I
HARRY CLOUGH Imperial Chemical Industries Ltd., Billingham-on-Tees, England
Gasoline Production from Coal Tar Oils By minor operating changes and improving catalysts, this British hydrogenolysis plant produces from creosote, motor and aviation gasolines that compete with rising octane numbers of premium grades
HIGH
PRESSURE hydrogenolysis, or destructive hydrogenation, will eventually find its most extensive use in producing from coal, liquid fuels ranging from gasoline to heavy fuel oils. When shifts in relative cost of coal and crude petroleum forced coal hydrogenation back to small-scale production, some Western German hydrogenation plants (7) used this process to make motor gasoline and Diesel fuel from petroleum resi-
dues (9), and after the wartime plant for producing aviation gasoline from petroleum oil gas (6, 8) in Great Britain was converted, the Billingham plant of Imperial Chemical Industries Ltd. (7, 2) continued to produce motor and aviation gasolines from creosote. Hydrogenolysis of solid and liquid feed stocks has been generally reviewed (3, 4 ) . This report describes in more detail, developments in the process used
T h i s article i s a reminder that sources of liquid fuels other than natural petroleum must not be overlooked. Petroleum reserves appear to be maintained despite increasing annual consumption; but, nevertheless, cost of producing crude i s continually increasing. In due course, when either costs shift or petroleum supplies become inadequate, synthetic fuels based on coal and, probably to a less extent, on shale will become sources of liquid fuels. When contemplating the long term future, i t i s indeed reassuring that world energy reserves as coal are about 500 times greater than those as petroleum. Certain countries have already had considerable full-scale experience in synthetic fuel production both by coal hydrogenation and the Fischer-Tropsch synthesis. Before the war, German oil flelds had not been discovered and during World War II, Germany relied almost entirely on synthetic fuels; her hydrogenation plants no longer treat coal or tar but have been integrated wifh refineries processing indigenous and imported crudes and permit these reflneries economically to meet Germany's particular pattern of demand for light, middle, and heavy fuels. In Great Britain, pureey economic factors have put coal hydrogenation out of the picture and most of the equipment and experience in high pressure techniques available to IC1 at its Billingham factory have been put to chemical uses such as hydration of oleflns to alcohols, carbonylation of olefins, production of synthetic phenols, and hydrogenolysis of animal and vegetable oils to alcohols. But motor fuel production from tar oils still continues at Billingham in part of the plant; this article shows how the process has been developed to
at Billingham for producing motor gasoline from creosote. The Process The creosote fraction of coal tar which is hydrogenated boils from approximately 200' to 340" C. It is highly aromatic, particularly when obtained from high temperature carbonization processes, and contains aromatic hydrocarbons such as substituted benzenes,
keep abreast of the demand for continually higher quality gasolines over the 21 years of its operation. Recent improvements in catalysts have in fact now made feasible production of 100 octane gasoline by creosote hydrogenation, even at the low level of 1.25 ml. TEL per U. S. gallon which i s used in Great Britain. Probably the most striking feature of this process i s its flexibility not only in feedstock but also in products which have ranged from fuel oils through Diesel oils to aviation gasolines. Performance required from most petroI%eum products i s more exacting today than when coal hydrogenation processes were last operated extensively, but advances made during the last 10 years in the theory and practice of catalysis will be just as rewarding in other destructive hydrogenation processes as they have been for creosote hydrogenation at Billingham. One important point, however, must not be overlooked. Production of liquid fuels from bituminous coal will require about 10,000 cubic feet of hydrogen per barrel of fuel produced, and this will probably have to come from coal by some gasification process. Even hydrogen from catalytic reforming processes, available so cheaply today, will have greater value since it will, of course, be produced by gasiflcation of coal. The cost of hydrogen will inevitably affect not only the cost of synthetic fuels but also that of chemicals based, for instance, on ammonia and methanol which are today prepared so cheaply from refinery hydrogen. The most urgent problem over the next few years, therefore, seems to be not so much detailed improvement of hydrogenation processes but development of some completely new process for making cheaper hydrogen from coal.
