US3272734A

US3272734A - Hydrofining and hydrocracking process

Info

Publication number: US3272734A
Application number: US304163A
Authority: US
Inventors: Donald D Maclaren
Original assignee: Exxon Research and Engineering Co
Current assignee: ExxonMobil Technology and Engineering Co
Priority date: 1963-08-23
Filing date: 1963-08-23
Publication date: 1966-09-13
Anticipated expiration: 1983-09-13
Also published as: NL6409703A; GB1071467A; NL6409702A; FR1410457A; GB1034547A; DE1297794B; DE1250039C2; DE1250039B; NO115299B; NL146529B

Description

United States Patent 3,272,734 HYDROFININ G AND HYDROCRACKING PROCESS Donald D. MacLaren, Plainfield, N.J., assignor to Essa Research and Engineering Company, a corporation of Delaware Filed Aug. 23, 1963, set. No. 304,163 13 Claims. or. 208-110) This invention relates to a hydrocarbon oil conversion process. Particularly, it relates to a hydrocarbon oil conversion process comprising an initial hydrofining treatment of a nitrogen-containing feed stock to partially reduce its total nitrogen content, and subsequent hydrocracking of the partially hydrofined feed stock with a nitrogen-tolerant hydrocracking catalyst to produce valuable lower boiling petroleum products. More particularly, the invention is concerned with an integrated hydrofining and hydrocracking process utilizing a hydrocracking catalyst comprising a crystalline alumino-silicate zeolite containing or supporting a platinum group metal.

Hydrocracking has recently become a subject of considerable interest within the petroleum industry because of certain particularized advantages it oiiers over conventional catalytic cracking processes. Chemically, hydrocracking may be considered as a combination of hydrogenation and catalytic cracking, and is eflected in the presence of a suitable bifunctional catalyst capable of simultaneously cracking high boiling hydrocarbons to lower boiling fractions and hydrogenating olefinic and aromatic materials into saturated paraflins and naphthenes. Among the advantages offered by hydrocracking are the ability to selectively convert a wide variety of feed stocks, including refractive heavy aromatic feeds, to lower boiling distillates, with significantly less gas and coke yield and higher quality liquid products than are usually produced by catalytic cracking; the adjustability of hydrocracking selectivity to produce high yields of specific liquid products, e.g. gasoline, middle distillate, etc., and the ability to maintain a high selectivity to the desired products over a wide range of conversion levels up to and including extinction recycle.

Generally, hydrocracking finds its highest utility in cracking hydrocarbons boiling in the heavy naphtha and gas oil range; however, it may also be applied to a variety of feed stocks which may include virgin and catalytic naphthas, gas oils, kerosenes, cycle oils from conventional cracking operations, shale oils, alkyl aromatics, etc. The hydrocracking process is accomplished by passing a high boiling feed stock over a suitable hydrocracking catalyst in the presence of hydrogen and at elevated temperature and pressure suflicient to produce the desired conversion to lower boiling products, e.g. products boiling within the gasoline range. A fixed, moving, or fluidized catalyst bed may be used.

It is well known in the prior art that conventional hydrocracking catalysts may be poisoned to a large degree by contact with nitrogen-containing compounds. The activities of such conventional catalysts (e.g. cobalt molybdate-alumina, nickel sulfide-alumina, and other metallic compounds on alumina and/or silica-alumina) are accordingly reduced with a resultant loss in operating efficiency, less desirable product distribution, and necessarily frequent regenerations. Furthermore, since the reaction temperature necessary to obtain the desired conversion level has been found to be directly proportional to the nitrogen content of the feed stock, the heat requirement for processing high nitrogen content feeds is accordingly increased. In addition, the resulting high temperatures often alter the selectivity of the catalyst thereby causing an increase in the yield of undesirable products, e.g. light gases and coke.

