CATALYTIC UPGRADING OF REFRACTORY HYDROCARBON STOCKS

Abstract: A process for catalytically upgrading refractory hydrocarbon charge stocks, such as coker gas oil, to premium products including gasoline, kerosine, diesel fuel and low sulfur cracking stock by catalytically hydrotreating said stock and cascading the hydrotreater effluent through a hydrocracking zone containing a catalyst comprising ZSM-5 type zeolite in combination with a metal hydrogenation/dehydrogenation component.

Description: Catalytic upgrading of refractory hydrocarbon stocks This invention relates to catalytic upgrading of refractory hydrocarbon charge stocks, such as, for example, a coker gas oil or catalytic cracking cycle stock, in a dual bed hydrotreating/hydrocracking cascade system. Present refinery practice in upgrading coker gas oil generally involves hydroprocessing to remove hetero-atoms, followed by fluid catalytic cracking or high pressure hydrocracking in a two step operation. Such processes have entailed high equipment costs. Thus, fluid catalytic cracking of even hydroprocessed material requires high severity operation causing excessive coke formation, leading to reduced catalyst life. The necessity to use high pressure has placed an additional economic burden on the overall process. In accordance with the present invention, there has been discovered a dual bed hydrotreating/hydrocracking cascade process for effectively upgrading refractory hydrocarbon stocks characterized by a Bromine No. greater than about 10 and an aromatics content of at least about 40 weight percent, such as coker gas oil and catalytic cracking cycle stock which process can be carried out at substantially lower pressures than those heretofore employed in the aforenoted two stage high pressure process. The present process has been found to provide unexpectedly good yields and selectivity for the conversion of the above characterized refractory hydrocarbon feeds to yield useful products including gasoline of high octane number, kerosene of low freeze point, Diesel or home heating fuel, low in pour point and sulfur and low sulfur distillate suitable for use as a catalytic cracking charge stock. The refractory hydrocarbon feed, contemplated for upgrading in accordance with the present process contains at least about 40 and generally between about 40 and about 70 weight percent aromatics and has a Bromine No. in excess of about 10 and usually in the approximate range of 10 to 60. Exemplary of such feeds are low hydrogen refractory materials, such as catalytic cracking cycle stocks and coker gas oils, which are mostly poly-aromatic in structure and may contain appreciable amounts of sulfur, capable of being effectively removed in accordance with the present process. The dual catalyst bed cascade process of this invention is conducted at a pressure within the approximate range of 790 to 12,512 kPa (109 to 1800 psig) and preferably between about 3549 and about 10443 kPa (500 and about 1500 psig). The temperature is generally within the approximate range of 2880C to 5100C (5500F to 9500F), with an increasing temperature gradient, as the feed passes initially through the bed of hydrotreating catalyst and thereafter through the bed of hydrocracking catalyst. Suitably, the temperature in the hydrotreatlng catalyst bed will be within the approximate range of 288"C to 4540C (5500 to 8500F) and in the hydrocracking catalyst bed within the approximate range of 3430C to 5100C t650 F to 9500F). The feed is conducted through the catalyst beds at an overall space velocity between about 0.1 and about 5 and preferably between about 0.2 and about 2, along with hydrogen initially present in the hydrotreating zone in an amount between about 1000 and about 10,000 standard cubic feet per barrel of feed, corresponding to a ratio of between about 2.4 and about 24 moles of hydrogen per mole of hydrocarbon and thereafter present in