VOL. 49, NO. 4
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CONVERTER
MAKE -UP He CIRCULATOR 1
REFINED OIL STORAGE LIQUOR
t----]
{his creosote hydrogenation process yields gasoline in two steps naphthalene, naphthalene derivatives, and more highly condensed aromatic hydrocarbons, plus appreciable quantities of phenols, bases, and sulfur compounds. From typical elementary analyses of creosotes-0.5 to 0.7y0 nitrogen, 0.7 to 0.9% sulfur, and up to 2% oxygen-the hydrocarbon content of creosote oil is estimated at only about 70%. Its average molecular weight is in the region of 160 to 170 compared with approximately 100 for motor gasoline; the hydrogen content of high temperature creosotes is about 8 to S.5yO compared with 14 to 15Yo for a low olefin content gasoline. The first destructive hydrogenation processes put in full-scale operation yielded gasoline directly in a single step; nitrogen, sulfur, and oxygen were eliminated by the well-known hydrofining reactions which proceeded simultaneously with reduction of molecular weight by hydrocracking reactions. This process was operated at a pressure of 250 atm. (3700 pounds per square inch gage). The oil, together with a stream of circulating hydrogen, was led through heat exchangers to a gas-fired preheater where complete vaporization was effected a t 400" C.; the reactants then passed through converters packed with pelleted catalyst. The highly exothermic reaction-550 to 600 kcal. per kg. of gasoline produced-had a high positive temperature coefficient and the catalyst was therefore arranged in beds separated by gas-mixing chambers into which cold-shot hydrogen could be introduced to keep temperatures within a range of some 10" to 15'C.
674
Products leaving the converter passed through the interchangers and cooler to a high pressure gas-liquid separator. The hydrogen was recycled to the inlet of the plant and the liquid product was let down to atmospheric pressure in stages. Dissolved gases released at intermediate pressures, consisting mainly of methane and hydrogen, were used as fuel; those coming out of solution at low pressure were chiefly propane and butane. The liquid product was distilled into gasoline overheads and middle oil bottoms which were recycled for further treatment; the gasoline content of the product amounted to between 50 and 70% according to the type of gasoline being produced. The converters are just over 4 feet in internal diameter and 50 feet long; the wall thickness is inches and each weighs about 160 tons.
catalyst, however, rapidly lost activity in the presence of organic nitrogen compounds and, while certain petroleum fractions could still be processed directly over it, coal and tar oils required a preliminary hydrofining treatment to reduce nitrogen to not more than 5 p.p.m. before hydrocracking; at the same time, sulfur was reduced to less than O . O l ~ o and oxygen to 0.1 to 0.276. Hydrofining was carried out in a plant identical with that described for vaporphase hydrocracking but was a oncethrough as distinct from a recycle process. The catalyst employed for the refining stage was the same tungsten sulfide previously used for the single-stage process, but space velocity and temperature were adjusted to give minimum hydrogenation consistent with the required nitrogen removal. Despite an inevitable reduction in boiling range of the oil caused, for example, by conversion of phencls to the corresponding hydrocarbons and by saturation of aromatics to lower-boiling naphthenes, it was possible to limit production of gasoline-boiling range material to less than 15% of the oil processed; conversion of oil to gasmethane, ethane, propane, and butanes -was only 1 to 29" by weight. This resulted from careful attention to design of the converter internals which permitted accurate temperature regulation throughout the converter (Figure 1). Life of 6434 catalyst was improved by acid xvashing the hydrofined oil to 1 p.p.m. of nitrogen or less. Control of the hydrofining operation was then less critical-products hydrofined only to 10 p,p.m. basic nitrogen could be processed satisfactorily after washing with dilute acid. Less severely hydrofined oils xvere unsatisfactory, even after acid
Early Improvements in Vapor-Phase Hydrogenation Catalysts
For a number of years, pelleted tungsten sulfide (5058) had been the catalyst used in hydrogenating vaporizable middle oils. It was a robust catalyst, active in the presence of high concentration of nitrogen, sulfur, and oxygen compounds; but even with highly aromatic feed stocks, it gave gasolines which were largely naphthenic and paraffinic in composition and therefore of relatively low octane number. I n 1935, a catalyst (6434), developed by I. G. Farben, partly avoided this oversaturation of the gasoline; it consisted of tungsten sulfide supported on a high surface-area natural earth. This
INDUSTRIAL AND ENGINEERING CHEMISTRY
The 160-ton reactor being removed for catalyst change
H Y D R O G E N I N THE PETROLEUM I N D U S T R J Table 11.