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Because of the sensitivity to nitrogen of the abovementioned conventional hydrocracking catalysts, it has heretofore generally been deemed essential that the feed stock possess certain characteristics; for example, that it be essentially free of nitrogen-containing compounds, e.g. less than about 10 p.p.m., preferably about 1-2 p.p.m. nitrogen; or that it have a substantial aromatic content, e.g. at least about 30%, so as to obtain reasonable yields of good quality gasoline. The feed stocks which are most suited for hydrocracking, however, usually contain an appreciable quantity of nitrogen, e.g. at least about 200 p.p.m., more usually at least about 500 p.p.m., and a reduction in nitrogen content is, therefore, necessary prior to contact with the hydrocracking catalyst. This is conveniently accomplished by hydrofining the feed stock to convert organic nitrogenous compounds into a readily removable form, such as ammonia gas. However, the required severity of this hydrofining pretreatment, which has usually proved quite high, particularly with heavy feed stocks such as cycle oils, residuals, etc., has inherent disadvantages. For example, the severity of hydrofining conditions required to obtain a negligible level of nitrogen content is both costly and time-consuming. It has been particularly found that the most diflicultly convertible fraction of organic nitrogen compounds is represented by a total nitrogen content of below about 50 p.p.m., and that conversion and removal of this fraction requires a very extensive hydrofining treatment with an increase in the severity of hydrofining conditions to an even greater degree than is initially required. As a result of these necessarily severe hydrofining conditions, partial conversion of the feed stock is difficult to avoid, since hydrocracking conditions may well be approached. It will thus be readily apparent to those skilled in the art that premature conversion in the hydro-fining stage may adversely affect the desired conversion, selectivity, and quality (e.g. octane number) of the products leaving the hydrocracking stage, due to alteration of the composition of the hydrofined feed. For example, severe hydrofining of the feed stock may result in premature hydrogenation of multi-ring aromatics to form condensed naphthenes, with resulting overconversion in the hydrocracking stage and a lower gasoline yield being thereby obtained. It is apparent, therefore, that a reduction in the severity of hydrofining conditions which would reduce the nitrogen content of the feed stock to an intermediate range, as opposed to an essentially nitrogen-free state, would be economically attractive and would provide product distribution and quality advantages, provided that the activity of the hydrocracking catalyst could be suitably maintained in the presence of the partially hydrofined nitrogen-containing feed.

The present invention is concerned with an improved hydrocarbon conversion process involving an initial hydrofining stage and a subsequent hydrocracking stage, wherein the hydrofining stage is performed under relatively mild operating conditions as compared to the severe conditions conventionally utilized, so as to obtain an essentially unconverted hydrofined efiluent containing an intermediate range of nitrogen content as opposed to a substantially nitrogen-free etfluent; and wherein the hydrocracking stage is performed in the presence of a recently developed nitrogen-tolerant hydrocracking catalyst capable of maintaining a relatively high catalytic activity in the presence of a substantial amount of nitrogen. By means of this process, a substantial economic saving is realized over the aforementioned conventional processes which involve essentially complete removal of nitrogen rather than the partial removal herein contemplated. Additionally, the overall performance of the process is unexpectedly improved by hydrocracking the partially hydrofined feed containing the predetermined intermedi- 3 ate nitrogen content range, as evidenced by superior selectivity and distribution of products in the efiiuent from the hydrocracking stage.

The process of the present invention comprises an integrated hydrofining and hydrocracking treatment of a suitable feed stock substantially (e.g. 95%) boiling above about 400 F. and initially containing at least about 100 p.p.m. nitrogen. The hydrofining step is effected in the presence of hydrogen and a conventional hydrofining, i.e. hydrogenation, catalyst, at elevated temperature and pressure sufiicient to convert the desired proportion of nitrogen compounds to a readily removable form of nitrogen, e.g. ammonia, without substantial conversion of the hydrocarbon materials in the feed. The hydrofined effluent produced should have an intermediate nitrogen content range of about 15 to about 200 p.p.m., preferably about 15 to about 100, more preferably about 15 to about 50 p.p.m. total nitrogen. This desired intermediate range of nitrogen content should be produced with essentially no change, e.g. less than about a 5% reduction in the total aromatics content of the feed. However, hydrogenation of polycyclic aromatics to monocyclic aromatics is permissible, if not desirable, provided that the formation of condensed naphthenes is avoided and the total aromatics content thereby remains essentially unchanged. Additionally, conversion of 430 F.+ hydrocarbon fractions in the feed to fractions boiling below about 430 F. should be minimized in order to avoid overconversion in the hydrocracking stage. It will be appreciated, however, that some production of lower boiling fractions will be inevitable due to the general lowering of the boiling point usually associated with hydrogenation, desulfurization and denitrogenation reactions, and that the actual quantity of lower boiling fractions produced in the hydrofining stage will be dependent upon the characteristics of the particular feed used. Specifically, conversion of the feed hydrocarbons in the hydrofining stage to fractions boiling below about 430 F. should be limited to less than about volume percent, preferably less than about 5 volume percent, based on the volume of feed.

After separation of the ammonia gas formed in the hydrofining stage, the hydrofined feed is subjected to hydrocracking to convert the high boiling petroleum fractions into valuable lower boiling fractions. Hydrocracking is effected at elevated temperature and pressure in the presence of added hydrogen and a hydrocracking catalyst comprising a crystalline alumino-silicate zeolite having a platinum group metal deposited thereon or incorporated therein. As hereinbefore mentioned and to be hereinafter illustrated, this hydrocracking catalyst is significantly more nitrogen-tolerant than the conventional hydrocracking catalysts of the prior art and is characterized by a high activity maintenance in the presence of nitrogen compounds thus enabling the initial hydrofining treatment to be conducted under less severe conditions than conventionally employed. The catalyst is also characterized by a reversible nitrogen poisoning effect; i.e. it has been found that any poisoning action of a high nitrogen feed is only temporary depending upon continued contact with nitrogen and that the catalyst will recover its original activity when subsequently contacted by a lower nitrogen content or nitrogen-free feed. Additionally, the yield, distribution and quality of the products obtained with this catalyst in the presence of nitrogen are markedly superior to those obtained with conventional catalysts. In summary, the zeolitic catalyst employed in the hydrocracking stage of the present process exhibits superior properties as compared to the conventional hydrocracking catalysts of the prior art, which properties include high activity and activity maintenance, nitrogen-tolerance, regenerability, and high selectivity to desired products, with accompanying economy of operation. The use of this catalyst determines to a large degree the successful operation of the process of the present invention and is, therefore, regarded as critical.

