US3458299A - Hydrocracking process - Google Patents

Description

United States Patent 3,458,299 HYDROCRACKING PROCESS Robert H. Hass, Fullerton, Califl, assignor to Union Oil Company of California, Los Angeles, Calif., a corporation of California No Drawing. Filed June 23, 1964, Ser. No. 377,372 Int. Cl. Cg 35/06 US. Cl. 48-213 14 Claims ABSTRACT OF THE DISCLOSURE A process for manufacturing propane-butane mixtures by passing naphtha and/or light gas oil through a bed of crystalline zeolite hydrocracking catalyst in the presence of hydrogen and correlating pressure, temperature and space velocity to yield at least a 40 volume percent yield of total propane and butane per pass, based on feed.

This invention relates to the catalytic hydrocracking of hydrocarbons, and more specifically is directed to the objective of converting naphtha and/or light gas oil fractions to propane-butane mixtures, conventionally referred to in the art as liquefied petroleum gas (LPG).

In broad aspect the process consists in passing the vaporized feed in admixture with hydrogen through a bed of hydrocracking catalyst comprising a crystalline, siliceous zeolite cracking base upon which is supported a Group VIII metal hydrogenat-ing component, at relatively low temperatures between about 500 and 800 F., under conditions of pressure and space velocity such that a substantial per-pass yield of LPG is produced, amounting to at least about 40% by volume of the feed. The most critical aspects of the invention reside in the use of crystallize zeolite catalysts which provide a large number of acid cracking centers (ion-exchange sites) per unit of surface area, and in operating at low temperatures which minimize the production of methane-ethane (dry gas). The provision of conditions such that the feed is totally in the vapor phase in at least a major portion of the reaction zone is also critical to obtain maximum conversion rates to LPG with resultant increase in permissible space velocities for a given conversion per pass. The latter is important in view of the low temperatures required in order to minimize dry gas production.

The basic objective of the invention is to convert naphthas and/or light gas oils selectively to LPG, with a minimum production of coke and dry gas, while also operating at economical high space velocities, low pressures, and low catalyst deactivation rates with resultant long run lengths between regenerations. Other objects will be apparent from the more detailed description which follows.

The modern development of catalytic hydrocracking has been directed almost exclusively toward the objective of converting light and heavy gas oils and residua to gasoline, jet fuel and diesel fuel products. In these processes, a primary objective has been to minimize the production of hydrocarbons in the C C range. This however is dictated only by the prevailing market demands in most areas of the United States. But in some local areas there is actually a surplus of products boiling in the gasoline and naphtha ranges, and a relative shortage of LPG. The same is true in many foreign countries. To satisfy the needs of these localized areas, it would be highly desirable to provide a simple and economical means for converting surplus naphtha fractions to LPG, if this can be achieved without substantial conversion to dry gas (which merely competes with natural gas). The process of this invention is well adapted toward the satisfying of these specialized requirements.

To achieve the results herein desired, namely selective conversion to LPG in preference to other products such as gasoline, dry gas and coke, the critical process variables center around the nature of the feedstock and the hydrocracking temperature. For the preferred operations carried out at economical high space velocities with minimal catalyst deactivation rates, the principal determining factors reside in the nature of the catalyst, the hydrogen partial pressure, and the presence or absence of a liquid phase in the hydrocracking zone, the latter being a function of feed boiling range (principally the end-boilingpoint), hydrogen-to-oil ratio, pressure and temperature employed. The criticality of these variables will be more apparent from the following description.

Feedstocks Feedstocks which may be employed herein may comprise any mineral oil fraction boiling between about 120 and 650 F., which range includes light gasolines, heavy gasolines, light gas oils, kerosene, or any desired mixtures or fractions thereof. The preferred feedstocks are those which boil below about 500 F., and still more preferably, between about 200 F. and 400 F. These feedstocks may be free of impurities, or they may contain substantial amounts of organic sulfur and/ or nitrogen compounds. A particularly desirable feature of the process resides in the fact that the zeolite catalysts employed are substantially more tolerant of sulfur and nitrogen compounds than are conventional amorphous hydrocracking catalysts such as those base on silica-alumina cogel. Feedstocks containing up to about 3 weight-percent of sulfur, and up to about 400 parts per million of total nitrogen may be utilized without prior hydrofining; for feedstocks containing higher proportions of sulfur and/or nitrogen, a hydrofining pretreatment is normally desirable.