Composition of an Aviation Naphtha
(135' C. end point, from creosote using catalyst 231) Vol.
T CONVERTERS
PREHEATER
Figure 1.
Temperature gradient through reactor system
washing. This was attributed to nonbasic nitrogen compounds that persisted in the oil. At first, 6434 catalyst was regarded merely as an extended or diluted form of 5058 but it was soon recognized that the activated earth was in fact responsible for much of the hydrocracking activitytungsten sulfide protected the earth against fouling by carbonaceous deposits. By 1939, work in IC1 laboratories on catalysts in which tungsten sulfide was replaced by less active components not required themselves to contribute to hydrocracking, had led to the development of catalyst 231 consisting of iron on a hydrofluoric acidactivated montmorillonite. Although
still capable of preventing carbonaceous deposits, the iron brought about less hydrogenation of aromatics to naphthenes. This was reflected in the motor method octane number of the gasolines, but of much greater importance was its marked effect on the supercharge rating of aviation naphthas produced by hydrogenation-e.g., a performance number of 137 with 231 catalyst, compared with 122 using 6434 catalyst. Development of 231 catalyst was a major factor in enabling 100/150 Pursuit Grade gasoline to be produced in Great Britain in 1944. Various stages of development discudsed so far are summarized in Table I. The improvement in octane number obtained by separating the refining and
Table 1.
Typical Operating Conditions and Yields for Vapor Phase Hydrogenation of Creosote Oil [ l o lb./sq. inch R.V.P.motor gasoline (45% D 4- L at 100' C.) from creosote]
Process Catalyst Operating conditions Fresh oil rate, kg./l cat. cat./hr. Total oil fed Conversion, yo Vapg. gas rate, I./kg./ total feed Cooling gas rate yo Hs in inlet gas Total pressure, Ib./sq. inch Av. temp., C. Yield based on fresh feed Gasoline, vol. yo Excess liq. butane VOl. 70 Gas, wt. yo Hydrogen used, wt. % Hydrogen used, cu. ft./ bbl. Gasoline compn., POI. Yo Straight chain paraffins Branched paraffins Naphthenes Aromatics M.M. octane No. Clear 1.25 ml. TEL/US gal. 3.0 ml. TELfUS gal. R.M. octane No. Clear 1.25 ml. TEL/US gal. 3.0 ml. TEL US gal.
