THE USE OF LOW ACIDITY HIGH SILICA-TO-ALUMINA RATIO LARGE PORE ZEOLITES FOR DISTILLATE PRODUCTION IN CATALYTIC CRACKING

Abstract: This invention relates to a process for the catalytic cracking of gas oil to produce distillate utilizing a low acidity high silica-to-alumina ratio large pore zeolite having an alpha value no greater than 3.

Description: THE USE OF LOW ACIDITY HIGH SILICA-TO-ALUMINA RATIO LARGE PORE ZEOLITES FOR DISTILLATE PRODUCTION IN CATALYTIC CRACKING This invention relates to a process for the catalytic cracking of gas oil using a particular zeolite catalyst which favors the production of distillate. The present invention provides a process for catalytic cracking of gas oil to produce products boiling in the naphtha and distillate fuel range wherein the gas oil is contacted at elevated temperatures over a catalyst comprising a zeolite, which comprises utilizing as a catalyst a composition comprising a matrix and a crystalline aluminosilicate zeolite characterized by (1) a framework silica-to-alumina ratio greater than about 100; (2) a pore size greater than at least 7 x 10-10m (7 Angstrom) units and (3) an alpha value of no greater than 3. The catalyst used in the process of this invention is highly active for the gas oil conversion while exhibiting low gas and coke selectivities and favoring the production of a 216 to 3430C (420 to 65O0F) distillate as evidenced by a substantially lower gasoline to distillate ratio than is exhibited by commercial catalysts tested under the same conditions. The catalysts utilized in the process of this invention are large pore zeolites, i.e. those materials having a pore size of at least 7 x 10-10m units (7 Angstrom), having a high framework silica-to-alumina ratio, i.e. greater than 100, wherein the zeolite is used in the alkali metal form. Large pore zeolites which can be utilized in the process of this invention are zeolite Y, zeolite beta, and zeolite ZSM-20, providing that the zeolites, particularly zeolite beta and zeolite Y have been subjected to a dealuminization procedure so as to remove part of the aluminum originally present in the framework and to obtain a zeolite having a substantially higher silica-to-alumina ratio. Ultra-stable zeolite Y of varying silica-to-alumina ratios are well known in the art and are disclosed in many United States patents, including U.S. Patent Nos. 4,293,192; 3,449,070; 3,354,077; 3,493,519; 3,513,108; 3,506,400; and 4,036,739. Ultra-stable zeolite Y having a framework silica-to-alumina ratio greater than 100 is also known in the art and is described in an article entitled "DEALUMINATE) FAUJASITE TYPE STRUCTURES WITH SiGyAlO2 RATIOS OVER 100", Journal of Catalysis, Vol 54, pp. 285-288 (1978), The present invention is concerned with utilizing a material such as the dealuminized faujasite Y having a high framework silica-to-alumina ratio, i.e. over 100, (known in the prior art) and converting that material to its alkali metal form by contacting that material with a source of sodium ions in an acidic medium and thereafter utilizing that material in the catalytic cracking of gas oil to obtain a product enhanced in distillate with low coke and gas selectivities. The process of this invention is concerned with zeolites which have large pores, i.e. greater than 7 x 10 10m units (7 Angstrom), which are ultra-stable, which have a high framework silica-to-alumina ratio, i.e. greater than 100, and which have low acidity. The most preferred zeolite is ultra-stable zeolite Y having a silica-to-alumina ratio greater than 100. The catalytic cracking of gas oil with low silica-to-alumina ratio sodium Y (e.g. about 3-6) is known in the art and has been identified as being ineffective not only because of the unacceptable steam stability, but also because