the hydrocracking zone in an amount between about 2 and about 23 moles of hydrogen per mole of charge to said hydrocracking zone. The hydrotrcatin catalysts employed are generally metals or metal oxides of Group VIB and/or Group VIII deposited on a solid porous support such as silica and/or metal oxides such as alumina, titania, zirconia or mixtures thereof. Representative Group VIB include molybdenum, chromium and tungsten and Group VIII metals include nickel, cobalt, palladium and platinum. These metal components are deposited, in the form of metals or metal oxides, on the indicated supports in amounts generally between about 0.1 and about 20 weight percent. Initial hydrotreating of the refractory hydrocarbon feed serves to convert sulfur and nitrogen derivatives of hydrocarbon to hydrogen sulfide and ammonia while depositing metal contaminant from hydrodecomposition of any organometal compounds. The entire effluent from the hydrotreating zone containing hydrogen, hydrocarbons, hydrogen sulfide and ammonia is passed to a hydrocracking zone over catalyst containing a crystalline aluminosilicate zeolite, characterized by a silica/alumina ratio greater than 12 and a constraint index, as hereinafter defined, in the approximate range of 1 to 12, such as, for example, zeolite ZSM-5 and a metal having activity to catalyze hydrogenation/dehydrogenation reactions. Representative of the latter metals are those of Group VIII and other metals commonly referred to as transition metals. Thus, the process of this invention provides a dual bed system for upgrading a nitrogen and sulfur-containing hydrocarbon charge by contacting a stream of such charge initially in a hydrotreating zone containing a hydrotreating catalyst under the aforenoted conditions and thereafter passing the entire effluent from the hydrotreating zone to a hydrocracking zone containing a particularly defined crystalline aluminosilicate zeolite-containing hydrocracking catalyst under the above specified conditions of reaction. The latter zeolite-containing catalyst would appear to have the unique ability to bring pressure requirements for the hydrocracking zone to within the range of pressure employed in the hydrotreating zone, i.e., permit use of lower pressures in the hydrocracking zone over those previously employed. Moreoever, the ability of the present process to convert the above stocks, such as coker gas oil to high octane gasoline, which can be directly blended to pool octane gasoline without the need for reforming is a significant feature differentiating the present operation from previous high pressure hydrocracking systems which produce low octane naphtha. The crystalline aluminosilicate zeolites utilized herein are members of a novel class of zeolites that exhibits unusual properties. Briefly, the preferred type zeolites useful in this invention possess, in combination: a silica to alumina mole ratio of at least about 12; a dried crystal density of not less than about 1.6 grams per cubic centimeter and a constraint index of about 1 to about 12. Crystal density constraint index and the measurement thereof are described in U.S. 4,118,431. An important characteristic of the crystal structure of this class of zeolites is that it provides constrained access to and egress from the intracrystalline free space by virtue of having an effective pore size intermediate between the small pore Linde A and the large pore Linde X. Although zeolites with a silica to alumina ratio of at least 12 are useful, it is preferred to use zeolites having higher ratios of at least about 30. Such zeolites, after activation, acquire an intracrystalline sorption capacity for normal hexane which is greater than that for water, i.e, they exhibit "hydrophobic" properties. When synthesized in the alkali metal form, the zeolite is conveniently converted to the hydrogen form, generally by intermediate formation of the ammonium form