Single Stage 5058
Hydrofining 5058
Hydrocracking 6434 23 1
0.59 1.1 54
1.2 1.2 100
0.85 1.3 65
0.85 1.3 65
2000 1700 85
2800 1900 87
1600 900 83
2000 1000 83
3700 420
3700 385
3700 400
3700 370
120
(113)
109
8 10.0 8.8
2.5 5.6
5870
3740
16 31 50 3 68 78
Two-St age 5058, 5058, 6434 231 0.50
0.50
107
123
121
6.8 4.8 3.4
6.8 4.4 3.0
7.7 7.3 9.0
7.7 6.9 8.6
1970
1740
6000
5750
6 33 54 7
5 32 48 15
74 81
76 83 87.5
79 87
81 89 93
'
(28)
Paraffins Isopentane %-Pentane Branchedhexanes %-Hexane Branched heptanes n-Heptane Branched octanes n-Octane Naphthenes C yclopentane Methylcyclopentane Cyclohexane Dimethylcyclopentane Methylcyclohexane Dimethylcyclohexanes etc. Aromatics Benzene Toluene Xylenes and ethylbenzene
% 17.0 2.5 10.5 2.0 5.5 0.2 3.5 0.2 1.0 13.0 2.5 6.0 11.0 8.0
4.0 5.0 3.0
cracking stages required over-all only about 15 to 20Yo more catalyst volume; the volumetric yield of gasoline from the two-stage process is slightly better than from the single-stage process, despite the increase in aromatic content. This improvement in yield probably results from hydrocracking at a somewhat lower temperature in the two-stage process and the virtual elimination of methane and ethane production. Although less hydrogen was chemically absorbed in the two-stage proceses, the full saving could not be realized because of additional loss of dissolved hydrogen on depressuring between the hydrofining and hydrocracking stages. Also 70 to 75y0 of the butanes consist of isobutane. This unusually high degree of branching is characteristic of gasolines produced by destructive hydrogenation processes and occurs throughout the boiling range (Table 11). Two features stand out. First, the iso-normal paraffin ratio is always so much higher than that corresponding to equilibrium under reaction conditions that the is0 structures must result from a particular pattern of molecular degradation; second, even though hydrocracking catalysts show little activity for paraffin isomerization, they are extremely effective in six-ring to five-ring naphthene isomerization. The first circumstance contributes significantly to the engine rating of hydrogenation gasoline; fortunately, the second has little effect. Composition of the starting material is too complex and possible combinations of reactions are too multitudinous to permit drawing up detailed schemes; however, the mechanism by which various types of hydrocarbon arise can VOL. 49, NO. 4
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be simply illustrated for naphthalene, one of the major constituents of creosote. Hydrocracking of aromatics to fragments of lower molecular weight cannot occur until a point of attack has been provided by saturation of at least one of the rings. An exception to this, of course, is the removal of side chains, but in creosote these are usually so short that their removal would take place only a t temperatures much higher than those used in hydrocracking. The condensed ring structure of the components of coal tar, together with the isomerization of six-ring to five-ring naphthenes is clearly the reason for the branched nature of the paraffins formed. Postwar Problems
From 1940 to 1953, only one grade of gasoline was marketed in Great Britain and this corresponded roughly to prewar regular grade-viz., 72 octane number motor method at 1 . 5 ml. of tetraethyllead/ US gal. max. When oil companies were again permitted to market their own premium grades, an octane race rapidly developed, which received impetus from the fact that since the war many new refineries had been constructed incorporating the latest petroleum processes. Gasolines also began to be specified in terms of Research octane number rather than Motor Method octane number. This was naturally advantageous to gasolines containing olefinic components such as
Table 111.
cracked gasoline and polymer gasoline. Olefins are, of course, inherently excluded from hydrogenation gasolines which fundamentally rely on isoparaffins for Motor Method octane number and aromatics for Research octane number. Large proportions of naphthenes in hydrogenation gasolines render their octane ratings sensitive to end point because higher-boiling naphthenes have relatively poor engine performance. By cutting back more severely to give higher volatility gasolines and recycling the heavier fraction to hydrocracking, octane ratings can be significantly improved, but a t the expense of additional hydrogen consumption and reduced gasoline yield (Table 111). Butane dehydrogenation and polymerization minimize the extent to which volatility must be raised and to some extent restore over-all gasoline yield. Pl'evertheless, volatility increase remains an expensive step because of increased hydrogen usage. Both technically and economically, the most attractive way of meeting increased octane number requirement, particularly the Research octane number, is to increase still further the aromaticity of the gasoline. There were three obvious lines of attack. A process had already been developed by IC1 just before the war for dehydrogenating six-ring naphthenes in gasolines to aromatics using a platinum on charcoal catalyst (5). Yield of aromatics on naphthenes converted was
Gasoline from Creosote
(10lb./sq. inch gage RVP using catalyst 231) 165 Gasoline end point, C. 200 50 35 Volatility, yo D L at 100' C. 119 Gasoline yield, vol. yo 125 10 2 Excess liquid butanes, vol. % 8.8 8. 1 Hz consumed, wt. % 77 73 M.M. clear 84 80 1.25 ml. TEL/US gal. 82 79 R.M. clear 90 86 1.25 ml. TEL/US gal.