The process of the present invention is applicable to a variety of feed stocks having higher initial boiling points than gasoline, i.e. higher than about 400 F., and boiling ranges of about 405 to 1100 F. or higher, preferably 405 to 850 F. For certain applications, e.g. for middle distillate or jet fuel production, 5% boiling points of about 500650 P. will be desirable. Additionally, in order to be advantageously used in the present process, the feed stocks should have a nitrogen content of at least about p.p.m., preferably at least about 200 p.p.m., and most preferably at least about 500 p.p.m. While those feeds having nitrogen contents between 100 and 200 p.p.m. will already have the desired intermediate range of 15 to 200 p.p.m. nitrogen, it will still be advantageous to hydrofine them down to the more preferred ranges of 15 to 100 p.p.m. or 15 to 50 p.p.m. Examples of suitable feed stocks include whole crudes, heavy naphthas, gas oils, kerosenes, refractory catalytically cracked cycle stocks, high boiling virgin and coker gas oils, residual oils, etc. These materials, when subjected to hydrocracking, are converted to gasoline, middle distillate, etc. as determined by control of the hydrocracking operating conditions.

The hydrofining stage of the present process is effected in the presence of a conventional hydrofining or hydrogenation catalyst. Such catalysts are well known in the art and are characterized by their resistance to poisoning, particularly sulfur poisoning. Suitable catalysts of this type include the oxides or sulfides of metals of Groups VI and VIII of the Periodic Table, either alone or in admixture with each other, composited with an inert metal oxide support, wherein the metal oxide is selected from the group consisting of silica, oxides of metals in Groups II-A, IIIA, and IV-B of the Periodic Table, and mixtures thereof. Examples of the Group VI metals are tungsten and chromium, with the preferred Group VI metal being molybdenum in the form of a molybdate; examples of the Group VIII metal components are cobalt and nickel; and examples of the metal oxides in the inert support are alumina, zirconia, magnesia, titania, ceria, thoria, etc. In particular, such catalysts are typified by nickel sulfide-tungsten sulfide, molybdenum sulfide or oxide, combinations of metal sulfides or oxides such as ferric oxide, molybdenum oxide or sulfide, and cobalt oxide, all of which are supported on the above inert metal oxide supports. A particularly prefered catalyst is cobalt molybdate on alumina. Preferred composite catalysts will contain from about 1 to about 6, preferably 3 to 4 wt. percent of a Group VIII metal oxide, e.g. cobalt oxide, and from about 9 to 18, preferably 12 to 15 wt. percent of a Group VI metal oxide, e.g. molybdenum oxide (M00 on a metal oxide support, e.g. alumina or silica-alumina.

The hydrofining treatment preferably involves a fixed bed operation with the oil feed flowing downwardly over the catalyst bed. Suitable operating conditions for hydrofining the feed stock in accordance with the present invention will of course vary with the nitrogen content of the raw feed in order to produce the desired intermediate range of nitrogen content in the hydrofined efiluent to be fed to the hydrocracking stage. In general, hydrofining operating conditions for a feed stock initially containing from about 100 to 3000 (or higher) p.p.m. nitrogen will include a temperature of from about 500 to about 800 F., preferably 625 to 725 F.; a pressure of from about 500 to about 3000 p.s.i.g., preferably 1200 to 2000 p.s.i.g.; a liquid hourly space velocity of from about 0.1 to about 10, preferably 0.1 to 3 volumes of feed per volume of catalyst per hour; and a hydrogen gas rate of from about 1,000 to about 20,000, preferably 2,000 to 12,000 standard cubic feet (s.c.f.) per barrel of feed. It will be understood that with nitrogen-containing feed stocks having nitrogen contents higher than about 3000 p.p.m. a greater degree of hydrofining will *be required to reduce the nitrogen level to within the desired range. This may be accomplished by increasing the severity of hydrofining conditions of temperature and pressure or by reducing the space velocity, i.e. residence time, provided that substantial conversion of the hydrocarbon materials in the feed is avoided. If substantial hydrocarbon conversion cannot be avoided, the hydrofining conditions may be maintained at a suitably mild level and the hydrofined effluent recycled to the hyd-rofiner until the nitrogen content of the feed is reduced to within the desired range.