as a result of ammonium ion exchange and calcination of ammonium form to yield the hydrogen form. In addition to the hydrogen form, other forms of the zeolite wherein the original alkali metal has been reduced to less than about 1.5 percent by weight may be used. Thus, the original alkali metal of the zeolite may be replaced by ion exchange with other suitable ions of Groups IB to VIII of the Periodic Table, including, by way of example, nickel, copper, zinc, palladium, calcium or rare earth metals. Generally, however, the zeolite either directly or via initial ammonium exchange followed by calcination, is preferably hydrogen exchanged such that a predominate proportion of its exchangeable cations are hydrogen ions. In general, it is contemplated that more than 50 percent and preferably more than 75 percent of the cationic sites of the crystalline aluminosilicate zeolite will be occupied by hydrogen ions. The class of zeolites defined herein is exemplified by ZSM5, ZSN-ll, ZSM-12, ZSM-23 ZSN-35 and ZSM-38, defined respectively by the X-ray diffraction data presented in U.S. 3,702,886, U.S. 3,709,979, U.S. 3,832,449, U.S. 4,076,842, U.S. 4,016,245 and U.S. 4,046,859. The preferred zeolite is ZSM-5. In practicing the desired conversion process, it may be desirable to incorporate the above described crystalline aluminosilicate zeolite in another material resistant to the temperature and other conditions employed in the process. Such matrix materials include synthetic or naturally occurring substances as well as inorganic materials such as clay, silica and/or metal oxides. The above zeolite, either as such or after comositing with a matrix, is combined with a minor amount, genially between about 0.1 and about 20 weight percent of a metal having hydrogenation/dehydrogenation promotion properties. Preferred metals are those of Group VIII of the Periodic Table. Palladium is highly effective, as are the other Group VIII noble metals platinum, iridium, osmium, ruthenium and rhodium. Nickel, cobalt, etc. are effective. Other metals, particularly those called transition metals may be employed. The metals may be used alone or in combination, e.g., palladium and zinc. The metals may be incorporated in the finished catalyst by any of the techniques well known in the art such as base exchange, impregnation and the like. Conditions for effective hydrotreating are well known and need no detailed review except to note that cascading the hydrotreater effluent to the second stage requires that sufficient hydrogen be supplied with charge to the hydrotreater zone in order that requirements of both stages shall be satisfied. Pressure in the system described herein will be the same for the hydrotreating and hydrocracking zones and generally within the approximate range of 790 to 12512 kPa (100 to 1800 psig) and preferably between about 3549 and about 10443 kPa (500 and about 1500 psig). Generally, it will be found desirable to employ higher temperature in the second than in the first stage to achieve high conversion to lower boiling products in the second stage. This is accomplished by inter-stage heating. Space velocities for the two stages are adjusted by sizes of the catalyst beds. The following examples will serve to illustrate the process of the invention without limiting the same: EXAMPLE 1 A coker heavy gas oil was processed, along with hydrogen, in a system made up of an initial bed of hydrotreating catalyst and second bed of hydrocracking catalyst. The charge stock had the following properties: Analysis, Elemental, % Carbon 85.9 Hydrogen 11.49 Iron .00006 Nickel < .00001 Nitrogen, Total .13 Nitrogen, Basic .05 Oxygen .11 Sulfur 2.30 Vanadium .00001 Bromide No. 12.4 Carbon Residue, Conradson,% .01 Distillation, OC Vol.% IBP 181 C (3570F) 5 263 (505) 10 292 (559) 20 314 (598) 30 329 (623) 40 340 (643) 50 350 (662) 60 361 (682) 70 374 (705) 80 387 (729) 90 403 (759) 95 415 (779) Gravity, API 22.9 Pour Point, OC 15.60C (600F) Type Analysis, Wt. % Paraffins 20.6 Mononaphthenes 9. 