+
676
INDUSTRIAL AND ENGINEERING CHEMISTRY
130 73 110 23 9.8 79 87 84 92
stoichiometric-no hydrocarbon gas was produced; five-ring naphthenes, however, were not converted. Although these, and to some extent paraffins, could be aromatized in the newly developed catalytic reforming processes using platinum on alumina, these processes gave considerably lower gasoline yields; also, it seemed likely that when applied to hydrogenation gasolines, part of the octane improvement arising from aromatization would be offset by isomerization of paraffins to a less favorable isonormal ratio. Addition of gasoline dehydrogenation or reforming, of course, implied new plant construction. Alternatively, saturation of aromatics to naphthenes could be limited by hydrogenating a t much higher temperatures where, even under the high partial pressure of hydrogen employed, the equilibrium constant is not so overwhelmingly in favor of naphthenes. A process employing these principles was used in the early 1930's before catalyst 5058 was developed. The catalyst used was a mixture of molybdenum trioxide with magnesium and zinc oxides which, in single-stage operation on unrefined middle oils at temperatures over 500" C., gave a gasoline of the following composition :
Gasoline from Creosote, Using Catalyst 64 Vol.
% Aromatics Naphthenes Branched paraffins Straight-chain paraffins
olefins
Octane No. M.M., clear
43
33
7 14 3 80
The aromatic-naphthene ratio is greatly improved, but this again is offset to some extent by the accompanying deterioration in iso-normal paraffin ratio. Operation at temperatures above 475' C. was not attractive because considerable modification of the equipment would be needed. Also, loss in yield of gasoline to methane and ethane, which consumes expensive hydrogen, is considerable. The third line of attack was to reopen work on developing new catalysts for both hydrofining and hydrocracking stages which would operate with less hydrogenation of aromatics to naphthenes. This was finally chosen. The Hydrofining Stage
Tungsten sulfide (5058) has been claimed as an effective catalyst for a wide variety of hydrocarbon reactions-e.g., hydrogenation of olefins and aromatics and isomerization of naphthenes and paraffins, as well as hydrogenolysis and
H Y D R O G E N IN T H E P E T R O L E U M I N D U S T R Y Tungsten Sulfide Compared with Cobalt Molybdate-Type Catalyst (3000 lb./lq. inch gage; 97% hydrogen) Catalyst
Wt. % Ref. No. Tungsten sulfide 7424 7457 7493
co
Mo
1.6 0.8 8.4
6.3 6.5 6.5
...
