TECHNOLOGY REVIEW
Hydrocracking: A Review N. Choudhary and D. N. Saraf* Deparfrnent of Chemical Engineering. lndian institute of Technology. Kanpur-2080 16. lndia
This paper conforms to the usual concept of an embellished literature survey and as such is useful as a review of technology. I n a n area of such broad scope no authors can be expected to offer a meaningful assessment of differing effects from the so numerous variables. There is value in the paper as an up-dating of events since the notable p u b lications “Aduances in . Hydrocracking” (Aduan. Petrol. Chem. Refining, 8 , 169 (1964)) and “Hydrocracking To-day and Tomorrow” (PetrolChem Eng., 37, 32-35 (Oct 1965)).
THEEDITOR.
During the past decade hydrocracking has taken shape into a major refining process (20, 134, 259, 261, 262). Ever since the announcement of the modern version of hydrocracking (272) the worldwide on-stream capacity has increased steadily (46, 47, 48, 92, 162, 274). At the beginning of the year 1973, the total hydroprocessing capacity of Western Europe was 4,631,000 barrels/day (92), while the total hydrocracking capacity of U.S.A. and Canada at the same time were 865,050 and 51,700 bbl/stream day, respectively (46). On-stream hydrocracking capacity indexes (1967 = 100) for the United States for consecutive years from 1969 to 1972 were 149, 180, 207, and 210, respectively (274, 162, 248, 47, 48). Hydrocracking is e x k n sively practiced commercially in petroleum refining to produce high quality gasoline (187, 248), jet fuel (15, 113, 148, 315, 324), low pour point diesel fuel (42), high quality lubricants (31, 45, 263, 287, 313), and LPG (126, 127, 234). Limited use of hydrocracking to produce light paraffins and aromatics has also been reported (17, 94, 169, 175, 259, 260, 278, 316, 331). Hydrocracking is the most versatile of modern petroleum processes (245, 262, 301). Flexibility of operation in respect of both the feed-stock and product have been reported (106, 171, 243). This provides the most economical refinery balance in respect of demand and supply. This flexibility in operation may be assigned to the development of specific families of the catalyst, the design of processing schemes that allow these catalysts to function efficiently, and the optimum refining relationships between hydrocracking ahd other refining processes. Whereas the commercial feed-stocks range from naphtha to residua, there is a wide choice for the product of a hydrocracker (262). It is possible to produce high octane gasoline at one time and middle distillates such as jet fuel, diesel oil, fuel oil etc. at another time (326).Proportion of motor and jet fuel production can be varied as and when needed ( I 19). Hydrocracking processes for converting polyethylene waste into gasoline and kerosine have been discussed by 74
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Takesue (282b). Processes for hydrocracking of paraffin wax to jet fuel (15) and to a-olefins (303) have been patented. Processes for selective hydrocracking of polycylic hydrocarbons by separating the feed-stocks into a polycyclic rich and polycyclic lean streams are licensed (334, 335). These indicate a fuhher extension to the handling capability of a hydrocracker. Advances in hydrocracking processes have been reviewed from time to time (129, 140, 182, 259, 262, 301, 328). Molenda et al. (161) have described the significance of hydrocracking process in modern refining. Murphy et al. (182) have established the superiority of hydrocracking to fluid catalytic cracking and have discussed their relative effects on overall refining systems. Factors responsible for the recent increase in hydrocracking capacity, refinery scheme incorporating various types of hydrocrackers and its future application under European conditions have been discussed by Winsor (328). Scott and Patterson (259) have considered the commercial processes and their world-wide applications. The hydrocracking of feed-stocks too contaminated to be catalytically cracked, integration of hydrocracking and other refining processes, and advances in residuum hydroprocessing have been discussed by these authors. Scott and Bridge (262) have described the state of hydrocracking technology and their catalytic requirements for a wide range of applications. The relation between the laboratory kinetic measurements and the actual performance in commercial units have also been discussed. Nelson (188-207) has written a series of review articles dealing with various aspects of hydrocracking such as the prediction of quality of the product in terms of their octane rating and characterization factor, the distribution of products and their yields, and design factors including hydrogen requirement, amount of heat liberated in the exothermic reaction, etc. Stormont (270-278) has reported initial commercial success of many processes that followed the modern version of hydrocracking (272). The effect of pressure on hydrocracking over a D-6 type chromia-alumina catalyst was investigated by Abidova et al. ( 5 ) . The hydrocracking activity of the catalyst was evaluated in terms of the extent of conversion into diesel fractions. Between 380 and 400°C an increase in pressure gave a higher yield of diesel fraction while a t higher temperatures (450-475°C) an optimum pressure was reached and further increase in pressure decreased the yield of diesel fraction. Filimonov et al. (95) have experimentally shown that an increase in temperature results in a gradual decrease in the reactions of aromatic and benzonaphthene hydrocarbons, and an increase in the reactions of paraffins and naphthene hydrocarbons, Chernozhukov et al. (57) have reported the effect of operating conditions on the hydrocarbon composition of aviation fuel obtained from hydrocracking of petroleum residues. Manshilin et al. (167) and Musial and Rutkowski (183) have reported the effect