of poor coke and gas selectivities. Materials of this type are disclosed in an article entitled "ACIDIC CRYSTALLINE ALUMINOSILICATES" by C.J. Plank et al, I. & E.C. PROD. RES. AND DEV. Vol. 3, No. 3, September 1964. Although this invention is concerned with materials which are steam stable, nevertheless, it is still surprising and unexpected that a high silica-to-alumina ratio Y material in its alkali metal form would have excellent coke and gas selectivities when contrasted to its low silica-to-alumina ratio counterparts. The expression "low acidity" as used throughout the specification and claims is inte ded to mean a zeolite which has an alpha value of no greater than 3 and preferably no greater than 0.3. As is known in the art, the alpha value is an approximate indication of catalytic cracking activity of a catalyst compared to a standard catalyst and it gives the relative rate constant (rate of normal hexane conversion per volume of oxide composition per unit time) It is based on the activity of the highly active silica-alumina cracking:catalyst taken as an alpha of one This test is described in U.S. Patent No. 5,354,078 and in the Journal of Catalysis, Vol. IV, pp. 522-529 (August 1965) In order to obtain the low alpha values of this invention, it is necessary to contact the ultra-stable high sca-to-alumina ratio large pore zeolites, in particular ultra-stable' Y, with Åa source of alkali metal ions The reason for this is that these materials are prepared with a final step which involves treatment of the Y zeolite with an acid and, as such, materials of this type generally have an alpha value of about 10 or higher and this is simply too much activity in order to obtain the particular results of this invention. Quite surprisingly, it has been found that if the ultra-stable high silica-to-alumina ratio zeolites are contacted with the source of alkali metal ions at a pH greater than 7, that the resulting zeolite simply is no longer stable and, in fact, reverts to. an amorphous form. On the other hand, if the zeolite is contacted with a source of sodium ions or other alkali metal ions at a pH less than or equal to 7r and, more particularly, at a pH of 4 to 6, that a very stable and active catalyst is obtained. A convenient method of accomplishing this exchange is to utilize a solution of sodium nitrate Quite obviously, other sources of sodium ions can be used with the only limitation being that the pH should range from 4 to 6 and, more preferably, from 4 to 5. Following contact with the alkali ions, the zeolite is thereafter dried and heated in an atmosphere such as air or nitrogen at atmospheric, subatmospheric, or super-atmospheric pressures for between 1 and 48 hours at temperatures ranging from 100 to 6000C (212 to 1112 F) The aluminosilicate zeolite component prepared in the foregoing manner is combined, dispersed, or otherwise intimately admixed with an inorganic oxide gel which serves as a base, binder, matrix or promoter, in such proportions that the resulting product contains fran 2 to 95% by weight and preferably 5 to 5a; by weight of the aluninosilicate in the final composite. Generally, the catalyst composition is dried between 66 C and 316 C (150 F and 600 F) and thereafter calcined in air or an inert atmosphere of nitrogen, hydrogen helium, flue gas or other inert gas at temperatures ranging from 26000 to 81600 (5000F to 1500 F) for periods of time ranging from' 1 to 48 hours or more -It is to be understood that the active aluminosilicate component can be calcined prior to incorporation with the inorganic oxide gel The aluminosilicate-inorganic oxide gel compositions can be prepared by several methods wherein the aluminosilicate having a particle size of less than 4W microns,' preferably within the range of 