A particularly desirable feature of the process which renders it more economical than other known processes resides in the fact that feedstocks containing between about 5 and parts per million of total nitrogen may be converted at the low temperatures required without encountering excessive catalyst deactivation rates. Feedstocks in this range of purity are often available as virgin stocks. In other cases, they may purposely be obtained from an incomplete hydrofining pretreatment, this being much more economical than a complete hydrofining to reduce nitrogen contents to below about 1 part per million. These hydrofining pretreatments may be carried out as a separate step with intervening removal of ammonia and hydrogen sulfide, or they may be carried out integrally with the hydrocracking step in which case the ammonia and hydrogen sulfide formed during hydrofining is passed through the hydrocracking zone. The zeolite catalysts of this invention are unique in their ability to maintain activity in the presence of substantial amounts of ammonia.

The end-boiling-point of the feedstock is important herein for two principal reasons. Firstly, high-end-point feedstocks will inherently result in the production of higher proportions of dry gas due to the more extended thermal cracking which takes place during conversion of the heavy ends down to the LPG range. Secondly, it is more difficult to maintain the high-end-point feedstocks entirely in the vapor phase during the hydrocracking. Where a liquid phase is present, there is a tendency to decrease the rate of formation of C -C hydrocarbons, as compared to the rate at which hydrocarbons in the C C range are produced. This is apparently because, when the surface area of the catalyst is coated with a liquid phase, the catalyst functions primarily to crack the heavy liquid hydrocarbons in the C -C range down to hydrocarbons mostly in the C C range. The rate of formation of LPG hydrocarbons by the hydrocracking of C -C hydrocarbons is relatively reduced because of the more limited access of the C -C hydrocarbons (which are mostly in the vapor phase) to the active cracking centers, the latter being coated with a liquid phase diffusion barrier. It is therefore desirable to operate in the vapor phase in order to provide maximum cracking rates for the C C hydra carbons, from which LPG hydrocarbons are primarily produced. It is therefore preferred to employ feedstocks having an end-boiling-point below about 5000 F., since these feedstocks are more easily maintained in total vapor phase under the optimum hydrocracking conditions described herein.

Hydrocracking conditions Hydrocracking conditions contemplated herein fall in general within the following ranges:

Of the foregoing variables, the most important is probably temperature. At temperatures above about 800 F., the rate of thermal cracking becomes very significant, with resultant sharp increase in the production of dry gas. Also the rate of coke formation increases, with resultant rapid catalyst deactivation.

Optimum temperatures also vary depending upon the end-boilingpoint of the feedstock. Relatively high temperatures are generally optimum for low-boiling feeds, while lower temperatures are generally preferred for feedstocks in the higher boiling ranges. Assuming the normal distribution of hydrocarbons found in most mineral oil fractions, wherein at least about 3 volume-percent thereof boils within 50 F. of the end-boiling-point, optimum temperatures are in general defined by the following equation, wherein T represents hydrocracking temperature and E represents the end-boiling-point of the feed, both in degrees F.:

It will be understood that temperatures are adjusted within the prescribed ranges to obtain the desired conversion per pass to LPG, preferably within the range of about 40-100 volume-percent per pass, based on feed.

A very desirable feature of the process resides in the relatively high space velocities which may be utilized at the low temperatures specified. Conventional hydrocracking catalysts based upon amorphous cracking composites such as silica-alumina cogel generally require space velocities below about 0.5 in order to achieve the desired conversions at the low temperatures utilized herein. The high space velocities useable herein permit a marked decrease in reactor size and catalyst volume required for obtaining a given conversion and throughput at a given efficiency level. Efficiency is herein defined as pounds of LPG produced per pound of dry gas produced. Efficiency levels above about 20, and ranging up to about 300 are readily obtainable herein, although levels above about 100 are usually not obtainable unless the end-boiling-point of the feedstock is below about 450 F.

An important effect of hydrogen partial pressure herein lies in its determining effect upon catalyst activity maintenance, and thus run lengths between regenerations. For obtaining run lengths of at least about 4 months, and up to 1-2 years, it is generally preferred to employ pressures above about 1,000 p.s.i.g. This however depends to some extent on the feedstock employed, high-end-point feeds requiring higher pressures to obtain a given run length. Where the feed is in the vapor phase, hydrogen partial pressure also has a marked effect on reaction rates, more so than when a liquid phase diffusion barrier is present. For this reason also, pressures above 1,000 p.s.i. g. are preferred, especially at low hydrocracking temperatures.

With temperature and pressure determined by the above criteria, the principal remaining variable, hydrogen/oil ratio, is preferably adjusted so as to insure that the feedstock is substantially completely in the vapor phase in at least most of the contacting zone. Preferably, sufficient hydrogen is used to provide a total vapor phase operation.

The hydrocracking conditions prescribed above seldom provide for conversion per pass to LPG and dry gas. To achieve such a result in a single pass would result in marked lowering of the efiiciency. For total conversion of the feed to LPG and dry gas, it is preferable to operate at the conversion levels per pass as above-prescribed and to recycle to the reactor, or to a second reactor, any unconverted material boiling above the product LPG.