hydrocracking (70). Although the relative activity in these various reactions naturally varies according to temperature and pressure employed, it is clear that tungsten sulfide cannot be regarded as a selective catalyst. Probably because of the tungsten shortage in Germany during the war, I. G. Farben developed a new catalyst (8376) to replace 5058; this consisted of 25% tungsten sulfide and 3y0 nickel sulfide on activated alumina. I t was claimed that this catalyst removed impurities from middle oils with less saturation of the aromatic components, but there was a suggestion that its life was inferior to tungsten sulfide. I n the last few years a large number of processes have been announced for desulfurizing petroleum oils, particularly gas oils, in which cobalt molybdate supported on alumina is used as hydrofining catalyst under moderate hydrogen pressure (200 to 800 pounds per square inch). The catalyst has long life and, under suitable conditions, a high degree of sulfur removal can be achieved with little more than theoretical hydrogen consumption. The good performance of this desulfurization catalyst, coupled with the I. G. Farben use of an alumina-based refining catalyst and some promising results obtained a t Billingham in 1944 on cobalt thiomolybdate, led to preliminary experiments in 1954 on use of cobalt molybdate-alumina under conditions similar to those used with tungsten sulfide in the hydrofining of creosote middle oil. I t was confirmed that under 3000 pounds per square inch hydrogen pressure, nitrogen could be reduced to about 10 p.p.m. with considerably less saturation of aromatics than when using tungsten sulfide. Somewhat higher temperatures (400" to 425O C.) were required than with catalyst 5058, but no loss of activjty was encountered even when the catalyst was operated for long periods a t pressures as low as 1500 pounds per square inch. A large number of cobalt molydate on alumina catalysts was then prepared and the catalysts were examined for hydrofining performance in small plants processing about 75 ml. per hour of creosote. Catalyst variables examined in-
...
Av. Temp. for 10 P.P.M. N
Aromatics, Vol. %
365' C. 390 403 425
28 40 50 53
cluded the atomic ratio of cobalt to molybdenum a t constant total metal content, varying total metal content at constant atom ratios, various grades of alumina, and pelleting pressures. For a given degree of nitrogen removal-e.g., 10 p.p.m. in the product-the temperature required a t a liquid throughput of 1 kg. per liter per hour increased with decreasing total metal concentration, the atomic ratio of the two metals had only a second-order effect. I t appeared, however, that for a given degree of nitrogen removal, lower metal concentrations and catalysts with a low cobalt-molybdenum ratio gave better aromatic retention. Of course, increased aromaticity of products from low metal and low ratio catalysts may not be caused by hydrogenolysis specificity; it may result from the effect on kinetics of less negative freeenergy change for aromatic hydrogenation a t the higher temperatures necessary with such catalysts. Effect of preparation method was examined in a limited number of cases but no general principles could be established. With some compositions, impregnation by molybdenum followed by cobalt, gave more active catalysts, whereas in others the reverse order, cobalt followed by molybdenum, gave more active catalysts. Some work was also carried out with manganese and iron molybdates on alumina; again performance of the catalyst depended markedly on precise method of preparation. Required nitrogen elimination was achieved only a t temperatures 30' C. or more above those required for cobalt molybdate, but products were correspondingly more aromatic. Some of the more promising catalysts were then examined in greater detail on a semitechnical scale-Le., in plants processing 1 to 2 liters per hour of creosote oil. The above tabulation shows the superiority of these cobalt molybdatetype catalysts over tungsten sulfide. Middle oils refined over cobalt molybdate catalysts werw then hydrocracked on the small scale using 231 catalyst. Value of the aromatics was confirmedresearch ratingsof the gasolines were 2 to 3 units higher than for gasolines from 5058refined oils. The more aromatic middle oils were only a little more resistant to