Subsequent to the hydrofining stage, the ammonia and hydrogen sulfide formed from the hydrogenation of nitrogen and sulfur compounds are removed prior to introduction of the hydrofined stream into the hydrocracking stage by conventional gas separating means. The partially hydrofined oil stream, freed of ammonia and hydrogen sulfide and having a nitrogen content within the aforementioned range, is then preferably passed downwardly over a fixed bed of the zeolitic hydrocracking catalyst to be hereinafter fully described.

Hydrocracking of the partially hydrofined feed stock is accomplished in the presence of added hydrogen and catalyst at a temperature of from about 450 to about 800 F., preferably 600 to 750 F; a pressure of from about 500 to about 3000 p.s.i.g., preferably 1200 to 1800 p.s.i.g.; a liquid hourly space velocity of from about 0.1 to about 10, preferably 0.5 to 3 volumes of feed per volume of catalyst per hour; and a hydrogen feed rate of from about 1000 to about 20,000, preferably 2000 to 12,000 s.c.f. per barrel of feed. The conversion, expressed as volume percent conversion of fractions boiling above about 430 F. to products boiling below about 430 F. for the production of gasoline (or to products boiling below about 500-650 F. for the production of higher boiling distillates), will generally be maintained at about 30 to 80%, preferably 50 to 75%, which conversion is readily attainable at operating conditions within the above ranges. Such conversions have generally been found to achieve Optimum economy of operation. However, for certain desired results, higher or lower conversion ranges may also be utilized, as will be determined by appropriate adjustments of the various operating variables, i.e. temperature, pressure, and space velocity. Such adjustments will be readily determinable by those skilled in the art with a view towards obtaining the desired selectivity of products, product distribution, and product quality, all of which are functions of the operating conditions and conversion levels utilized in the hydrocracking reactor.

The hydrocracked efiluent is fractionated to separate the desired products, and the high boiling fractions are preferably recycled to extinction by return to the hydrocracking reactor. Thus, as the process proceeds, all material boiling above the desired cut-off point is preferably eventually converted to lower boiling products with in the desired ranges. The cut-01f points will, of course, vary depending upon the boiling range of the products desired; e.g. naphtha (boiling in the range of about 375 to 425 E), jet fuel (boiling in the range of about 450 to 550 F.), middle distillate (boiling in the range of about 650 to 750 F.), etc.

The hydrocracking catalyst specifically useful in the process of the present invention comprises a crystalline metallo alumino-silicate zeolite, well known in the art as a molecular sieve, having a platinum group metal (e.g. palladium) deposited thereon or composited therewith. These crystalline zeolites are characterized by their highly ordered crystalline structure and uniformly dimensioned pores, and have an alumino-silicate anionic cage structure wherein alumina and silica tetrahedra are intimately connected to each other so as to provide a large number of active sites, with the uniform pore openings facilitating entry of certain molecular structures. It has been found that crystalline alumino-silicate zeolites having effective pore diameters of 6 to A., when composited with the platinum group metal, and particularly after base exchange to reduce the alkali metal (e.g.

6 Na O) content of the zeolite to less than about 10 wt. percent, are effective hydrocracking catalysts, particularly for the hydrofined feeds herein contemplated, i.e. those having nitrogen contents within the aforementioned ranges.

Naturally-occurring large pore crystalline alumino-silicate zeolites may be exemplified by the mineral faujasite. Synthetically produced alumino-silicate zeolites having large pore diameters are also available and will be preferred in the present invention. In general, all crystalline alumino-silicate zeolites, in natural or synthetic form contain a substantial portion of an alkali metal oxide, normally sodium oxide.

More specifically, the support for the hydrocracking catalyst used in the present invention is a crystalline alumino-silicate zeolite having an effective pore diameter of about 6 to 15 A., wherein a substantial portion of the alkali metal, e.g. sodium, has been replaced with a cation (either a metal cation or a hydrogen-containing cation, e.g. NHJ) so as reduce the soda (Na O) content to less than 10 wt. percent and preferably to about 1 to 5 wt. percent (based on zeolite). The anhydrous form of the base-exchanged large pore crystalline alumino-silicate zeolite prior to compositing with platinum group metal may be generally expressed in terms of moles by the formula:

Wherein Me is selected from the group consisting of hydrogen and metal cations (so that the Na O content is less than 10 Wt. per-cent of the zeolite), n is its valence and X is a number from 2.5 to 14, preferably 3 to 10 and most preferably 4 to 6. Crystalline zeolites having these ratios have been found to be highly active, selective and stable.