Hydrocracking catalysts As suggested above, the most critical aspect of the catalysts employed herein resides in the nature of the cracking base upon which the hydrogenating metal is distended. These crystalline siliceous zeolites are sometimes referred to in the art as molecular sieves, and are composed usually of silica, alumina and one or more exchangeable cations such as sodium, hydrogen, magnesium, calcium, etc. They are further characterized by crystal pores of relatively uniform diameter between about 4 and 14 A. It is preferred to employ molecular sieve zeolites having a relatively high SiO /Al O mole-ratio, between about 3.0 and 12, preferably between about 4 and 8 Suitable zeolites found in nature include for example mordenite, stilbite, heulandite, ferrierite, dachiardite, chabazite, erionite, and faujasite. Suitable synthetic molecular sieve zeolites include for example those of the B, X, Y and L crystal types, or synthetic forms of the natural zeolites noted above, especially synthetic mordenite. The preferred zeolites are those having crystal pore diameters between about 8-12 A., wherein the SiO /Al O moleratio is about 3-6. A prime example of a zeolite falling in this preferred group is then synthetic Y molecular sieve.

The naturally occurring molecular sieve zeolites are normally found in a sodium form, an alkaline earth metal form, or mixed forms. The synthetic molecular sieves normally are prepared first in the sodium form. In any case, for use as a cracking base it is preferred that most or all of the original zeolitic monovalent metals be ionexchanged out with a divalent metal, or with an ammonium salt followed by heating to decompose the zeolitic ammonium ions, leaving in their place hydrogen ions and/or exchange sites which have actually been decationized by further removal of water:

(NHtl)X +)XZ +I s A (H+ xz z H2o Mixed divalent metal-hydrogen zeolites may be prepared ion-exchanging first with an ammonium salt, then partially back-exchanging with a divalent metal salt, and then calcining. In some cases, as in the case of synthetic mordenite, the hydrogen forms can be prepared by direct acid treatment of the alkali metal sieves. Hydrogen or decationized Y sieve zeolites of this nature are more particularly described in Belgian Patents Nos. 598,582, 598,682, 598,683 and 598,686, and US. Patent No. 3,130,- 006.

There is some uncertainty as to whether the heating of the ammonium zeolites produces a hydrogen zeolite or a truly decationized zeolite, but it is clear that, (a) hydrogen zeolites are formed upon initial thermal decomposition of the ammonium zeolite, and (b) if true decationization does occur upon further heating of the hydrogen zeolites, the decationized zeolites also possess desirable catalytic activity. Both of these forms, and the mixed forms, are designated herein as being metal-cation-deficient.

The essential active metals employed herein as hydrogenation components are those of Group VIH, i.e., iron, cobalt, nickel, ruthenium, rhodium, palladium,- osmium, iridium and platinum, or mixtures thereof. The noble metals are preferred, and particularly palladium and platinum. In addition to these metals, other promoters may also be employed in conjunction therewith, including the metals of Groups VI-B and VII-B, e.g., molybdenum, chromium, manganese, etc.

The amount of hydrogenating metals in the catalyst can vary within wide ranges. Broadly speaking, any amount between about 0.1% and 20% by weight may be used. In the case of the noble metals, it is normally preferred to use about 0.2% to 2% by weight. The preferred method of adding the hydrogenating metal is by ion exchange. This is accomplished by digesting the zeolite, preferably in its ammonium form, with an aqueous solution of a suitable compound of the desired metal wherein the metal is present in a cationic form, as described for example in Belgian Patent No. 598,686.

Following addition of the hydrogenating metal, the resulting catalyst powder is then filtered olf, dried, pelleted with added lubricants, binders, or the like if desired, calcined at temperatures of, e.g., 700-1,200 F. in order to activate the catalyst and decompose zeolitic ammonium ions.

The followingexamples are cited to illustrate suitable techniques and results obtainable in the practice of the invention, but are not to be construed as limiting in scope.

Example I Six hydrocracking runs were carried out under various hydrocracking conditions, using as feedstock a heavy Kuwait naphtha having the following characteristics:

TABLE 2 Gravity API 52.4 Sulfur wt. percent-.. 0,216 Nitrogen p.p.m 9 Boiling range F 217-500 Aromatics vol. percenL- 13.2

The catalyst employed throughout was a copelleted mixture of about 80 weight-percent of a Y molecular sieve zeolite containing 0.5 weight-percent palladium, and an activated alumina binder containing 0.3 weight-percent palladium. The Y molecular sieve cracking base had a SiO /Al O mole-ratio of about 4.7, about 35% of the zeolitic ion-exchange capacity being satisfied by magnesium ions (3 weight-percent MgO), about 10% by sodium ions, and the remainder by hydrogen ions. This zeolite base was prepared by first subjecting a sodium Y molecular sieve to ion exchange with an aqueous ammonium salt solution, and then partially back-exchanging the resulting ammonium zeolite with magnesium sulfate solution. Hydrocracking conditions employed in the various runs, and the respective results were as follows:

were of excellent blending quality, having iso/normal parafiin ratios ranging between about 2.57 and 7.05, and having octane numbers (F1+3 n11. TEL) ranging between about 97.0 and 98.3.