hydrocracking, and conversion to gasoline could be maintained by raising temperatures to about 5" C. above those required for oils refined over catalyst 5058. 1,n an experimental run on the full scale plant, a cobalt molybdate catalyst has operated satisfactorily with no significant decrease in activity for several months. Basic nitrogen content of the refined product can be allowed to rise to 25 to 30 p.p.m. prior to the dilute acid wash without sufficient nonbasic nitrogen compounds remaining to cause difficulty in hydrocracking over catalyst 231. This less severe refining leads, of course, to a still higher aromatic content product. Desulfurization is less effective than when using catalyst 5058 but the remaining 0,05y0 of sulfur is reduced to a few parts per million during hydrocracking; oxygen content of the refined oil, 0.01 to 0.0270, is considerably lower than when tungsten sulfide is used. Mainly as a result of the improved retention of aromatics, hydrogen consumption in the refining stage is now considerably less than when using tungsten sulfide. Greater selectivity of cobalt molybdate-alumina for hydrofining-i.e., its relatively low activity for saturating aromatics and breaking C-C bonds-has made the full-scale operation and control of the hydrofining process much easier than when using tungsten sulfide. Thus, any local high temperatures do not cause hydrocracking reactions which are themselves exothermic and have high temperature coefficients. Consequently, the converter system is more stable and can operate a t throughputs up to 25 to 30% higher than those possible when using tungsten sulfide in the same converter. This increase in throughput requires an average catalyst temperature only 10" C. higher. Using the same hydrocracking catalyst 231, the Research octane number (1.25 ml. of TEL/ US gal.) of the gasoline produced on the full scale plant rose by 1.5 units at the same volatility. One further advantage over 5058 is the ability of these catalysts to operate continuously at total pressures considerably below 3700 pounds per square inch. This, of course, is more important to new plants than existing installations, but even in existing installations some advantage results from reducing the quantity of hydrogen removed from the high pressuresystem by solution in the product. No doubt additional gains in aromaticity could be obtained from further study of the hydrofining stage and its catalyst. With the current hydrocracking catalyst 231, however, approximately one half of the aromatics in a refined middle oil were destroyed during conversion to gasoline; it was considered more profitable a t this point to switch VOL. 49, NO. 4
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Table IV. Comparison of Catalysts 231 and 7679 (Pressure 3000 lb./sq. inch, 97% hydrogen; throughput, 1 kg./l./hr. of cobalt molybdaterefined creosote: aromatics, 46 vol. %) 6 Lb./Sq. Inch RVP Gasoline Properties Temp., Convn., D L at F.B.P., 1.25M1. TEL/US Gal. Catalyst O c. % ’ 100’ C., % O C. Aromatics M.M. R.M.
+
231 7679
375 3 80
68 75
50 51
attention from hydrofining to hydrocracking.
The Hydrocracking Stage The simplified reaction scheme or mechanism given earlier brought out that hydrogenolysis of aromatics can proceed only after saturation of at least one ring has provided a point of attack. I t is desirable, however, that saturation activity of the catalyst should so relate to its hydrogenolysis activity that naphthenes formed do not greatly exceed those which can be split. In the absence of other considerations, this balance should readily be achievable by varying the concentration of hydrogenating metallic element on the cracking catalyst. In contradistinction to catalytic cracking processes, however, high pressure hydrocracking does not allow frequent catalyst regeneration; therefore, lay-down of carbonaceous deposits must be prevented by saturation of their olefinic precursors before they polymerize on the catalyst surface. I n the past, concentration of metal necessary to ensure satisfactory catalyst life had to be such that aromatic saturation exceeded splitting capacity of the catalyst. It seemed that improvements could be brought about in two waysfirst, by replacing the natural earths with more active and selective synthetic alumina-silicas; and secondly, by using as the protecting metal, an element having high olefin saturation activity but only moderate aromatic saturation activity. A large number of catalysts has been examined-metals of Groups V1 to V I I I on both proprietary cracking catalysts and on specially prepared aluminasilicas. Among the first formulations tried was a nickel on silica-alumina (7679) which gave considerably better results than those obtained with the standard 231 catalyst; it was decided to develop this type of catalyst for fullscale operation while continuing the search for even better catalysts. The performance of 7’679 compared with that of 231 catalyst is illustrated in Table IV. Semitechnical-scale experience with catalyst 7679 soon showed other advantages. Its hydrocracking activity was not affected by nickel content up to 5% and aromaticity was only slightly affected. Pressure could be reduced to 1800 pounds per square inch with only
678
163 146
18 23