The processes for synthetically producing crystalline alumino-silicate zeolites are well known in the art. They involve crystallization from reaction mixtures containing: A1 0 as sodium aluminate, alumina sol and the like; SiO as sodium silicate .and/ or silica gel and/ or silica sol; and Na O as sodium hydroxide. Careful control is kept over the soda (N 0) concentration of the mixture, as well as the proportions of silica to alumina and soda to silica, the crystallization period, etc., to obtain the desired product. A conventional scheme for preparing a crystalline alumino-silicate zeolite having a silica to alumina mole ratio of about 4 to 6 is as follows:

Colloidal silica is mixed with a solution of sodium hydroxide and sodium aluminate at ambient temperature to produce a reaction mixture having the following molar ratios of reactants:

TABLE I Reactants: Molar ratio N|a O/SiO 0.2 8 to 0.45 SIO /Al O 8 to 30 H O/Na O to The reaction mixture may then be allowed to digest at ambient temperatures for up to 40 hours or more in order to aid crystallization, after which period it is heated at to 250 F., e.g. 200 F., for a sufficient time to crystallize the product, e.g. 24 to 200 hours or more. The crystalline, metallo alumino-silicate is separated from the aqueous mother liquor by decantation or filtration and Washed to recover a crystalline product.

For hydrocracking use the zeolite is preferably baseexchanged with a hydrogen-containing or metal cation to reduce the soda content to below 10 wt. percent. Suitable metal cations include ions of metals in Groups I to VIII and rare earth metals, and preferably metals in Groups II, III, IV, V, VI-B, VII-B, VIII, and rare earth metals. Where a hydrogen-containing cation is used to replace the sodium, the hydrogen or decationized form of the zeolite is produced. A convenient method of preparing the hydrogen or decationized form is to subject the zeolite to base-exchange with an ammonium cation solution followed by controlling heating at elevated temperature, eg 600 to 1000 F., to drive off ammonia and Water.

The base-exchanged crystalline zeolite is then composited with the platinum group metal by treatment (e.g. wet impregnation or base-exchange) with a platinum or palladium salt or ammonium complex, e.g. ammonium chlonoplatinate, palladium chloride, etc. The amount of catalytic metal in the finished catalyst is ordinarily between 0.01 and about 5.0 wt. percent, preferably 0.1 to 3.0 wt. percent, based on the zeolite. Normally the catalyst is subjected to a heat or hydrogen treatment at elevated temperatures, e.g. 500 to 1500 F., to reduce the platinum group metal, at least in part, to its elemental state.

The zeolitic catalyst may be utilized in the abovedescribed form or may be suitably embedded in an amorphous material such as silica gel, or a cogel of silica and at least one other metal oxide wherein the metal is selected from Groups II-A, III-A, and IV-B of the Periodic Table, e.g. alumina, titania, magnesia, etc. The use of such composite materials, eg crystalline noble metalcontaining zeolite embedded in .a silica-alumina cogel matrix, has been found useful in certain applications. For example, the activity of the pure crystalline zeolite is very often too high for some applications and to achieve desired selectivities in the hydrocracked product streams it is often desirable to utilize a composite form of zeolitemat-rix catalyst. Additionally, for fluidized and moving bed operations, the crystals of zeolite are often too fine for successful fluidization due to excessive attrition and carry-over losses. The use of a composite catalyst comprising zeolite crystals embedded in a suitable matrix, e.g. a silica-alumina matrix, obviates these difiiculties since the composite material can be formed into particles of a desired size range. A convenient means of forming the composite form of catalyst is to incorporate the zeolite crystals into a suitable hydrogel, erg. a silicaalumina hydrogel; subject the resulting mixture to high agitation conditions with added water, if necessary, to produce a homogeneous fluid dispersion; and finally spray dry the resulting mixture to form particles of the desired size. The final composite may contain to 80 wt. percent zeolite.

The process of the present invention will now be described in further detail, with reference being made to the accompanying drawing which is a simplified fiow diagram of a preferred embodiment of the process.

A suitable oil feed charge is introduced through line 10 and is mixed with hydrogen being introduced through line 11. The resulting admixture of feed charge and hydrogen is heated to hydrofining temperature in heater 12, and passed through line 13 into hydrofiner 14, where it fiows downwardly over hydrofining catalyst of the type hereinbefore described. The hydrofining operating conditions are determined by the nitrogen content of the feed charge and will be sufiicient to reduce the nitrogen content to within the desired range without substantial conversion of the hydrocarbons in the feed. In hydrofiner 14, the nitrogen-containing and sulfur-containing compounds in the feed are substantially converted to ammonia and hydrogen sulfide, respectively. The total effluent from hydrofiner 14 containing excess hydrogen, hydrogen sulfide, ammonia and substantially unconverted liquid petroleum hydrocarbon fractions is withdrawn through line 15, and passed to gas separator 16, to separate the gaseous phase from the liquid phase. The gaseous phase, including hydrogen, ammonia and hydrogen sulfide, passes through line 17 into gas treater 18, wherein the hydrogen sulfide and ammonia are removed by scrubbing or other suitable means. The remaining gas stream composed substantially of hydrogen is recycled to the system via line 19 which joins line 11. Alternatively, the hydrogen sul- 8 fide and ammonia may be removed by a gas treater in line 15.