of hydrodynamic conditions on hydrocracking product pattern. It is possible to predict the conversion of petroleum vacuum residue from the hydrodynamic conditions prevailing during hydrocracking (180). Empirical equations obtained by Chernakova et al. (56) can be used to calculate the thermal effects of the product depending on the degree of conversion of the raw material and the consumption of hydrogen. Jet fuel selectivity was studied by Stangland and Kittrell (268). Mathematical descriptions of the product of hydrocracking (338, 339) and use of microkinetic equations to determine the yield (214) have been reported. Using similarity theory, an attempt has been made to advance hydrocracking theory (250). Qader et al. (238) have indicated that by operating the process a t relatively higher severity of hydrocracking, gasoline production can be maximized. The cracking severities depend on the quality of feed-stock, reaction temperature and pressure, and activity and selectivity of the catalyst. Hydrocracking Catalyst These days, it is impossible to conceive of a development of chemical progress in oil refining without the aid of a highly efficient catalyst. Hydrocracking is in no way an exception to it. In fact, its progress has been very much dependent on the development of catalysts of requisite activity and selectivity (32, 33, 245, 259, 262, 286, 301, 386, 328). The history of catalyst development in Europe also bears witness to this fact (52, 70, 212, 228, 265, 266). In a way, the beginning of hydrocracking may be pinned down to the year 1927 when the first Bergius plant for hydrogenating brown coal at Leuna, Germany, was put on stream. This led to supplementary processes for hydrocracking of vaporizable middle oils from coaltar in Germany, England, and a little later, in the United States (133). Pelleted tungsten sulfide was the most successful early hydrocracking catalyst (70, 227). Operating conditions were 250 atm pressure and 400°C temperature with hydrogen circulation rate of 1700-2700 volumes (STP) per volume of oil fed. Multiple bed of catalyst with cold hydrogen injection was used to control temperature of this highly exothermic reaction. The catalyst had very good tolerance for high concentrations of impurities like nitrogen, sulfur, and oxygen compounds, but it failed miserably to produce high octane gasoline even from highly aromatic feeds because of extensive hydrogenation. By 1935, tungsten sulfide on montmorillonite catalyst was developed which succeeded partly to avoid this oversaturation of gasoline but its nitrogen tolerance was poor and hence coal tar oil had to be pretreated. In 1937, ESSO used tungsten sulfide on HF-treated montmorillonite catalyst to obtain aviation gasoline (299). The European coal tar hydrogenator used two-stage operation to obtain good octane gasoline. Unsupported tungsten sulfide was used in the first stage for the removal of nitrogen and its compounds by suitably adjusting the operating parameters so that aromatic saturation was minimized. Tungsten sulfide supported on HF-treated montmorillonite was used in the second or hydrocracking stage (70). Iron on HF-treated montmorillonite catalyst developed by ICI, England, in 1939 proved to be a good catalyst for their second stage hydrocracker. By this time it was known that the hydrocracking catalyst was bifunctional. Nickel on HF-treated montmorillonite catalyst developed by ESSO could produce better antiknock quality gasoline but fairly rapid sulfur poisoning of nickel was its serious limitation. Two-stage hydrocracking, using tungsten sulfide on HF-treated montmorillonite for the first stage to remove sulfur, and nickel on HF-treated montmo-
rillonite for the second stage was commercially practiced during 1940-44 to produce aviation gasoline (299). Further development resulted in nickel on silica-alumina catalyst (70), which was more nitrogen resistant. At a given conversion level of the feed, this catalyst preserved still more aromatics in the gasoline fraction to further enhance its octane rating. Good summaries and review articles on hydrocracking catalyst development are available (32, 262, 286, 301). The hydrocracking catalyst is dual functional. The two functions that it serves are (1) cracking of high molecular weight hydrocarbons, and (2) hydrogenation of the unsaturates formed either during the cracking step or otherwise present in the feed stock. A typical cracking catalyst is silica-alumina and a base or a noble metal serves as a hydrogenating catalyst. In a way, hydrogenation helps cracking. The metal hydrogenation sites keep the acid sites of cracking catalyst clean and active by hydrogenating the coke precursors. This cooperation between these two activities depends on the amount of metal deposited on the acidic support but flattens out after a particular metal loading point is reached (245). Cracking Catalysts. Almost all commercial cracking catalysts are combinations of silica-alumina of one type or another. They are (1) the acid-treated aluminosilicates, (2) amorphous synthetic silica-alumina combinations, and (3) the crystalline, synthetic silica-alumina combinations better known as zeolites or molecular sieves. All of these are high temperature acids and their catalytic activity is attributed to their acidity (26, 109, 110, 159, 167, 168, 174, 183, 186,217, 229, 230, 282, 284, 316). In case of amorphous silica-alumina, both Lewis and Bronsted acid sites operate (23, 134, 159, 168, 174, 217, 320), their ratio being dependent on the degree of catalyst hydration (217, 259). The composition of silica-alumina determines the number of active sites and their specific activity (12). Generally speaking the catalysts having highest number of acid sites are the most active ones (172, 309). During the last five years, interest in zeolite catalyst has been greatly intensified. Zeolites X, Y and mordenite have been the center of attraction (24, 43). There is no consensus of opinion regarding the source of activity and nature of active sites. However, it is generally believed that both the Bronsted and Lewis acid sites are required and that optimum activity is obtained with equal number of each (246, 247). In the understanding of these, significant contributions have been made by infrared and thermal analytical studies ( I I , 51, 304). Zeolite crystal lattice has both small and big holes which are interconnected and regularly spaced; 5 8, and 8 A are the two sizes of “windows” with “rooms” of 13 A. Orderly silica-alumina tetrahedra make the walls of cages and super cages. Whereas amorphous silica-alumina may be viewed as a graft polymer of alumina on a backbone of silica having no ion-exchange capacity, zeolites may be regarded as a copolymer of alumina and silica. Regularly spaced alumina and silica form the backbone of the copolymer. Each atom of aluminum and silicon is joined by a single valence bond to each of four oxygen atoms. There is a negative charge in the copolymer for every aluminum atom because aluminum has a positive balance of 3. These negative sites are usually balanced by a rare earth metal ion to make REX (abbreviation for X-sieves) and REY catalysts (86). These rare earth ions, being trivalent, leave strong dipoles in the catalyst (86, 211) and catalyze coke burning reactions (107). Maher et al. (166) have discussed the relationship between activity and different rare earth cations. Correlation of existing literature data on Ind. Eng. Chem.. Prod. Res. Dev., Vol. 14, No. 2, 1975
