, F. Viscosity S.S.U. a t 130" F. S.S.U. a t looo F. Water Conradson carbon. wt. yo Total sulfur, wt. yo Nitrogen, wt.%
Charge Stock 19.8 80 139.8 Tiace 4.7 0.90 2.2
Combined Distillate 34.1 +5
...
35 Trace 1.05a 0.86 1.6
360 405 443 485 551 584 606 892
... ...
...
a
Recovery, vol. % On 10% bottoms.
e99 57
98
September 1951
I N D U S T R I A L A N D E N G I N E E R I N G CHEMISTRY
2175
te er
Shale ailcharge pump
Hot oil cycle pump
Figure 1.
Residuum Light g a s oil
Fractionotar reflux and storage p u m p
duction mid hyc1rogt.n coiisuiiiptioii :I:(' incrciiscd. 111 the riln :It 1.13 sl):icc' velocity, cxliniination o f nitrogvn illcreased almost in direct proportion to inere:iwti hyclrogyii consumption. A similar al".ough less pron01~11("l t r ~ i ~ n c I.l
F ~ MIJiayram Recycle Delayed Coking
70tor 260 hours. The xiiiliric. point of tlic fraction boiling above 400" F. \ v a ~almost constant; only a slight trend dowiiward is observed. The erratically loiv tempc>raturc attained during the first 40 to 50 hours of the run \vas reflected in the dry gas (butanes and 1ightt.r) production b!. low values during this pcsriod, but product.ion subsequcntl?; remained c~ssc~ntiallyconst,ant. This ttniount of hydrogen : ~ l ~ o r l decreased rd with timri of opcration. The iiitrog(,ii content of the l o w boiling fraction of t hc product continued t o increase to 0.10 weight yo at t,litI of t h e r u n a t which time the hydiogcsn absorption h a d dc eased to less than 1100 cubic feet per barrc'l of fwd. (.%I1 gas volumes given in this papc.r :ire%rc,portcd s t local metering conditioiis: 68" F. anti T 4 O m m . of mercury barometric prwsure.) .X saniplc~ of' the disrharged catalyst aftvr \v:+shing Lvitti t.oluene to rcniove oil was found to c.ont:Liii 6.0 \vc,ight. % carbon which is Lvithin t h i , pc~rniiwiblclimits for regeneration of the catalyst. The converter was rech:irged using the arrangement of catalyst and inert material dtwribed, and a second run (SO-2) \vas made. To f'urthrr drcrease the nitrogen contcnt of thr product, thc feed rate was rcduced t o 15.4 pounds p t r hour corresponding t o a liquid hourly space vdoaity of 0.9 volume per volume. As a result of the experience with the first run, analyses of the product oil were limited to those of the coniposittl samples takcn every 24 hours; the run was terminated after 320 hours a t which time enough product oil had been accumulated to supply evaluation test, requirements. The nit,rogen coiitent o f the total product oil, the light. hydrocarbon gas production, and the hydrogen consuniption are plotted against time in Figure 5 . T h r :tverage nitrogen content of thc total pr0duc.t oil \vas found to be less than 0.05 weight Yo for t>hc e n t i r e run; the dry gas pivdu:tion \vas ii('iw1.v const,ant ; the hytirogi~nc*ori?;uml)tiondata.