The essentially unconverted liquid petroleum stream is withdrawn from gas separator 16 through line 20, and passed via lines 21 and 22 into heater 23, wherein it is heated to hydrocracking temperature. The heated liquid stream passes through line 24, mixes with a hydrogen stream flowing through line 25, and the combined admixture of hydrogen and oil streams enters hydrocracker 27 through line 26, where it fiows downwardly over hydrocracking catalyst of the type hereinbefore described. The hydrocracking operating conditions in hydrocracker 27 are sufficient to obtain the desired degree of conversion per pass. The etliuent stream from hydrocracker 27 is withdrawn through lne 28, and passes to gas separator 29, which serves to separate the liquid and gaseous phases. The gaseous phase containing hydrogen, and minor amounts of ammonia and hydrogen sulfide leaves gas separator 29 through line 30 and passes to gas treater 31, where the ammonia and hydrogen sulfide may be removed by scrubbing or other suitable means. The use of gas treater 31 is optional depending upon the amounts of ammonia and hydrogen sulfide in the recycle gas. The gas stream composed substantially of hydrogen is then recycled through line 32, along with make-up hydrogen being introduced through line 33. The combined hydrogen stream is heated to the desired operating temperature in heater 34, passes through line 25, and joins with the partially hydrofined oil feed in line 26 as hereinbefore described. The use of heater 34 is optional.

Liquid petroleum products are withdrawn from the bottom of gas separator 29 through line 35, and are conducted to fractionator 36, wherein light hydrocarbons and gases are removed as overhead through outlet 37. Alternatively, a stabilizer can be utilized in line 35 for the removal of light hydrocarbons and gases. Intermediate fractions of light naphtha, heavy naphtha and light fuel oil may be withdrawn respectively from fractionator 36 through

lines

38, 39 and 40. The heavy oil bottoms product is withdrawn through outlet 41 and recycled through line 42, where it joins with the hydrofined separated liquid oil stream in line 20 for recycle to hydrocracker 27 via

lines

21, 22, 24 and 26. The process is thus characterized by a recycle to extinction of the fractionator bottoms which are converted to lower boiling fractions in hydrocracker 27.

It is to be understood that the process of the invention is subject to minor variations and is not to be limited to the specific embodiment described. For example, while the present process as specifically described, includes a single hydrofining stage and a single hydrocracking stage, it will be realized that additional stages may be included. It may thus be preferable for certain applications to utilize a multiple stage hydrocracking operation, with the fractionator bottoms product being hydrocracked in a separate hydrocracking reactor, rather than the recycle to extinction process hereinbefore described. Such conventional processing variations, being well within the understanding of those skilled in the art, are intended to be included within the scope of the present invention.

The invention will be further understood by reference to the following examples, which are not intended to be limiting.

Example I A crystalline alumino-silicate zeolite hydrocracking catalyst was prepared by the following procedure:

A mixture of sodium aluminate, sodium hydroxide and silica sol was heated at about 212 F. in such proportion and for such time as to produce a crystalline zeolite. Specifically, a slurry mixture of 35 grams of commercial sodium aluminate containing 65 wt. percent NaAlO and grams of sodium hydroxide (97% NaOH), contained in 447 grams of water, was added with rapid stirring to 846 grams of a commercially available colloidal silica sol containing 30 wt. percent silica (Ludox'solution, supplied by E. I. du Pont de Nemours & Co. Inc.). Mixing was conducted at ambient temperature of 75 F. Stirring was continued to form a homogeneous mixture and the composite slurry was then kept at ambient temperature with stirring for a digestion period of about 1 /2 hours, after which time it was heated in an open vessel to 212 F. over a period of 30 minutes. The vessel was then sealed to eliminate evaporation losses and heated at 212 F. for a period of 69 hours to attain maximum crystallinity as determined by periodic sampling and X-ray diffraction analysis. The product crystals were then filtered, and ion-exchanged with a 28 wt. percent aqueous ammonium sulfate solution by slurrying 100 grams of crystals in about 350 grams of solution followed by filtering and water washing. After three ion-exchange and washing operations the resulting ammonium form of the zeolite was separated by filtration and treated with a solution containing the ammonium complex of palladium chloride, in an amount suflicient to produce about a 0.5 Wt. percent palladium catalyst. The catalyst was then filtered, washed, oven-dried at 212 F, and finally calcined at 1000 F. It contained 0.48% palladium, had a uniform pore size of about 15 A. and a silica to alumina mole ratio in the zeolite component of about 55:1.