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the dependence of the catalytic exchange activity of X and Y zeolites on structure, chemical composition, and temperature has been proposed by Jerocki et al. (139). There is a distribution of strength over the catalyst surface (106, 316), some being stronger than others. Strongest sites are not the most favorable because the slow rate of desorption can retard the catalytic reactions. The most favorable sites are the weak sites which are sufficiently strong to accomplish the desired chemical conversion (105). Amorphous silica-alumina has a higher percentage of the total acidity in its strong acid range (131, 216) and this is one of the reasons for its inferiority to zeolite catalyst. Increased cracking selectivity results from the disappearance of the stTong acid sites. To obtain this selectivity, a hydrothermal treatment is given to the catalyst (85, 180). In this hydrothermal treatment, zeolite is heated up to the desired temperature and then subjected to a steam pressure of 1 atm for 17 hr (180). Weisz et al. (322) have related the selectivity of zeolite catalysts to the size of the cages and super cages present in their sieve lattice. Increased catalyst activity has resulted in more economic processing conditions, such as lower operating pressure or higher through-put rates. This higher activity may be used for increased liquid yields, greater selectivity for midbarrel product, or reduced catalyst inventory (316). Some of the major points of agreement between amorphous and crystalline silica-alumina catalysts are that both owe their activity to their acidity (109, 167, 174, 282, 284), show promotional effect of olefins (85, 217), and act through carbonium ion mechanism (186, 229, 229b). The main points of disimilarity between the two lie in the differences in ion-exchange activity and in acidity distribution as mentioned earlier. Whereas the exchange of sodium in the zeolites with rare earth ions produces very active catalyst, this is not the case with amorphous silica-alumina (211). Zeolite, being a superior cracking catalyst, permits a decrease in the reaction temperature for the same level of conversion of the feed-stock (311). Both activity and selectivity advantages have been reported (211). This superiority of the zeolite catalyst can be attributed partly to 50 to 70 times greater number of “active sites” present in zeolites (105, 135, 180) than that in amorphous silica-alumina. Further, zeolites yield 2 to 3 times higher amounts of isostructuial products than amorphous alumino-silicates (135). This superiority is well born out by Fedorov (93), who claimed that if the amorphous aluminosilicate catalyst presently being employed in 40 catalytic cracking units of U.S.S.R. be replaced by zeolites, there would be a net saving of 40 million rubles by way of additional production of gasoline. Hydrogenating Catalysts. The major hydrogenating components are platinum, nickel, palladium, molybdenum, cobalt etc. These may be altered by promotion with another metal (4, 28, 40, 76, 91, 149, 232, 258, 289) or by some pretreatment such as sulfiding (87, 151). Voorhies and Smith (299) have shown that nickel can produce a higher octane gasoline than either tungsten or iron. Higher boiling product formation with a decreased amount of light hydrocarbons is favored if nickel is replaced by a noble metal. Coonradt et al. (75) reported better selectivity to a liquid product with a platinum catalyst than with rhodium, ruthenium, osmium, or iridium. A better service with Pt-catalyst is ensured if the feed is pretreated to remove sulfur and nitrogen. Oxygenated compounds in feed-stock and conventional air regeneration of the catalyst both contribute significantly to permanent deactivation of the catalyst (184). The relationship between the hydrogenating activity of supported Pt-catalyst and catalyst acidity was studied by 78
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Lanon et al. (157). They observed that with increasing platinum content, a selective adsorption of Pt by acid sites caused a reduction in catalyst activity. With a strong hydrogenating catalyst such as platinum on silica-alumina, extensive direct isomerization resulting in a high ratio of intermediate to low molecular weight paraffins was observed (74). This indicates less secondary splitting which can be used advantageously to minimize the ratio of gas to liquid products and the hydrogen consumption. Nitrogen compounds depress the hydrocracking activity of both the noble metal and base metal catalysts, but noble metal catalysts can better preserve their activity and selectivity, if operated at high temperature to maintain the same level of conversion. Base metal catalysts suffer much deterioration with time (301). Nitrogen causes irreversible loss of activity for Pt on silica-alumina and nickel sulfide on silica-alumina (29, 299). However, it has been possible to regain the activity of Pt-catalyst on certain other supports (299).Both the activities of the hydrocracking catalyst depend on their respective composition in the overall catalyst operating conditions as well as the composition of the charge stock. In any specific application, the best choice of a hydrocracking catalyst demands a peculiar balance between the cracking and hydrogenating-dehydrogenating components of the catalyst. A comparative study of the different hydrocracking catalyst activities has been discussed by Voorhies and Smith (299). Depending on the feed-stock to be handled and the end product desired, not only the proportion of these two components of the hydrocracking catalyst are different but other characteristics of the catalyst like surface area and porosity me also different. As an example, the catalyst characteristics for the production of middle distillates from gas-oil are moderate acidity and strong hydrogenating activity with moderate to high porosity while that for the production of