Feed tank
To
( 3 r d
pressor s u c I i o n
b-l
Recycle-gas meter
Preheater
3" I. D. x I l'-6'
converter
2116
O F SHALE OIL COKER DISTILLATE TABLE 111. HYDROGENATION
Operating Time, Days 1 2 3 4 5 6 7 8 9 10 11 12 13 Average 1
2 3
9 10 11 12 13 14 Average a
b
Vol. 43, No. 9
INDUSTRIAL AND ENGINEERING CHEMISTRY
(Reactor temp., 835' F.; pressure, 1500 Ib./sq. inch gage) Hz Liquid Product Oil Analysis. Wt. % .4bsorbed, Recoverya, Gas Vol. % Cu. E.t./Bbl. Cu. Ft./Bbl. Carbon Hydrogen Nitrogen Sulfur
Drg
100.6 98.5 98.3 97.4 96.9 97.8 98.0 97.0 98.5 97.9 98.2 97.7 97.4 98.0 97.7 98.1 96.9 96.8 96.9 95.9 96.1 95.8 96.1 97.7 95.6 94.2 95.5 96.1 96.4
Run SO-1, 1.13 Space Velocity 21 2 1174 86.29 300 1275 86.29 307 1227 86.21 86.30 341 1323 351 1291 86.30 317 1341 86.35 312 1214 86.37 348 1255 86.45 289 1144 86.35 323 1117 86.21 86.27 309 1186 86,30 327 1073 353 1040 86.23 315 1204
13.54 13.63 13.69 13.61 13.58 13.58 13.53 13.45 13.51 13.62 13.58 13.57 13.63
0.16 0.07 0.09 0.08 0.10 0.06 0.09 0.08 0 .13 0.13 0.14 0.12 0.13
0.01 0.01 0.01 0.01 0.02 0.01 0.01 0.02 0 0 .. 0 0 14 0.01 0.01 0.01
Run 50-2, 0.9 Space Velocity 1314 86.23 1314 86.19 1468 86.40 1384 86.23 1413 86.21 1358 86.24 86.19 41 1322 ..3 . 1342 86.22 413 1383 86.13 432 1256 386 86.28 1483 443 86.30 1285 412 86.23 1287 424 86.35 1268 401 86.32 1348 406
13.67 13.75 13.58 13.76 13.77 13.74 13.78 13.75 13.85 13.68 13.65 13.72 13.61 13.64
0.06 0.04 0.01
0.04 0.02 0.01 0.01 0.02
388 367 390 405 395 420
Includes pentanes (by calculation). Methane to C1, by calculation.
observable in the second run. Comparison of the nitrogen contents of the oil products for the two runs shows average values of approximately 0.11 weight yo for SO-1 and less than 0.02 weight yo for SO-2-a good improvement for a small decrease in space velocity. I n addition, the product oil shows a slightly higher hydrogen content at the lower space velocity. Expressed in volume-per cent, the liquid recovery includes the pentanes-andheavier fraction of the product gas. This fraction amounted to 5.3% of the total liquid recovery for run SO-1 and 7.5% for SO-2. The dry gap includes butaneg. Evaluation of products. T o evaluate the product oil in respect t o its conformance t o jet and Diesel fuel specifications, two composite samples were prepared. The sample of product of the first pilot plant run covered 200 hours of operation and did not include material from those periods where high nitrogen contents were observed; the sample representing the second pilot plant run included product from the entire run. These samples were distilled a t atmospheric pressure in two separate runs in a 500-gallon batch still using a Raschig ring-pa.cked column approximately equivalent t o 15 plates. The distillate, 60 volume yo of each sample, was evaluated as jet fuel and the bottoms as Diesel fuel. Table I V shows the properties of the jet fuel from both pilot plant runs. There are included, also, data obtained on a composite sample from run SO-2 to which inhibitor wasadded. The jet fuels were found t o conform t o Air Force-Navy Specification for J e t Fuel, AN-F-58, Grade JP-3, except for Reid vapor pressure which was below the specified minimum value of 5.0 pounds per square inch. (This designation has been superseded by MIL-F-5624, January 26, 1950, with little or no significant change in requirements. ) This departure from specification stems from lack of pilot plant facilities for condensation and recovery of light hydrocarbons. Some 5 t o 7 volume % of the calculated yield of liquid product is made up of pentanes and heavier contained in the exit gas streams. These hydrocarbons and ' butanes are present in sufficient quantity to impart the required vapor pressure to the finished product if they were condensed and recovered as liquids; in addition, this proccdure probably would
.. .. ..
0.02 0.01 0.01
...
0.02 0.02 0.03 0.02 0.03
...