The above catalyst was utilized in the integrated hydrofining-hydrocracking process of the present invention using the processing scheme hereinbefore described and illustrated in the accompanying drawing with the exceptions that ammonia and hydrogen sulfide gases were removed from the system by water scrubbing the streams in lines 20 and 22, and a stabilizer was used in line 35 to remove light hydrocarbons and gases (C4 The hydrofining catalyst utilized was a cobalt molybdate on alumina catalyst containing 16.5 wt. percent cobalt molybdate (3.9 wt. percent 000, 12.6 wt. percent M The feed stock contained 758 p.p.m. nitrogen and was a blend of 28.7 vol. percent coker gas oil, 35.1 vol. percent coker heating oil and 36.2 vol. percent catalytic cycle oil.

Two runs were performed to demonstrate the desirability of hydrofining the feed to an intermediate nitrogen content level, as opposed to an essentially nitrogen-free feed. In Run A the feed was hydrofined to a level of 1 p.p.m. nitrogen, and in Run B it was hydrofined to a level of 28 p.p.m. nitrogen. The results were compared to those obtained by hydrocracking the raw feed, designated as Run C.

The feed properties, operating conditions of the hydrofining and hydrocracking stages, yields and selectivity data for the three runs are summarized in the following table:

TABLE II Raw Hydrofined Feed Hydroeraeker Feed Run Number C B A Feed Properties:

Total Nitrogen p.p.m 758 28 1.0

Analine Point, F 121. 136 143 Pour Point, F. 35. 0 45. 0 50. 0

Gravity, API.. 27. 1 31. 9 32. 9

Sulfur, wt. percen 0. 49 0. 04 0. 04

R1. at 67 C l. 4881 1. 4637 1. 4594 Distillation Range:

I.B.P., F 366 341 328 5% at F 409 391 at F 428 409 50% at F 556 538 F.B.P., 774 780 Hydrofining Conditions:

Catalyst COM'OOJ/AlgOa Temperature, F 700 Space Velocity, v. 0.5-0.7

Pressure, p.s.i.g 1,500

Hydrogen Rate, s.c.f./b 10,000 Hydrocracking Conditions:

Catalyst Pd on crystalline zeolite Temperature, F 798 645 618 Space Velocity, v./v./hr 1. 09 1. 01 0. 94

Pressure, p.s.i.g, 1,340 1, 485 1, 500

H Rate, set/b 6, 630 8, 150 13, 800

Conversion to 430 F.-, v01. percent 38. 8 60 62. 8

l 0 TABLE II.-Continued Raw Hydrocracker Feed Run Number C B A Yields on Raw Feed:

(1 wt. percent 0 vol. pereent 0 vol. percent 5 4. (I /430 F., v01. percent 39. Selectivity- 1 A blended feed of two-thirds coker gas oil and onethird steam cracked cycle oil, containing 812 p.p.m. nitrogen, was hydrofined and hydrocracked in substantial accordance with the procedure of Example 1. Hydrofining was accomplished with the cobalt molybdate on alumina catalyst of Example 1, at two severity levels. Blend D was hydrofined to a level of 158 p.p.m. nitrogen and then hydrocracked with a 2% palladium hydrocracking catalyst prepared by the procedure of Example 1, except that a more concentrated palladium solution was used in the final metal deposition step. Blend E was severely hydrofined to a 10 p.p.m. nitrogen level and hydrocracked with the 0.5% palladium catalyst used in Example 1.

The feed properties of the initial blend and hydrofined blends, the operating conditions of hydrofining and hydrocracking, and the yields and product quality data resulting from the integrated hydrofining-hydrocracking operation are shown in the following table:

TABLE III Hydrofined Raw Hydrocracker Feed Feed Mild Severe Feed Properties:

Total Nitrogen, p.p.m 158 10 Gravity, API 27. 8 29. 6 Aniline Pt., F 124. 0 135. 5 Sulfur, Wt. percent 0.05 0. 00 R1. at 67 C 1. 4781 1. 4697 Distillation Range:

.B.P., 391 372 366 10% at F 465 453 446 50% at F. 544 538 532 95% at F 809 780 780 llydroiining C Catal COLIOO4/A1ZO3 Temperature, F"... 700 700 Spice Velocity, v./v./l1r 1.0 0.5 Pressure, p.s.i.g 1, 500 2,300

Hz Rate, s.e.i./b Hydrocracking Conditions:

Catalyst a Pd on crystalline zeoli 03-, wt. percent 1.8 0. 9

Ci. Vol. percent 7. 6 6.7

Q5/4130 F., V01. percent 69. 1 70. 4 05/430 F. Quality:

API 49.0 49. 7

RON+3 cc.. a. 88 82.5

MON+3 cc 0 83 Aromatics, Vol. percent 22 8 I On 65430 F. fraction.