2 to 7 microns, is intimately admixed with the inorganic oxide gel while the latter is in a hydros state such as in the form of a hydrosol, hydrogel, wet gelatinous precipitate, or a mixture thereof. Thus, finely divided active aluminosilicate can be mixed directly with a siliceous gel formed by hydrolyzing a basic solution of alkali metal silicate with an acid such as hydrochloric, sulfuric, etc The mixing of the two components can be accomplished in any desired manner, such as in, a ball mill or other types of kneading mills. The aluminosi'licate also may be dispersed in a hydrosol obtained by reacting an alkali metal silicate with an acid or an alkaline coagulant. The hydrosol is then permitted to set in mass to a hydrogel which is thereafter dried and broken into pieces of desired shape or dispersed through a nozzle into a bath of oil or other water-inIniscible suspending medium to obtain spheroidally shaped "bead" particles of catalyst such as described in U.S. Patent No. 2,384,946. The aluminosilicate-siliceous gel thus obtained is washed free of soluble salts and thereafter dried and/or calcined as desired In a like manner, the active aluminosilicate may be incorporated with an aluminiferous oxide. Such gels are well known in the art and may be prepared, for example, by adding ammonium hydroxide, ammonium carbonate, etc., to a salt of nitrate, etc., in an amount sufficient to form aluminum hydroxide which upon drying is converted to alumina. The aluminosilicate may be incorporated with the aluminiferous oxide while the latter is in the form of hydrosol, hydrogel or wet gelatinous precipitatev The inorganic oxide gel may also consist of a plural gel comprising a predominant amount of silica with one or more metals or oxides thereof selected fromm the Groups B, II, III, IV, V, VI, VII and VIII of the Periodic Table. Particular preference is given to plural gels of silica with metal oxides of Groups IIA, IIIB and IVA of the Periodic Table wherein the metal oxide is magnesia, alumina, zirconia, beryllia, or thoria. The preparation of plural gels is well known and generally involves either separate precipitation or copreipitation techniques in which a suitable salt of the metal oxide is added to an alkali.metal silicate and an acid or baser as required, is added to precipitate the corresponding oxide. The silica content of the siliceous gel matrix contemplated herein is generally within the' range of 55 to 100 weight percent with the metal oxide content ranging. from 0 to 45 percent Minor amounts of promoters or other materals which may be present in the composition include cerium, chromium, cobalt, tungsten, uranium, platinum, lead, zinc, calcium, magnesium, lithiuitt, nickel and their compounds as well as silica,. alumina, or other siliceous oxide combination as fines. Of course, the pH must not be allowed to go above 7 during compounding with the matrix. A preferred embodiment is to incorporate the zeolite in the matrix and then exchange the entire composition with alkali metal ions at the pH of 7 or less as previously described. It has been further found in accordance with the invention that catalysts of improved activity and having other beneficial properties in the conversion of hydrocarbons are obtained by subjectlng the treated aluminosilicate to a steam treatment carried out at elevated temperatures of 427 C to 8160C (8000F to 15000F), preferably at temperatures of 5380C to 7040C (1000 F to 1300 F). The treatment may be accomplished in an atmosphere of 100% steam or in an atmosphere consisting of steam and a gas which is substantially inert to the aluminosilicate The steam treatment apparently provides beneficial properties in the aluminosilicats It is particularly preferred, however, that the matrix with which the