0.02 0.02 0.02 0.02 0.03 0.02 0.02 0.01
adjust the observed freezing point of the SO-1 sample t o conform to the reauiwd value of -76" F. Some diffieulty was encountered with the SO-2 sample in meeting the minimum accelerated gum requirement of 20 nig. pcr 100 nil. The second SO-2 sample, t o which inhibitor was added, shows that improvemcnt in stability is possible; uti accelerated gum of 2.6 mg. was obtained, compared with 28 nig. for uni n h i b i t 4 jet fuel. The inhibited samplc showed no increase in accelerated gum value on 30 days' storage, whereas the uninhibited sample increased t o 9 i . 6 mg. To determine the Dossibi1it.c. of usine part of the jet fuel fraction as motor gasoline, octane ratings wcre obtained on jet fucl from run SO-2. R c d t s art' listed in Table V and indicate that additional processing for octane improvement is needed to obtain a suitable motor gasoline. Residue obtained by removing thc. jet fuel distillate fraction can br uwd directly as premium Diesel fui.1. The characteristics of the samplvs obtaincd from each of the two pilot p h n t runs are given in Table VI. Except for pour point and cloud point, these samplcs
-
Sk.Th 2.
Sk. Th. 4.
Sk. Th. 5.
KEY
Internal Th.-thermocouple
Figure 3. Detail of Catalyst Red
INDUSTRIAL AND ENGINEERING CHEMISTRY
September 1951
2177
direct injection of water t,o cool the product, Ammonia production can be estimated, however, by observing the difference in nitrogen cont,ent. of feed and product and assuming that all the eliminated nitrogen becomes converted t o ammonia. A n average of 5.7 pounds of ammonia per barrel of coker distillate hydrogenated was calcula.ted in these tests. Some 1.9 pounds of hydrogen sulfide formed per barrel of feed are similarly estimated from sulfur elimination. DISCUSSION
Over-all liquid recovery of jet and Diesel fuels from the original crude shale oil by recycle coking followed by hydrogenation of the coker distillate is 75 t o 78 volume %, using the operating conditions outlined here. Considerable variation in processing conditions are possible, and the yield and quality of products reported probably are not t'he best that can be obtained. The data presented are sufficient, however, when combined with Figure 4. Shale Oil Coker Distillate Hydrogenation knowledge previously obtained about the hyR u n S.0-1; 1500 Ib./square i n c h gage; 835' F.; 1.13 space velocity drogenat,ion characteristics of other nnatcrids over the cobalt molybdate catalyst employed ineet the requirements of S a v y Diesel Fuel Specification 7-0-2e ( 1 , 4 ) , t o evaluate theeffect of various factors on yields, product dated May 15, 1945. quality, and processing requirements. Analyses of hydrocarbon gases obtained in the hydrogenation The recycle-coking step exerts the greatest effect on over-all step are listed in Table VII. I t can be seen t h a t enough butanes volume recovery, t h a t for the hydrogenation step being nearly are available if required for blending with the jet fuel so a s to 100%. I n the operations described, the coker distillate end point meet specified vapor pressure. The C , t o CB hydrocarbons are was set arbitrarily at a maximum of 700" E'. in a n attempt t o present in sufficient quantity +oprovide a raw material for hydroensure a suitable pour point for the resulting Diesel fuel fraction gen production by reforming these hydrocarbons with steam. at the expense of liquid recovery. The recovery in the coking The ammonia produced by removal of ..itrogen during hydrostep might be increased t o 83 t o 87 volume yo by increasing the genation was not recovered in the gas stream because of the end point of the coker distillate t o 740" or 750" F. Hydrogenation of this distillate would yield a usable Diesel fuel which would have a higher pour point than that shown in Table VI. The TABLE IIr. JET-FCEL PROPERTIES Run~ KO. YO- 1
Freezing point, F. Corrosion Residue, nig. per 100 mi. Aromatics, yo S u l f u r , %o Gravity, API Reid vapor pressure, lb./sq. inch Water tolerance, ml. Accelerated gum. mg. per 100 ml. Bromine No. ASTM distillation, temp., I.B.P. 5% 10% 20% 30% 40% 50% 60% 70% 80% 90% 95% E.P. Recovery, vol. yo Residue, %
- 62
Pass
so-2
- 90
22.3 0.03 58.2 1.2-1 3 1 .o 14.4 1.3
Pass 0 8 20 5 0 1 51 2 1 6 1 0 28 0 0 9
163 217 237 264 288 309 331 352 376 40.5 446 484 514 98.0 1.0
162 214 225 257 279 298 318 342 363 392 430 471 500 98.0 1.0
1.1
_
so-25 - 83 1 0 20 6 51 2
2 2
1 0
2 6 0 8
F. 176 212 230 255 268 306 327 319 370 396 428 469 47 1 95.5 1.0
a Inhibited with 2.6-di-terl-butyl-4-methylphenol 1 pound inhibitor per 5000 gallons.