gasoline from the same feed-stock become strong acidity and moderate hydrogenating activity with low to moderate porosity. Effects of silicon dioxide content on physicochemical properties of hydrocracking catalyst were studied by Perezhigina et al. (222). Addition of Si02 to a CoMo04/A1203 catalyst increased its cracking and isomerization ability, conversion and the rate of iso-to-normal hydrocarbons in the gaseous and liquid products. Further, this addition of Si02 increased the mechanical strength of the catalyst by a factor of 1.5. However, it caused a loss of hydrodesulfurization activity of the catalyst. A 20% Si02 catalyst lost its hydrodesulfurization activity more rapidly than a 10% Si02 catalyst. Various factors affecting the selection of catalyst may be summarized as follows. Effect of Feed. The basic constituents in a feed can affect the acid component of hydrocracking catalyst. Therefore, if the feedstock is known to contain basic compounds of nitrogen then zeolite catalysts should be preferred to amorphous silica-alumina since these catalysts are known to be more nitrogen resistant. In the presence of nitrogen, catalyst temperature must be raised to sustain conversion. Normally the feed-stocks are pretreated to remove nitrogen compounds which allows the lower temperature operation of these strongly acidic catalysts. Effect of Acidity of the Catalyst. Acidity of the catalyst affects the temperature requirement of the process but it has limited effect on the type of reactions. By changing the relative hydrogenation activity of the catalyst, acidity can have major effects on the products (72). Scott and Bridge (262) have compared the effect of catalyst acid strength for Arabian straight run vacuum gas-oil hydrocracked over a moderately acidic catalyst in a two-
stage plant, In a single-stage operation, mild acidity catalyst can effectively produce low freezing-point jet fuel and, if operated at higher cut points, low pour-point synthetic diesel from high pour-point feed-stock. If strongly acidic catalyst is employed in the second stage of a two-stage plant, product distribution similar to single stage plant is obtained. However, with lower consumption of hydrogen, superior products like jet fuel having lower smoke point and naphtha of higher octane number are obtained. Effect of Pore Diffusion. Porosity is a very important parameter of a catalyst. If the intrinsic reaction rate is high compared with the diffusive influx of the reactants within the catalyst particle to the “active sites”, the catalyst is not being used effectively. In such cases the reaction rate and selectivity both become functions of particle size (61, 262). This change in selectivity has been used to advantage by Chen and Weisz (54). When processing heavy feeds, catalysts having high pore volume and pore diameter should be used. Increasing catalyst porosity or decreasing the catalyst particle size can alleviate the problem of pore diffusion limitation. However, the use of small particles of catalyst has the attendant problem of pressure over-shooting unless an ebullated bed type reactor is used. In case of residual hydrocracking, ?&-in. catalyst particles have given a better performance than Ihe-in. catalyst particles (101). Empirical correlation to evaluate liquid phase effective diffusivity in fine pores where the pore size is of the order of the size of the adsorbing molecule is available (172, 254). Effect of diffusion factors on the hydrocracking and hydrodesulfurization of Tuimazy Petroleum is reported by Katsobashvili et al. (141). Effect of Catalyst Poisons. Inorganic salts, water, metals, and organic compounds of sulfur, nitrogen, or oxygen present in petroleum act as poisons for the hydrocracking catalysts. Each of these foreign constituents affect the catalyst in a different way. Sulfur compounds inhibit the hydrogenation component while nitrogen compounds inhibit the cracking component of the catalyst. The metal contaminants in petroleum are deposited on the catalyst. These deposits, when in the active sate, promote various dehydrogenating reactions (125, 172) and also increase coke producing tendency (333). Oxygen compounds do not pose any serious problem to the refiner (123) although it is reported to cause permanent catalyst deactivation under certain conditions (184). Effective removal of sulfur and nitrogen is achieved by employing hydrodesulfurization (9, 19, 20, 33, 38, 42, 50, 71, 100, 118, 120, 122, 129, 136, 144, 160, 177, 181, 251, 255, 269) and hydrodenitrogenation (6, 49, 95, 128, 132, 206, 207, 215), respectively. However, beneficial uses of these impurities have also been reported (30, 44, 62, 291, 332). Catalysts for hydrocracking of high sulfur and nitro-
gen containing asphaltic residue have been enlisted by Larson (158) and Vlugter and Spijker (299). Kinetic effects due to poisons in hydrogenation process were also considered by Larson (158). Type of Hydrocracking. Depending upon the feedstock used, two types of hydrocracking are being practiced industrially. If the feed-stock is a heavy distillate obtained from straight run refining or cracking operation, it is called distillate hydrocracking. Residual hydrocracking is the name given to the process if the feed stock happens to be the residue of the straight-run refining. These residuals are usually lower in API gravity and higher in carbon residues and carbon to hydrogen ratio as compared to the distillates. Galbreath and Van Driesen (101) have shown that residual hydrocracking is clearly a different process than the distillate hydrocracking. In residuum hydrocracking, a different type of catalyst is used at rela-
tively higher temperature. This difference in the catalyst employed in residuum hydrocracking is caused by metals and asphaltenes present in the feed stock. While the support of the residual hydrocracking catalyst is mainly alumina, it is only one-fourth alumina and the rest silica in the case of distillate hydrocracking. The residual hydrocracking catalyst should possess surface area and pore volume over two times higher than that of the distillate catalyst. The thermal mode of cracking prevails in the former process where as it is catalytic in the latter process. The selection of the catalyst is, thus, governed chiefly by the type of feed used, the type of product desired, operating conditions of the process, contacting pattern of the reactants, and the type of hydrocracking to be used.