low activity high silica-to-alumina ratio. large pore zeolite be composited is one which is substantially inert to catalytic cracking1 i.e. it also has an extremely low alpha. Matrices of this type are well known in the art and include alumina, silica, etc. A preferred matrix for the process of this invention is an alumina matrix Cracking operations carried out with the catalyst prepared in accordance with this invention are effective at temperatures ranging from 3nOC to 6490C (7oO0F to 1200 F) under reduced, atmospheric or superatmospheric pressures The catalyst can be utilized in the form of spheroidal particles or beads disposed in a stationary bed, fixed fluid bed, or in the fluid procedures where the catalyst is disposed in a reaction zone to which the catalyst is continuously added and from which the catalyst is continuously removed. A particularly effective cracking process can be achieve when the catalyst is used to obtain the inherent advantages realized in fluid: catalytic processes The following Examples will illustrate the process of this invention In the Examples which follow, the charge utilized was a sour heavy gas oil having the following properties: : THE FEED Properties of 3430C (650 F) Joliet Sour Heavy Gas Oil C (wt %) 85.65 H (wt. %) 12.13 O (wt.%) 0.30 N (wt.%) 0.09 S (wt. %) 2.15 Ash (wt.%) 0.01 Ni ppm 0.5 V ppm 0.5 CCR (wt.%) 0.44 216 C-(420 F) (wt.%) 0.0 216-3430C (420-650 F) (wt. %) 3.5 343-454 C (650-850 F) (wt.%) 52.2 454 C + (850 F) (wt.%) 44.3 The Catalysts CATALYST A A high silica-to-alumina ratio ultra-stable Y zeolite was produced by subjecting commercially available sodium Y (silica-to-alumina ratio of 5) to the procedure described in Journal or Catalysis, Vol. 54, pp. 285-288,. 1978, previously referred to. This procedure involved contacting said zeolite Y with a source of ammonium ions followed by deep-bed calcination at 760 C (1400 F) and re-exchange to the ammonium form. The material is again deep-bed calcined at 815 C (1499 F) followed by treating this material with two-molar hydrochloric acid at 900C (1940F) for 2 1/2 hours. The material obtained had 9 of the crystallinity of the starting sodium Y and had an overall silica-to-alumina ratio of 131. Ammonia desorption studies revealed that the effective silica-to-alumina ratio based on framework aluminum was 283. This material had an alpha of 9.û. CATALYST s Catalyst 8 was prepared by subjecting Catalyst A to sodium base exchange using sodium bicarbonate at a pH of 8 to 9 The resulting material was amdrphous and had an alpha value of 0.01. CATALYST C Catalyst C was prepared by subjecting Catalyst A to sodium exchange utilizing an aqueous solution of sodium nitrate at a pH of 4 to 5. This treatment resulted in a material which had its sodium content raised from 70 ppm (Catalyst A) to 0.4 weight percent. This material retalned 90% of the crystallinity of original low silica-to-alumina ratio sodium Y and had an alpha value of 0.1. CATALYST D Catalyst D was prepared in a manner similar to Catalyst A, with the exception that treatment with hydrochloric acid was carried out for 20 minutes as opposed to 2 1/2 hours The resulting composition had an overall silica-to-alumina ratio of 56, and a framework silica-to-alumina ratio of 118 as determined by ammonia absorption - It had a crystallinity of 90% of the staring low silica-to-alumina ratio zeolite Y. CATALYST E Catalyst E was prepared by subjecting Catalyst D to alkali metal exchange with a sodium nitrate solution at a pH of 4 to 5. Material was obtained which had good crystallinity retention ( > 90% vs. the low Si/Al NaY parent) and an alpha value less than 0.1. CATALYST F Catalyst F was prepared in a manner similar to Catalyst A, with the exception that a 5 normal hydrochloric acid solution