9 Polynaphthenes 19.2 Aromatics 50.4 Viscosity, cs, 400C 9.856 Viscosity, cs, 1000C 2.494 Operating conditions included a pressure of 6996 kPa (1000 psig), a hydrotreating catalyst bed temperature of 371-3890C (700-7320F) , a hydrocracking bed temperature of 424-464"C (796-8680F), hydrogen in the amount of 2500 standard cubic feet of hydrogen per barrel of charge and an overall space velocity of 0.3-0.5. The hydrotreating catalyst was in the form of an extrudate 2 having a surface area of 129 m /gram and containing, expressed as weight percent oxides, the following metals: Cobalt 1.75% Molybdenum 10 % Nickel 1.75% Titanium 30.7 % Zirconium 30.7 % Aluminum 25 % Effluent from the hydrotreating zone was passed directly without interstage separation to a hydrocracking zone containing catalyst. The hydrocracking catalyst was HZSM-5 containing 0.5 weight percent palladium and 0.6 weight percent zinc. This catalyst, having a surface area of 315 m2/gram, was prepared by ion-exchanging HZSM-5 with aqueous Zn (NO3)2, followed by water washing and impregnation of the wet cake with aqueous palladium tetramine nitrate, after which, the composite was dried and calcined at 5380C (10000F). The resulting material was wet slurried with 35 weight percent A1203, extruded in the form of 1.6mm (1/16") particles, dried and calcined in air at 5380C (10000F). The results obtained in runs extending from 6 to 41 days are set forth in the following Table: TABLE 1 Run No. 1 2 3 4 Time on Stream, Days 6 10 14 20 Temperature, C, Ave. Hydrotreating Catalyst Bed 371 (700 F) 372 (702 F) 385 (726 F) 383 (722 F) Hydrocracking Catalyst Bed 424 (796 F) 424 (796 F) 439 (821 F) 438 (820 F) Pressure, kPa 6996 (1000 psig) 6996 (1000 psig) 6996 (1000 psig) 6996 (1000 psig) H2 Flow, SCFB 2500 2500 2500 2500 Space Velocity, vv-1Hr-1 Hydrotreating Catalyst Bed 1 1 1 1 Hydrocracking Catalyst Bed 1 1 1 1 Overall 0.5 0.5 0.5 0.5 Charge Yields, Wt % C1 0 .36 .26 .49 .42 C2 0 .73 .51 .91 .81 C3 0 4.13 3.15 5.54 5.06 C4 0 5.10 4.16 6.52 5.71 C5-166 C (330 F) 0 11.2 11.9 11.7 11.3 166-216 C (330-420 F) .25 3.33 3.09 4.40 3.42 216-281 C (420-538 F) 6.24 12.1 12.4 12.6 12.8 281-366 C (538-690 F) 51.19 39.3 40.4 37.2 38.9 366-430 C (690-805 F) 32.79 18.4 18.6 16.0 16.8 430 C+ (805 F+) 9.55 4.2 4.0 3.44 3.56 NH3 .11 .12 .14 .11 H2O .11 .11 .12 .11 H2S 2.22 2.10 2.27 2.25 Total 101.25 100.94 101.29 101.31 TABLE 1 CONTINUED 1 2 3 4 H2 Consumption, SCFB 718 523 736 756 Conversion to 216 C (420 F), net 24.6 22.0 29.3 26.5 Conversion to 282 C (540 F), net 30.5 29.0 35.7 33.0 Selectivity (C4-216) x 100/216 C- 78.8 82.9 76.3 76.2 (C4-282) x 100/282 C- 82.8 86.5 80.5 81.0 Total Yield of Upgraded Naptha & Distillated (C4-430 C) 89.4 88.0 88.4 88.9 Liquid Analysis, Wt. % Hydrogen 11.36 11.94 11.78 11.72 11.88 Nitgrogen 0.13 .042 .033 .02 .024 Oxygen .11 .012 .018 .01 .01 Sulfur 2.31 .255 .39 .212 .23 TABLE 1 CONTINUED Run No. 5 6 7 Time on Stream, Days 27 34 41 Temperature, C Ave. Hydrotreating Catalyst Bed 383 (721 F) 378 (712 F) 389 (732 F) Hydrocracking Catalyst Bed 451 (844 F) 450 (842) 464 (868) Space Velocity, vv-1 Hr-1 Hydrotreating Catalyst Bed 0.6 0.6 0.6 Hydrocracking Catalyst Bed 0.6 0.6 0.6 Overall 0.3 0.3 0.3 Yields, Wt. % C1 .89 1.05 1.54 C2 1.75 1.99 2.82 C3 11.4 11.3 13.7 C4 11.2 8.4 8.66 C5-166 C (330 F) 15.7 10.8 11.2 166-216 C (330-420 F) 3.2 3.30 3.47 216-381 C (420-538) 11.3 12.0 12.4 281-366 C (538-690) 29.9 33.9 31.2 366-430 C (690-805) 11.8 13.7 11.8 430 C+ (805 F) 2.43 2.89 2.56 NH3 .15 .15 .14 H2O .12 .12 .12 H2S 2.39 2.38 2.36 Total 102.12 102.01 102.05 TABLE 1 CONTINUED Run No. 5 6 7 H2 Consumption, SCFB 1206 1147 1175 Conversion to 420 F, net 43.9 36.6 41.1 Conversion to 540 F, net 48.9 42.3 47.3 Selectivity (C4-216) x 100/216 C- 68.0 60.8 56.1 (C4-282) X 100/282 C- 71.3 66.1 61.8 Total Yield of