Hydrocracking Reactions Reactions occuring during hydrocracking have been studied extensively (10, 22, 24, 25, 29, 33, 67, 69, 72-75, 87, 88, 235-237, 239-240, 279-280, 327, 337). The chemistry of hydrocracking is essentially the carbonium ion chemistry of catalytic cracking coupled with the chemistry of hydrogenation (285). Almost all the initial reactions of catalytic cracking take place but the presence of hydrogen inhibits some of the secondary reactions. This is why similar products are formed in these two refining process. Like catalytic cracking, very small amounts of C1 and CZ fractions are produced in hydrocracking but unlike catalytic cracking, only saturated products are formed in hydrocracking. This difference is due to the presence of hydrogen and that of the hydrogenating component of the hydrocracking catalyst. Cq fractions rich in isobutane are formed because of a great tendency to form tert-butylcarbonium ion. Also the catalyst remains active for a relatively longer period because olefins and other catalyst deactivating materials are rapidly hydrogenated. This suppresses coke formation. At lower temperature, rapid cracking was accomplished because of hydrogenation and relatively high working pressure. Also the need to regenerate the catalyst was greatly reduced. Reaction temperature and the relationship of acidity to the hydrogenating-dehydrogenating activity are chiefly responsible to decide the product distribution. Although the hydrocarbon processes are complex in nature, yet the overall reaction kinetics can be expressed in simple terms. The mechanism proposed by Qader and Hill (240) to explain product distribution in hydrocracking of gas-oil with nickel (6%) tungsten (19%) sulfide on silica-alumina catalyst envisaged a combination of simultaneous and consecutive bond-breaking reactions followed by isomerization and hydrogenation. They reported that hydrocracking, hydrodesulfurization, and hydrodenitrogenation reactions were first order with the cracking reactions on the acidic sites of the catalyst being rate controlling. In another study of hydrocracking of polycyclic aromatics, Qader (235) observed the sequential occurance of hydrogenation, isomerization, and cracking reaction. Hydrocracking reactions followed first-order kinetics and the kinetic data were compatible with the dual site mechanism of Langmuir-Hinshellwood (132, 235). Langlois and Sullivan (156) have reviewed the chemistry of hydrocracking. When the reactants are paraffins, cycloparaffins, and/or alkylaromatics, the products obtained from both hydrocracking and catalytic cracking are similar, but when the reactant is polycyclic aromatics wide difference in the product from these two refining processes are obtained. This is a comprehensive review and should be referred to for details. In the present article only the classwise salient features of the hydrocracking reactions are discussed. Ind. Eng. Chem., Prod. Res. Dev., Vol. 14, No. 2, 1975
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Reactions of Paraffins. The reaction scheme of hydrocracking of paraffins can be related to the mechanisms proposed for catalytic cracking of paraffins (109, 110, 284). It starts with the formation of olefins from paraffins a t metallic centers and the formation of carbonium ions from these olefins at acidic centers (73, 155, 170, 257, 323). Olefins, being the precursor of carbonium ions, can hydrocrack more readily than paraffins. This explains why a small amount of olefins when added to the feed-stock increases the cracking rate of some of the hydrocarbons (85, 21 7, 285). The reactions of paraffin hydrocracking have been studied extensively (9, 14, 24, 72, 73, 75b, 97, 115, 155, 223, 227, 256, 257, 260). Carbon-carbon bond cleavage, hydrogen transfer, hydrogenation-dehydrogenation, isomerization, disproportionation, and cyclization are the reactions that can take place with a varying degree of importance (156).Extensive catalytic cracking followed by hydrogenation to form isoparaffins appears to be the primary reaction of the paraffins (301). Generally, isoparaffins react in the same fashion as n-paraffins with the only difference in rate which is much more rapid in case of isoparaffins as compared to n-paraffin of the same carbon number. The effluent stream from hydrocracking of n-paraffin contains only a small quantity of isoparaffin products with the same molecular weight as the reactant. This appears reasonable as the isoparaffin formed may readily undergo subsequent reaction (156). A characteristic of this family of hydrocarbons is that with increase in molecular weight of paraffin reactants there is a corresponding increase in the hydrocracking rate (14, 97). Flinn et al. (97) have reported 95% conversion of n-hexadecane a t conditions of operation a t which only 53% n-octane was converted. Another notable feature of paraffin hydrocracking is that there are very high ratios of branched-to-straight chain paraffins in the product. This iso-to-normal ratio of product paraffins is substantially greater than the thermodynamic equilibrium (29, 87, 100). The mechanism for the production of isoparaffins proposed by Flinn et al. (89),Archibald e t al. (14), and Barnet et al. (22) explains the high iso-to normal ratio