was utilized as opposed to a 2 normal solution, and contact was carried out for 5 hours and the entire procedure was repeated an additional two times. A material was obtained which had 90% of the crystallinity of the starting material, had a framework silica-to-alumina ratio approaching infinity and an overall silica-to-alumina ratio of 168. The alpha value of this material was 0.2. Catalysts A, B, C, E and F were evaluated for the catalytic cracking of gas-oil as well as a conventional sodium Y having a silica-to-alumina ratio of 5, a control which contained no catalyst but merely contained Vycor chips, and a commercial equilibrium catalyst comprising about l. rare earth exchanged zeolite Y having a silica-to-alumina ratio of 5. in a silica-alumina matrix (identified as 75F). Also, a physical mixture of 12% REY and t-alumina, steamed at 7600C (14000F) for 4 hours (100% steamj to simulata equilibrium catalyst conditions was tested. Catalysts A, 3, C, E and F were pelleted with g-alumina (65K zeoliter 35 wt. X alumina). When the sodium form of the catalyst was utilized, the alumina was pretreated at a pH of 8 to 9 with sodium bicarbonate. When a catalyst was run in the acid form, e.g. Catalyst A, the &gamma; &gamma;-alumina used during pelIeting was not sodium treated. EXAMPLES 1 AND 2 Over the range of conditions investigated, shown in the tables below, (Net 3430C- (6500F-) + coke) background conversion, ascertained by reacting feed Vycor chips, was 10 wt. % or less. When alkali metal exchange of high silica/alumina ratio was attempted at pH > 7, an amorphous material resulted which was no more active than Vycor. MAMPLES 5-15 A 343 C+ (650 F+) sour heavy gas oil previously described was catalytically cracked over various catalysts in a fluidized bed reactor at temperatures of 505#5 C and at atmospheric pressures Weight hourly space velocities (WHSV) ranged from 1.5 to 5.8. In cyclic runs, 10 minute cracking intervals were employed - each cracking interval being followed by a catalyst regeneration period. Continuous runs cracked charge for 100 minutes prior to coke removal. The results obtained, as well as additional operating parameters, are shown in the following table: TABLE 1 343 C+ (650 F+) Jaliet Sour Heavy Gas Oil Cracking Over Na or H Zeolite Y Cracking Time: 10 min. (except where noted) Temperature, 505 C (9410F); Pressure, 101 kPa (1 atm.) Example 3 4 5 Catalyst A B C Framework SiO2/Al2O3 283 283 283 Acidity (&alpha;) 9 0.1 0.1 WHSVa 3.0 3.0 1.5 Cat/Oil (wt/wt)b 2.0 2.0 4.0 Products (wt. %) Charge C3- - 9.3 4.4 6.6 C41s - 10.8 3.5 1.6 C5-216 C (4200F) - 44.5 29.8 42.2 216-343 C (420-650 F) 3.5 17.9 27.5 23.6 343-454 C (650-850 F) 52.2 9.4 22.2 17.5 454 C+ (850 F+) 44.3 4.0 9.4 4.2 Coke - 4.1 3.2 4.3 (net 3430C- (650 F-) + coke) - 86.1 67.2 77.5 Conversion (wt.%)C Selectivities Gas 24.2 12.2 11.0 Coke 4.9 4.9 5.7 C5-216 C (420 F) (G) 53.6 46.0 56.4 216-343 C (420-650 F) (D) 17.3 37.0 26.9 G/D 3.1 1.2 2.1 TABLE 1 (cont.) 3430C+ (6500F+) Joliet Sour Heavy Gas Oil Cracking Over Na or H Zeolite Y Cracking Time 10 min. (except where noted) Temperature, 5053C (941 F); Pressure, 101 kPa (1 atm.) Example 6d 7 8 Catalyst C F E Framework SiO2/Al2O3 283 "o" 118 Acidity (&alpha;) 0.1 .2 < 0.1 WHSVa 3.0 3.0 3.0 Cat/Oil (wt/wt(h 2.0 2.0 2.0 Products (wt. %) C3- 2.2 4.2 5.0 C4's 0.8 3.8 2.1 C5-216 C (4200F) - 36.6 19.6 27.8 216-343 C (420-650 F) 26.4 24.5 29.7 343-454 C (650-850 F) 25.5 37.2 26.0 454 C (850 F) 6.6 7.6 5.7 Coke 1.9 3.1 3.7 (net 343 C- (650 F-) + coke) 66.7 53.6 67.2 Conversion (wt.