Upgraded Naphtha & Distillates (C4-430 C) 83.1 82.1 78.7 Liquid Analysis, Wt. % Hydrogen 11.48 11.68 11.36 Nitrogen .013 .014 .018 Oxygen .004 .005 .006 Sulfur .10 .098 .133 It will be seen from the above tabulated data that the same show net conversion as high as 49% to 2820C (5400F-) and 44% to 2160C- (4200F ) . The total yield of naphtha plus distillate C4-4540C (8500F) was as much as 89. Moreover, it will be seen that the catalyst exhibited good stability as evident from results obtained after an uninterrupted on-stream period of 41 days. Separation of representative liquid products was effected by fluorescent indicator adsorption into volume percent aromatics1 olefins and saturates. Results for the liquid products obtained in Run Nos. 1, 5, 6 and 7 of Table 1 are shown below: TABLE 2 Run No. 1 5 6 7 Conversion to 282 C- (540 F-) 31 49 43 48 FIA, vol. % Aromatics 59.0 56.8 57.6 60.3 Olefins 1.1 1.3 1.4 1.7 Saturates 39.9 41.9 41.0 38.0 Results from inspections of appropriate boiling range products are shown in the table set forth below: : TABLE 3 Boiling Range, C IBP 235 C (455 F) 235-335(455-636 F) 335+(636 F+) Wt. % Analysis, wt. % Nitrogen, Total .0043 .0059 .051 Nitrogen, Basic .002 .002 .004 Hydrogen 13.27 11.58 11.04 Sulfur .193 .051 .315 Gravity, API 60/60 - 24.2 PONA, wt. % Paraffins 42.61 8.3 Olefins 8.87 {34.7} Naphthenes 16.86 Aromatics 31.31 57.1 Cetane Number - 36 FIA, vol. % Saturates 57.4 Olefins 11.8 Aromatics 30.8 RONO, micro 85.6 RON3, micro 94.7 Pour Point, C < -53 ( < -65 F) Smoke Point 7.0 Bromine No. 24.6 2.3 It will be evident from the foregoing that useful product fractions show a pool octane quality gasoline containing 31% aromatics (86 RON-Clear); a light gas oil suitable for diesel fuel or No. 2 fuel (.05% sulfur) and a heavier fraction with 0.3% sulfur useful as a cracking stock.

Claims: WHAT IS CLAIMED IS: 1. A dual bed process for upgrading a refractory hydrocarbon feed selected from coker gas oil and catalytic cracking cycle stock and having a Bromine No. greater than about 10 and an aromatics content of at least about 40 weight percent, characterized by contacting a stream of said feed initially in a hydrotreating zone containing a bed of hydrotreating catalyst under reaction conditions which include a pressure within the range of about 790 to 12512- kPa, a temperature between about 2880C and about 4540C in the presence of between about 2.4 and about 24 moles of hydrogen/mole of hydrocarbon charge; passing the entire effluent from said hydrotreating zone to a hydrocracking zone containing a bed of hydrocracking catalyst comprising a crystalline aluminosilicate zeolite characterized by a silica to alumina ratio of at least about 12 and a constraint index within the approximate range of 1 to 12 in combination with a metal component exhibiting hydrogenation/dehydrogenation activity under reaction conditions which include a pressure within the above stated range, a hydrogen concentration between about 2 and about 23 moles of hydrogen per mole of charge to said latter zone and a temperature between about 3430C and about 5100C . 2. The process of Claim 1 characterized in that said hydrocarbon feed contains between about 40 and, about 70 weight percent of aromatics and has a Bromine No. in the approximate range of 10 to 60. 3. The process of any preceeding claim characterized in that said hydrocarbon feed is a catalytic cracking cycle stock. 4. The process of any preceeding claim characterized in that said hydrocarbon feed is coker gas oil. 5. The process of any preceeding claim characterized in that said zeolite is ZSM-5. 6. The process of any preceeding claim characterized in that said metal component is palladium, zinc or a combination of palladium and zinc.

Inventors: GORRING, ROBERT LEE, SMITH, ROBERT LLOYD
Applicants: MOBIL OIL CORPORATION
Application year: 1980
Main classification: C10G65
IPC codes: C10G1, C10G65, C10G47, F02B3