As shown in the above table, the octane number of the C /430" F. gasoline, fraction obtained from the mildly hydrofined feed was significantly greater than for the severely hydrofined blend. A distinct product quality advantage is therefore offered by limiting the severity of the hydrofining treatment to produce an intermediate nitrogen content range. I

What is claimed is:

1. A process for hydrocracking a hydrocarbon oil feed having a nitrogen content within the range of about 28 to about 200 parts per million comprising contacting said feed in the presence of hydrogen with a hydrocracking catalyst comprising a crystalline alumino-silicate zeolite composited with a platinum group metal, said zeolite having pore openings of about 6 to about 15 Angstrom units and containing less than about 10 wt. percent alkali metal oxide, at hydrocracking conditions sufiicient to obtain the desired conversion to lower boiling product.

2. The process of claim 1, wherein said zeolite is composited with 0.01 to 5.0 wt. percent platinum group metal.

3. The process of claim 1, wherein said platinum group metal is palladium.

4. The process of claim 1, wherein said alkali metal oxide is Na O.

5. The process of claim 1, wherein said hydrocracking conditions include a temperature of from about 450 to about 800 F., a pressure of from about 500 to about 3000 p.s.i.g., a liquid hourly space velocity of from about 0.1 to about 10 volumes of feed per volume of catalyst per hour, and a hydrogen rate of from about 1000 to about 20,000 standard cubic feet per barrel of feed.

6. A process for hydrocracking a hydrocarbon oil feed having a nitrogen content within the range of about 28 to about 200 parts per million said feed substantially boiling above about 400 F. to obtain lower boiling product which comprises contacting said feed in the presence of hydrogen with a hydrocracking catalyst comprising a crystalline alumino-silicate zeolite composited with 0.01 to 5.0 wt. percent platinum group metal, said zeolite having pore openings of about 6 to about Angstrom units and containing less than about 10 wt. percent Na O, at a temperature of from about 450 F. to about 800 F., a pressure of from about 500 to about 3000 p.s.i.g., a liquid hourly space velocity of from about 0.1 to about 10 volumes of feed per volume of catalyst per hour, and a hydrogen rate of from about 1000 to 12 about 20,000 standard cubic feet per barrel of feed, and recovering lower boiling product.

7. The process of claim 6, wherein said zeolite is com posited with 0.1 to 3.0 wt. percent palladium.

8. The process of claim 6, wherein said hydrocracker feed has a nitrogen content within the range of about 28 to about 100 p.p.m. nitrogen.

9. The process of claim 6, wherein said hydrocracker feed has a nitrogen content within the range of about 28 to about p.p.m. nitrogen.

10. A process for hydrocracking a hydrocarbon oil feed substantially boiling above about 400 F. and having a nitrogen content in the range of about 28 to about 50 p.p.m. nitrogen to obtain lower boiling product, which process comprises contacting said feed in a hydrocracking zone, in the presence of hydrogen, with a hydrocracking catalyst comprising a crystalline alumino-silicate zeolite composited with a platinum group metal, said zeolite having pore openings of about 6 to about 15 Angstrom units and containing less than about 10 wt. percent Na O, at a temperature of from about 600 to about 750 F., a pressure of from about 1200 to about 1800 p.s.i.g., a liquid hourly space velocity of from about 0.5 to about 3 volumes of feed per volume of catalyst per hour, and a hydrogen rate of from about 2000 to about 12,000 standard cubic feet per barrel of feed, and recovering said lower boiling product from the efiluent of said hydrocracking zone.

11. The process of claim 10, wherein said zeolite is composited with 0.01 to 5.0 wt. percent platinum group metal.

12. The process of claim 10, wherein said zeolite is composited with 0.1 to 3.0 wt. percent palladium.

13. The process of claim 10 which further comprises separation of hydrocarbon fractions boiling higher than desired lower boiling product and recycle of said fractions to said hydrocracking zone.

References Cited by the Examiner UNITED STATES PATENTS 2,971,904 2/ 1961 Gladrow et al 20846 3,132,089 5/1964 Hass et al 208254 3,132,090 5/1964 Helfrey et al. 20889 3,159,568 12/1964 Price et al. 20889 DELBERT E. GANTZ, Primary Examiner.

S. P. JONES, Assistant Examiner.

Claims

1. A PROCESS FOR HYDROCRACKING A HYDROCARBON OIL FEED HAVING A NITROGEN CONTENT WITHIN THE RANGE OF ABOUT 28 TO ABOUT 200 PARTS PER MILLION COMPRISING CONTACTING SAID FEED IN THE PRESENCE OF HYDROGEN WITH A HYDROCRACKING CATALYST COMPRISING A CRYSTALLINE ALUMINO-SILICATE ZEOLITE COMPOSITED WITH A PLATINUM GROUP METAL, SAID ZEOLITE HAVING PORE OPENINGS OF ABOUT 6 TO ABOUT 15 ANGSTROM UNITS AND CONTAINING LESS THAN ABOUT 10 WT. PERCENT ALKALI METAL OXIDE, AT HYDROCRACKING CONDITIONS SUFFICIENT TO OBTAIN THE DESIRED CONVERSION TO LOWER BOILING PRODUCT.