%)c Selectivities Gas 4.6 15.5 11.0 Coke 3.0 6.0 5.7 C5-216 C (4200F) (G) 56.8 37.9 42.9 216-343 C (420-650 F) (D) 35.6 40.6 40.4 G/D 1.6 0.9 1.1 TABLE 1 (cont.) 343 CI (650 F) Joliet Sour Heavy Gas Oil Cracking Over Na or H Zeolite Y Cracking Time: 10 min. (except where noted) Temperature, 505 C (941 F); Pressure, 101 kPa (1 atm.) Example 9e 10 11 Catalyst E E NaY Framework SiO2/Al2O3 118 118 5 Acidity 65) < 0.1 < 0.1 < 0.1 WHSVa 3.0 5.8 3.0 Cat/Oil (wt/wt)b 0.2 1.0 2.0 Products (wt. %) EMI13.1 <tb> C3- <SEP> 5.3 <SEP> 4.2 <SEP> #16.9 <tb> C4's <SEP> 1.5 <SEP> 1.6 <SEP> # <tb> C5-216 C <SEP> (420 F) <SEP> 26.7 <SEP> 16.8 <SEP> 40.2 <tb> 216-343 C <SEP> (420-6500F) <SEP> 28.3 <SEP> 26.2 <SEP> 20.1 <tb> 343-454 C <SEP> (650-850 F) <SEP> 29.1 <SEP> 37.2 <SEP> 6.0 <tb> 454 C+ <SEP> (850 F+) <SEP> 7.8 <SEP> 11.1 <SEP> 2.2 <tb> Coke <SEP> 1.3 <SEP> 2.9 <SEP> 14.6 <tb> (net <SEP> 3430C <SEP> (6500F-) <SEP> + <SEP> coke) <SEP> 61.8 <SEP> 49.9 <SEP> 91.5 <tb> Conversion <SEP> (wt.%)c <tb> Selectivities Gas 11.4 12.0 19.1 Coke 2.2 6.0 16.6 C5-2160C (420 F) (G) 44.8 34.8 45.5 216-343 C (420-650 F) (D) 41.6 47.1 18.8 G/D 1.1 0.7 2.4 TABLE 2 3430C+ (6500F+) Joliet Sour Heavy Gas Oil Cracking Over Equilibrium Commercial 75F or REY + &gamma; -Alumina Cracking Time: 10 min. (except where noted) Temperature, 5050C (941 F); Pressure, 101 kPa (1 atm.) Example 12 13 14 15 Catalyst 75F 75F 75F Stmd. REY+ &gamma;Al Framework SiO2Al2O3 - - - Acidity (&alpha;) 0.4 0.4 0.4 WHSVa 2.2 5.3 6.6 4.2 Cat/Oil (wt/wt)b 2.7 1.1 0.9 1.4 Products (wt. %) EMI14.1 <tb> C3- <SEP> 6.0 <SEP> 4.2 <SEP> 2.7 <SEP> #4.4 <SEP> <tb> C4's <SEP> 4.3 <SEP> 2.3 <SEP> 1.8 <tb> C5-216 C <SEP> (420oF) <SEP> 55.7 <SEP> 37.9 <SEP> 26.5 <SEP> 40.9 <tb> 216-343 C <SEP> (420-6500F) <SEP> 17.6 <SEP> 21.7 <SEP> 21.0 <SEP> 21.8 <tb> 343-454 C <SEP> (650-850 F) <SEP> 11.3 <SEP> 22.3 <SEP> 36.7 <SEP> #28.6 <tb> 454 C+ <SEP> (850 F+) <tb> Coke <SEP> 2.3 <SEP> 2.4 <SEP> 1.9 <SEP> 4.3 <tb> (net <SEP> 3430C- <SEP> (6500F-) <SEP> + <SEP> coke) <SEP> 85.4 <SEP> 67.4 <SEP> 52.2 <SEP> 70.3 <tb> Conversion <SEP> (wt. <SEP> %)c <SEP> <tb> Selectivities Gas 12.5 10 8.9 6.5 Coke 2.8 3.7 3.8 6.3 C5-216 C (420 F) (G) 67.6 58.3 52.6 60.2 216-343 C (420-6500F) (D) 17.1 28.0 34.7 27.0 G/D 4.0 2.1 1.5 2.2 TABLE (cont.) a. Based on weight of catalyst in the bed (i.e, zeolita + gamma alumina). b. Weight of catalyst divided by weight of oil pumped prior to regeneration c (Net 3430C- (650 F-) liquids + gas + coke)/(343 C+ (650oF+) change). d. Catalyst used in Examples 4 and 5 was steamed at 7600C (14000F), 4 hrs, 100% steam and used in Example 6. e. Continuous run - no periodic regeneration: 100 minutes cracking time. The results depicted in the above table show the high conversion activity of the equilibrium commercial catalyst (75F-Example 12) and its preference for gasoline product, i.e. 68% naphtha selectivity and gasoline to distillate (G/D) of 4.0. The table also shows that. a conventional low silica-to-alumina ratio sodium zeolite Y, i.e. (Example 11) operating at the same conversion level as 75-F, (85-9) produced excessive amounts. of coke and gas as is obvious from the combined coke and gas selectivities which exceeded 35%. Also the G/D ratio exceeds 2. Producing an ultra-stable zeolite Y having a high silica-to-alumina ratio and testing this material in the acid form (Example 3) showed that it had about the same activity at the same operating conditions as the low silica-to-alumina sodium Y, i.e. Example II but, though it provided much reduced coke selectivity, its gas selectivity was still quite high and the gasoline to distillate ratio (G/D) exceeded 3. However, when the catalysts prepared in accordance with the process of this invention viz. alkali metal exchange in an acidic solution, i.e. Catalysts C and E, were used, much improved gas and coke selectivities were obtained and distillate production was greatly enhanced compared to the commercial catalyst at the same conversion level (Example 4 vs. Example 13 and Example 10 vs. Example 