TABLE V.
JETFUELOCTANERATINGS
Motor Method, clear 1 cc. T E L 3 cc. T E L Research, clear 1 cc. TEL -t 3 cc. TEL
+
+ +
39.5 52.1 65.1 44.5 55.1 69.7
_
TABLE VI.
DIESELFUELPROPERTIES Run N o . __
so-1 Flash point, O F. Pour point, F. Cloud point O F. Viscosity, s,'s.u.a t 1000 F. Water Sediment Carbon residue in 10% bottoms, % .ish, Corrosion S u l f u r , wt. yo Ignition, cetane No. Distillation, F. 10% 30% 50 %I 70% 90 %
TABLE
VII.
11
36 39.9 Trace Trace 0.05 0.003 Pass 0.1 54.6
0 0 Pass 0 2 60.4
450 475 504 538 581
430 466 489 523 558
AVERAGECOLIPOSITIOX
Methane Ethane Ethylene Propane Propylene Butanes Butylenes Carbon monoxide Carbon dioxide
Av. production,
cu. ft./bbl. feed
so-2
187 15 22 34 8
188
0 0 0 0 0 03
OF D R Y G A S
Constituent, Vol. 7% SO-1 SO-2 39.94 38.54 20.98 21.01 18.00
2.23 12.31 5.60 0.37 0.57 100.00
i?'.'Ss
3.77 11.96 6.50
-
.. 2
315
406
100.00
INDUSTRIAL AND ENGINEERING CHEMISTRY
2178 c
Vol. 43, No. 9
ACKNOWLEDGMENT
02
+-
I n the experiments described, the extent of :p 01 the operation and the necessity for utilizing equipment a t both Rifle, Colo., and Bruceton, 2s Pa., have obligated the authors t o many indis o 500 z ~ z viduals for their advice and assistance. It is impossible t o detail the part each individual has O2: $5-1 played in making this work possible, and the 400 En," authors' thanks must be directed t o the groups SLS in various sections and t o their supervisors. 8p'm nocc I503 300%:: The operators and maintenance men at the c L - - rO l Shale Oil Demonstration Plant Refinery who T, w conducted the recycle coking operation under 9 1400 the supervision of J. E. Phillips and the pilot z'ad ? a plant operating and maintenance groups a t 1300 Bruceton, particularly John Nanz, Arnold Pipilen, and William R. Wilson, who supervised ,": the hydrogenation experiments, can be credited 1200 0 50 100 150 200 2 50 300 350 with a large share of the success of the runs HOURS OF OPERATION reported. The analytical work was supervised by -4.S. Houghton at Rifle, Colo., and by W. R. Figure 5. Shale Oil Coker Distillate Hydrogenation Rosinski a t Bruceton. The authors are inRun S.0-2; 1500 lb./square inch gage; 835' F.; 0.9 space velocity debted t o R. A. Friedel and his staff for the inass spectrometric analyses and t o S. Wekselhydrogcbnation step results in sufficient reduction of molecular man and M. Corrado for the calcalution of data. The authors weight to produce a Diesel fuel having a 90% point of 558' F. also wish t o thank J. Vidosh and his staff for preparation of the from :t coker distillate with a n end point of 700" F. The maxidrawings. The Union Oil Co. donated the catalyst, and the helpmum permissible 90% point for Diesel fuel is a temperature of ful suggestions of Clyde Berg and H. C. Huffman of t h a t company 675" F. so the use of a 750' F. end point coker distillate should are appreciated. Thanks are due H. H. Storch, W. C. Schroeder, produce a Diesel fuel fraction of satisiactory boiling range. The Boyd Guthrie, and H. M. Thorne for their assistance in cohigh(.r molecular weight