85.5 87
91.5 94.5
small loss in conversion and no permanent or irreversible effect on the catalyst. Without damage to the catalyst, temperature could be raised to 450’ C. a t which conversion to gasoline was only 5% greater; but as a result of its higher aromatic content, this gasoline had a Research octane number of 96.6. Either of these operating conditions would have irreparably damaged catalyst 231. Finally, 7679 was less sensitive to nitrogen compounds; oils hydrofined to 80 to 100 p.p.m. and then washed to 3 to 4 p.p.m. basic nitrogen caused only temporary depression of hydrocracking activity. Product quality remained unaffected and conversion was fully recovered on reverting to normally refined feed. This catalyst has been installed in the full-scale plant and expectations based on semitechnical work have been fully realized. Leaded research octane numbers have risen by 2 to 3 units; gas make is no greater than with 231 catalyst; although satisfactory plant balances have not been established, increased aromaticity suggests a t least 10% saving in hydrogen consumption. Contrasted to the tolerant nature of 7’67’9 catalyst, striking cases of complete or almost complete inactivity have been encountered, even with nickel, on some specially prepared and some commercial alumina-silicas of proved catalytic cracking activity. Again, on some aluminasilicas certain elements such as nickel depress hydrocracking activity to a greater extent than considerably larger amounts of other elements such as iron. Therefore it is impractical to consider the functions of cracking and “protection” separately; without knowledge of the intimate relationship which obviously exists between alumina-silica and added metal, ad hoc experimentation remains the only effective approach to the problem. Platinum catalysts have ranged from almost complete inactivity to undesirably high activities-one catalyst gave complete saturation of aromatics until throughput was increased to 4 kg. per liter per hour and pressure decreased to 1200 pounds per square inch. Further work might lead to methods of controlling catalyst activity but platinum does not appear particularly attractive. Generally, for a given per cent of aromatics, a platinum catalyst product has a lower Research octane number than a nickel product, presumably because of undesirable paraffin isomerization.
INDUSTRIAL AND ENGINEERING CHEMISTRY
Cobalt seems to give a slightly higher octane number than nickel but a somewhat larger concentration is required to give satisfactory life; as a result, conversion suffers slightly. In the same way that 7679 could be operated satisfactorily at 450’ C., a number of the other new formulations also showed satisfactory life at this high temperature and at pressures as low as 1800 pounds per square inch. Some of these show an even greater lift in octane number under the more severe conditions, but extended life tests have not been run. Before they can be recommended for full-scale plant, more work is needed to establish yields and hydrogen consumptions. Conclusions
Within the past 18 months, new types of catalyst have been installed in both stages of the Billingham creosote hydrogenation plant. Although the plant was constructed over 21 years ago, flexibility of the hydrogenation process, permitting exploitation of developments in theory and practice of catalysis, has made it possible to keep abreast of the times. All this work relates to hydrocracking coal tar oils obtained predominantly by high temperature carbonization. These starting materials are essentially aromatic and the aim has been to preserve this aromatic character as far as possible. The’hydrocracking process is equally applicable to middle oils obtained from low temperature carbonization processes and from shale retorting. For these oils also, recently introduced catalysts will lead to improved product quality, even though lower aromaticity of the feed stocks will generally cause the products to be somewhat inferior to those from high temperature tar oils. literature Cited (1) Gordon, K., J . Inst. Fuel 9, 69 (1935). (2) Ibid.,21, 53 (1947). ( 3 ) Holrovd, R., J . Junior Znst. Engngrs. 56, ‘325 (1946). (4) Holroyd, R., Proc. United Nations Conf., Conservation and Utilization of Resources, vol. 111, pp. 96-8, 1949. ( 5 ) Imperial Chemical Industries Ltd., Brit. Patent 585,958; U. S. Patent 2,411,726. (6) Oriel, J. A., Gordon, K., World Power Conf. Sect. A3, No. 1, The Hague, 1947. (7) “Petroleum and Synthetic Oil Industry of Germany,” His Majesty’s Stationery Office, *London, England, 1947. (8) Petroleum Times, p. 423 (June 9, 1945). (9) Pier, M., World Petroleum Conf., Sect. III/L, Preprint 2, Rome, 1945. (IO) Pier, M., Z. Elektrochem. 53, 291 (1 949).
RECEIVED for review September 17, 1956 ACCEPTED February 9 , 1957