14). Example 8 demonstrates that the framework silica-to-alumina ratio is not narrowly critical once it is over 100. Thus, for example, the results of Example 8 having a framework silica-to-alumina ratio of 118 are very similar to Example 4 which had a framework silica-to-alumina ratio of 285. When the catalyst from Example 8 was used in a repeat experiment under the same conditions but run continuously wfthout periodic regeneration, i.e Example 9, results similar to Examples 8and 4 were obtained indicating little or no catalyst deactivation with the high Si/Al alkali metal exchanged catalysts. In an attempt to avoid. alkali metal exchange to produce low acid activity (&alpha;), Example 7 was- run wherein the catalyst had substantially no measurable framework aluminurtrfr Nevertheless, as can be seenF the product distribution from Example 7 indicated somewhat higher gas and naphtha selectivities and lower distillate at about the same conversion level as a high SiAl Na exchanged Y (Example 10) These results indicate a preference f'or alkali metal exchange over high AL removal to obtain low acidity¯ In order to demonstrate the steam stability of the high silica-ta-alumina ratia sodium Y zeolites, the catalyst which had been used in Examples. 4 and 5 was subjected to steaming at 7600C (14000F) for 4 hours at low steam. This treatment simulates the attainment of an "equilibrium" catalyst condition characteristic of commercial unit operation, X-ray analysis of the steamed catalyst revealed no apparent crystallinity loss compared to its unused, unsteamed high silica-to-alumina ratio sodium Y counterpart. When this material was used in Example 6, it did not exhibit any activity loss- vs. Example 4 at the same operating conditions In fact, compared to Example 4, this material exhibited a somewhat enhanced naphtha selectivity at the expense of gas and coke, distillate selectivity remaining essentially unchanged, indicating that steaming not only did not hurt the catalyst but may actually have improved its performance. REY in conjunction with either silica-alumina or &gamma;-alumina performs similarly at similar conversion levels (Examples 13 and 15). Though gas and coke selectivities differ somewiat, gas plus coke selectivity is about the same for both examples as are the individual naphtha and distillate selectivities. Thus, the advantages observed for alkali metal exchanged high silica/alumina ratio ultra-stable Y compared: to REY catalysts at the same conversion level (Examples 4, 6 and 8 vs. Examples 15 and 15) are not the result of influences due to the non-zeolite catalyst component

Claims: CLAIMS: 1. Process for catalytic cracking of gas oil to produce products boiling in the naphtha and distillate fuel range wherein thegas oil is contacted at elevated temperatures over a catalyst comprising a zeolite, which comprises utilizing as a catalyst a composition comprising a matrix and a crystalline aluminosilicate zeolite characterized by (1) a framework silica-to-alumina ratio greater than about 100; (2) a pore size greater than at least 7 x 10 10m (7 Angstrom) units and (3)- an alpha value of no greater than 3. 2. The process of claim 1 wherein the large pore crystalline aluminosilicate zeolite is zeolite Y, zeolite beta or zeolite ZSM-20. 3. The process of claim 2 wherein the zeolite is zeolite Y. 4. The process of claim 3 wherein the zeolite Y has had its activity reduced by treatment with alkali metal cations at a pH of 7 or less 5. The process of claim 4 wherein the pH is from 4 to 6. 6. The process of claim 5 wherein the zeolite Y has been treated with a solution of sodium nitrate.

Inventors: WALSH, DENNIS EDWARD
Applicants: MOBIL OIL CORPORATION
Application year: 1982
Main classification: C10G11
IPC codes: B01J29, C10G11