feed stock for the hydrogenation step ordinating the efforts of the various groups in the Office of may cause a decrease in efficiency of the catalyst for nitrogen Synthetic Liquid Fuels. elimination and may increase the fiequrncy of regenerationof the catttlyst. LITERATURE CITED Comparison of the two runs dehcribccl shows that decrease 111 Berg, Clyde, Bradley, W. E., Stirton, R. I., Fairfield, R. G., liquid space velocity results in increased production of gas. .4t Leffert, C . B., and Ballard, J. H., Chem. Eng. Progress, 43, the same time a decrease in nitrogen content of the product was 1-12 (1947). observed. A marked and costly increase in hydrogen consump. Byrnes, A. C., Bradley, W. E., and Lee, M. W., IND. ENG. CHEM., 3 5 , 1160-7 (1943). tion accompanies the increased gasification. The variation in Hendricks, G. W., Huffman, H. C . , Parker, R. L., and Stirton, liquid recovery is quite small in this step and is subordinate t o the R. I., presented before the Division of Petroleum Chemistry importance of nitrogen elimination and hydrogen consumption. a t the 109th Meeting, AMERICAK CHEMICAL SOCIETY, Atlantic Some indication of the composition of the products may be obCity, N. J. tained from the analytical data. The low octane numbers of the Hughes, E. C . , Stine, H. M., and Faris, R. B., IND.ENG.CHEM.. jet fuel fraction, in spite of a 20y0 aromatics content, indicate 42,1879-82 (1950). Hull, W. Q., Guthrie, B., and Sipprelle, E. M., Ibid., 4 3 , 2-15 the presence of a large amount of saturated normal paraffin (195 1 ). hydrocarbons. This conclusion is substantiated by the high Lankford, J. D., and Ellis, C. F., Ibid., 4 3 , 27-32 (1951). pour point of the Diesel fuel in spite of the rather low boiling Mills, G. A., Boedeker, E. R., and Oblad, A. G., J . Am. Chem. rhnge (for Diesel fuel). It thus appears t h a t hydrogenation over SOC., 72, 1554-60 (1950). cobalt molybdate catalyst t o remove sulfur and nitrogen proMorris, H. B., and Gilbertson, D. L., Petroleum Engr., 21, No. 9, C26-32 (August 1949). duces a material requiring additional processing if it is t o meet Reed, Homer, and Berg, Clyde, Petroleum Processing, 3, 1187gasoline specifications.
g:
-IL
f
92 (1948).
CONCLUSION
A technically feasible process for the conversion of crude shale oil into specification grade jet and Diesel fuel is the combination of recycle coking and hydrogenation of the coker distillate a t 1500 pounds per square inch pressure. The hydrogenate, on simple distillation, yields 60 volume yo jet fuel (overhead) and approximately 40 volume yoDiesel fuel.
Secretary of the Interior, Rept. Synthetic Fuels, Part 2 (1950) [U. S. Bur. Mines, Rept. Invest 4 6 5 2 ( 1 9 5 0 ) l . Thorne, H. M., Murphy, W. I. R., Ball, J. S., Stanfield, K. E., and Horne, J. W., IND.ENG.CHEM.,43, 2& 7 (1951). RECEIVED February 23, 1951. Contribution from the Synthetic'Fuels Research BranchFBureau of Mines, Brureton, Pa. Presented before the Division of Industrial and Engineering Chemistry at the 119th Meeting of t h e AMERICAN CHEMICAL SOCIETY, Cleveland, Ohio.