observed. Hydroisomerization, disproportionation and other side reactions contribute to this high ratio of iso-to-normal paraffins by converting n-paraffins formed during cracking step into branched products. Thomas (283) and Voge et al. (297) have related this high ratio of iso-to-normal paraffins to the catalyst used. A lower iso-to-normal paraffin ratio may be expected, if a strong hydrogenating catalyst is used (123). The preferential adsorption of other families of hydrocarbons on the catalytic sites may inhibit the reactions of n-paraffins in a mixture (24, 82, 302). Use of shape selective zeolite catalyst for the selective hydrocracking of normal paraffins in the presence of branched paraffins have been reported (55, 173, 322). With these catalysts only those molecules of reactant and product which conform to a particular size and shape can arrive a t “active sites” or depart from these active sites. Thus the geometrical configuration of the catalyst plays a limiting role (156). Reactions of Cycloparaffins. Hydrocracking reactions of a number of cycloparaffins have been reported (24, 25, 87, 103-105, 147, 163, 325). A typical cracking reaction known as paring reaction takes place to selectivity remove the methyl groups from the cycloparaffin rings without severely affecting the ring itself (24, 87, 147, 156, 325). This cracking reaction gives one acyclic and one cyclic product containing four carbons less than the parent cycloparaffin. Normally the main acyclic product is isobutane. There is a high ratio of iso-to-normal paraffins pres78
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ent in the product which is far greater than the thermodynamic ratio. A steep fall in both selectivity and cracking rate is reported if Clo cycloparaffin reactants are replaced by Cg cycloparaffins. Over 100 times greater rate of cracking is reported for Clo cycloparaffins than that of Cg cycloparaffins (87, 156). The composition of the product from 1,2-diethylcyclohexane (CH), n-butyl-4C, sec-butyl-CH, isobutyl-CH, and tert-butyl-CH- were found to be similar irrespective of the structure of the reactant. Decrease in temperature results in higher selectivity for the paring reaction of polymethylcyclohexanes. Cracking of cycloparaffins is preceded by the extensive isomerization (87, 156). The hydrocracking of a multiple ring naphthene such as decaline is more rapid than that of the corresponding n-paraffins. Higher iso-to-normal ratios of light paraffins and large yields of single-ring naphthanes are found. At conventional process conditions, single-ring naphthenes are resistant to further hydrocracking. These naphthenes contain a higher ratio of methyl cyclopentane to cyclohexane far in excess of thermodynamic equilibria. The Cs rings are desirable to produce good octane gasoline directly but are not amenable to subsequent reforming needed for further octane improvement (24). Hydrogenolysis of a number of polymethylcyclopentanes over Pt-on-alumina catalyst (103-105) and that of alkycyclohexanes over a Ni-on-alumina catalyst (147) have been reported. In the hydrogenolysis of polymethyl cyclopentanes, demethanation to produce lower polymethyl cyclopentanes, and ring cleavage to produce paraffins were the main reactions. Only little isomerization was observed, while with alkylcyclohexanes, methyl groups were removed successively from alkyl side chains without ring cleavage. Beelen et al. (25) have reported the reactions of cyclopropane with hydrogen on Ni and Cu-Ni alloy powders in a pulse reactor a t 40-350°C. Two parallel routes, namely an addition and a hydrocracking reaction, are reported. Both reactions are affected differently if Ni is alloyed with Cu. This differences is attributed to the formation of different intermediates by cyclopropane addition. Reactions of Alkylaromatics. The hydrocracking of a number of alkyl aromatics on acidic catalysts have been reported (63, 64, 81, 156, 279, 280). Isomerization, dealkylation, paring, and cyclization are the reactions that take place. Normally the effluent stream from the hydrocracking of aromatics shows wider variation and greater dependence on the structure of the reactants than that from the hydrocracking of cycloparaffins. Hydrocracking of alkylbenzenes containing side chains of 3 to 5 carbon atoms gives relatively simple products (87, 97, 279). The larger the alkyl side chains, the more complex the product distribution and a new cracking reaction known as cyclization is observed (279).The presence of a considerable quantity of Cg-C12 polycyclic hydrocarbons specifically tetralins and indanes in the product is intriguing since the formation of these bicyclic species is thermodynamically unfavorable. However, only little tetralin is formed in the hydrocracking of n-butylbenzene. A suitable mechanism to explain the formation of tetralin and indane type hydrocarbons is available (27). Hydrocracking of polyalkylbenzenes gives light isoparaffins and Clo and C11 methylbenzenes as the principal products. Ring cleavage is almost absent and hydrogenolysis to form methane is at a minimum. To explain the product distribution, a mechanism similar to the paring reaction was proposed by Sullivan et al. (156, 279). Mechanisms other than this are also available (63, 64, 81, 329).
Make-up hydrogen Aromatic products from hydrocracking, generally form mixtures approximating thermodynamic equilibrium. This is because these can undergo secondary reactions including alkyl transfer and migration of alkyl groups around the rings. The products from cracking of cycloparaffins, on the contrary, are generally far from thermodynamic equilibrium distribution as they do not react extensively. While the ring is conserved for both the families of hydrocarbons, only alkylaromatics undergo cyclization (156). If nonacidic catalyst is used, successive removal of methyl groups from the side chains is the principal reaction and the isomerization reaction is a t a minimum (27, Figure 1 . Two-stage hydrocracker flow diagram: 1, preheater; 2, 156, 219, 233, 264). Commercial uses of both catalytic and reactor; 3, heat exchanger; 4,separator; 5, fractionator. thermal hydrogenolysis to produce gasoline and naphthalene have been reported ( 17, 94, 169, 270,321). Reactions of Polycyclic Aromatics. Hydrocracking of a Make-up hydroqon Recycle hydrogen number of polycyclic aromatics has been reported (52, 97, I 235-237, 239, 240, 280, 327, 337). The principal reactions are reported to proceed through a multiple mechanism of hydrogenation, isomerization, and cracking (52, 235, 236, 237) and also via a different route of hydogenation, isomerization coupled with alkylation, paring, and cracking reaction (97, 279, 280). Unlike other class of hydrocarbons, the products obtained from hydrocracking and catalytic cracking of polycyclic aromatic are distinctly different. Catalytic cracking of phenanthrene over acidic catalyst produced only coke and small quantities of gases (13) while hydrocracking of Figure 2.One-stage hydrocracker flow diagram. the same gave low molecular weight cyclic products (BO). This difference in the product is caused by the hydrogenbe seriously affected. Zeolite catalyst having a greater tolating component of the catalyst and excess of hydrogen erance for nitrogen compounds than amorphous catalyst present in hydrocracking. After hydrogenation, these aro(41, 300, 301) is preferred in such applications. Hence, a matics which are refractory stocks in catalytic cracking unicracking-JHC process employing zeolite catalysts has are converted into readily cracked naphthenes (285).This an edge over other processes in this regard. hydrogenation is effective only a t higher severity of the reaction. Contacting Pattern. A fixed bed of catalyst with downDi- and polycyclic aromatics joined by only one bond ’ flow of reactants is normally used. Howeyer, an ebullated reactor bed with up-following reactants is used in H-oil rather than by two common carbon atoms are readily processes. It is suitable for processing heavy feed containcleaved by hydrogen and converted into single ring aromatics (223).To produce high octane gasoline, the partial ing metallic contaminants. hydrogenation of polycyclic aromatics, followed by extenReaction Stage. Either a single-stage or a two-stage sive splitting of the saturated rings to form monocyclic process is employed commercially. Single-stage is used advantageously to produce middle distillates from heavy aromatics, is a significant reaction (301). These monocyvacuum gas oils. The two-stage plant is intended mainly clic aromatics enhance the octane rating of gasoline. A to convert high boiling, high nitrogen containing hydrostrong hydrogenating catalyst would undesirably hydrogenate them and thereby necessitate the use of subsecarbons, to gasoline. Whereas the first stage removes basic quent reforming for further octane improvement (301). impurities the second stage hydrocracks the effluents from Activity and selectivity of different catalyst systems the first stage into gasoline. The yield of the desired prodwere investigated by Qader et al. (235). Cobalt sulfide on uct is maximized by suitably selecting the recycle point. silica (low) alumina was found to be the most active cataThe nature of the feed-stock and product desired decide lyst in the hydrocracking of polycyclic aromatic hydrocarthe selection of processing scheme (259, 262). The simplibons. fied form of commonly used flow arrangements is shown in Figures 1 and 2. Most of the processes described herein Hydrocracking Processes employ a flow scheme of one form or the other. Details of Seven hydrocracking processes offered for licensing have flow diagram for each of hydrocracking process may be been enlisted with details of flow diagram, commercial obtained from the literature referred to in general and applications, operating capacity, and operating conditions Scott and Paterson (259) in particular. A brief outline of by Scott and Paterson (259). Extensive references have each of the major hydrocracking processes is presented also been provided. Some of the general points of distinchere. tion among these processes may be summarized as given BASF-IFP Hydrocracking (134, 259, 301). This probelow. cess is especially suitable to handle heavy vacuum Catalysts. As stated earlier, it is possible to modify the straight run or cracked distillates and deasphalted vacuhydrocracking catalyst composition in order to obtain speum residuum to produce gasoline or any desired proporcific advantages in each of the important types of application of gasoline, jet fuel, and diesel oil by using either tions (115, 302). Each process uses a particular catalyst in one-stage or two-stage processes with a number of catathe most economical plant to meet its requirements. The lysts (36, 83, 134). A single-stage process operated either acidity of the strongly acidic hydrocracking catalysts is on a once through or recycle to extinction basis can maxipartially neutralized by nitrogen compounds present as mize middle distillate production from heavy vacuum gas impurity in the feed. Thus the working of the plant may oil. The characteristic functions of catalysts for use in
Ind. Eng. Chem., Prod. Res. Dev., Vol. 14, No. 2, 1975
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middle distillate production were reported by Oettinger and Reitz (213). Several designs have been offered for capacities ranging from 7000 to 20,000 bbl/stream day by Institute Francais du Petrole and Badische Anilin- and Soda-Fabrik AG. Houdry-Gulf Hydrocracking (134, 259). This process is operated in either one- or two-stage mainly to maximize gasoline production from residuum desulfurization and hydrocracking. High end-point furnace oil (79, 80) and light gas-oil (98) have been successfully employed as feed-stock. In the two-stage process, the first stage, operated at mild hydrocracking conditions, is used to lower the level of nitrogen, sulfur, and oxygen compounds. After cooling the effluents from this stage, a hydrogen rich gas stream is separated. The separated liquid is fed into the second stage. The second stage works on a recycle to extinction basis. In the case of one stage plant, the process flow is same as the second stage of the two-stage plant. Many commercial installations have been reported to be on stream (134, 259). Gulf Research and Development Co. and Houdry Process and Chemical Co. Division of Air Products and Chemicals, Inc., are the licensors. H-Oil Hydrocracking (102, 124, 134, 259, 292-293). This process uses heavy residues and asphalts as feed stock to convert them into lighter fractions. It is equally possible to use highly contaminated heavy gas-oil (Hyprocessing) as feed. This process uses an ebullated-bed reactor in which solid catalyst particles are maintained in a state of continuous random motion by up-flowing heavy gas-oils (100). Extremely efficient catalyst-oil-hydrogen contact, isothermal operation, ease of catalyst addition or withdrawal in situ are the chief advantages claimed. In the ebullated bed, since there is no problem of pressure overshooting, a relatively small particle size catalyst can be used. This means a more effective utilization of the catalyst. Many commercial units are in operation (140, 243). Hydrocarbon Research, Inc., and Cities Service Research and Development Co. own the process. ISOMAX Hydrocracking (259, 301). This process is a combination of certain features of the original Isocracking and Lomax processes (7, 259, 301). Both single-stage and two-stage versions of the process are being employed commercially. Feed stocks are generally heavy gas-oils, some having end points around 1100°F and containing metals contaminants (244, 248, 249, 259). The Isomax process is being evolved continuously (16, 77-78, 153, 279, 299, 319). The application of the process to residuum conversion has been particularly successful. There is a considerable reduction in the amount of sulfur and nitrogen in the product as compared to their concentration in the feed. A large number of plants are on-stream with a total capacity higher than most of the processes. This process is patented by Universal Oil Products Co. and Chevron Research Co. Unicracking-JHC Hydrocracking (134, 259, 301). Catalytic cycle oils, coker, virgin and thermal gas oils, and catalytic, thermal, and virgin naphtha are the feed stocks to this process which is used to produce high quality gasoline, jet fuel, and midbarrel products (41, 53, 134, 163). Light liquid product yield is generally 100 to 125 vol 70 and the leaded octane rating of light gasoline is 99 to 101 from all types of feed-stocks (134). This process employs molecular sieve zeolite catalyst in a fixed-bed reactor with normal process conditions in the range of 400-800°F at pressures of 500-1500 psia (300-301). A single stage with recycling to extinction (84) or a two-stage operation may be selected to suit the product requirement (220); 60 vol % of the feed can be con80
ind. Eng. Chem.. Prod. Res. Dev., Vol. 14, No. 2, 1975
verted to gasoline per pass in the first stage while the second stage works a t a conversion level of 50-70 vol 70 per pass on total feed (301). A great number of plants having a design capacity of 2000 to 3200 bbl/stream day are on production. The licensors are Union Oil Co. of California and Esso Research and Engineering Co. Varga Hydrocracking (37, 259). This process was developed by the late Professor J. Varga of Hungary to convert high asphaltene-containing Hungarian Crude residusem into diesel fuels and low sulfur fuel oils (37). It is a two-stage process. A vapor phase hydrocracking step is preceded by a sump liquid hydrocracking step. This process has successfully passed a semi-work trial test in a commercial coal-tar hydrogenating plant. Licensors: Hungerian Peoples' Republic, Budapest and German Democratic Republic, Bohebei, Leipzig. With increasing awareness of the problems created by environmental pollution the use of leaded gasoline is decreasing (90, 210). Hydrocracking, followed by catalytic reforming, is particularly suitable for the production of unleaded gasoline. The recent oil crisis has very badly hit the economy of the oil consuming nations (1, 2, 18, 209, 267, 313). The developing countries can hardly absorb the import bills for oil without upsetting their economic balance. Hydrocracking, with its versatility to convert residues and heavy oils to indispensible products like gasoline, jet fuel, diesel, and kerosine may prove to be a boon to these countries. Literature Cited Aalund, L. R., Oil Gas J . , 71 (50),45 (1973). Aalund, L. R., Oil Gas J . , 7 1 (45),19 (1973). Aalund, L. R., OiiGas J . . 69 (12). 86 (1971). Abidova, M . F., Sultanov, A. S., Midzhuraev, R., Uzb. Khim. Zh., 1 4 (1). 26 11970):Chem. Abstr.. 72. 27,211(19701. Abidova, M . F., Midzhuraev, R., Sultanov, A. S., Uzb. Khim. Zh., 12 (6),40 (1968); Chem. Abstr.. 70, 59,446(1969). Aboul-Gheit, A. K., Abdou, I. K.,J. Petrol. lnst., 59, No. 568, 188
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