Example H The results obtained in Example I indicate that the zeolite catalyst employed retains its activity for substantial periods of time in the presence of the nitrogen-containing feedstock employed. To demonstrate that this sustained hydrocracking activity is not obtainable using conventional hydrocracking catalysts, a catalyst consisting of 20% nickel impregnated on a coprecipitated gel-type cracking base (90% silica-10% alumina) was employed for hydrocracking a 580 F. end-point gas oil initially containing only about 2 parts per million of nitrogen. At 0.8 LHSV, 1,500 p.s.i.g., and 8,000 s.c.f./b. of hydrogen, this catalyst gave 60 volume-percent conversion of the feed to gasoline at a hydrocracking temperature of 585 F. Upon adding 100 parts per million of nitrogen to the feed (as t-butylamine), the conversion began to drop rapidly and at the end of 34 hours had declined to zero. Upon raising the hydrocracking temperature to 707 F., the conversion level was still only 26 volume percent. It is thus evident that, at the hydrocracking temperatures here concerned, the amorphous catalysts are much less active in the presence of nitrogen compounds than the zeolite catalysts.

Substantially similar results are obtained when other hydrocracking catalysts and feedstocks within the purview of this invention are substituted in the foregoing Example I. It is therefore not intended that the invention should be limited to the details of the examples but broadly as defined in the following claims.

I claim:

1. A process for the manufacture of propane-butane mixtures which comprises subjecting a mineral oil fraction boiling below about 650 F. to catalytic hydrocracking at a temperature between about 500 and 800 F. and at pressures about 400 p.s.i.g. in the presence of added hydrogen and a hydrocracking catalyst comprising a minor proportion of a Group VIII metal hydrogenation component deposited upon a crystalline, siliceous zeolite cracking base wherein a substantial proportion of the zeolitic ion-exchange capacity is satisfied by hydrogen ions and/or polyvalent metal ions, and correlating the hydrocracking conditions of temperature, pressure and space velocity so as to give at least about 40 volume-percent yield of total propane and butane per pass, based on feed.

2. A process as defined in claim 1 wherein the hydrocracking conditions of temperature, pressure and hydrogen/oil ratios are adjusted and correlated so as to provide complete vaporization of said mineral oil fraction in at least a major portion of the hydrocracking zone.

3. A process as defined in claim 1 wherein said Group VIII metal is a noble metal.

TABLE 3 Catalyst age, hrs 11-19 59-67 79-87 103-111 127-135 150-158 Pressure, p.s.i.g 500 500 500 500 500 1, 000 Temperature, F 701 675 700 700 725 651 LHSV,v./v./hr 1.0 1 0 1.0 2.0 2.0 1.0 Hz/Oil ratio, 1n.s.c.f./b 4 4 4 4 4 4 Product yields:

Dry gas, s.c.f./b., feed".--

84.2 52. 7 72.9 38. 6 65.4 33. 4 LPG, vol. percent feed 96.4 76. 3 79. 8 49. 3 63. 7 57. 2 05-09, vol. percent feed 37. 8 45. 1 40.1 43. 7 41. 9 42. 0 0 vol. percent feed l. 5 8. 3 12. 2 37. 0 29. 3 23. 8 Efficieney, LPG/dry gas, lbs/lb 29. 7 30. 9 27. 8 32. 4 24. 7 44. 7

is readily pp the above 'f that 4. A process as defined in claim 1 wherein said zeolite tifliclencyfif P Rroductlon'bear s a Substantlauy g cracking base is an alumino-silicate molecular sieve of line relationship wlth hydrocracklng temperature, 1g est the Y crystal type, having a Sick/A1203 m01e ratio efiiciency being obtained in low-temperature run 6, and lowest efficiency in high-temperature run 5. In all of the foregoing runs, the conditions were such that the feed was substantially wholly in the vapor phase in the hydrotween about 3 and 6.

5. A process as defined in claim 1 wherein said hydrocracking is carried out at a temperature between about cracking zone. The C -C gasoline fractions obtained 600 and 750 F.

6. A process as defined in claim 1 wherein said hydrocracking is carried out at a temperature defined by the equation:

T=800-0.3Ei40 wherein T is temperature in F. and E is the end-boilingpoint of said mineral oil fraction in F.

7. A process as defined in claim 1 wherein said mineral oil fraction contains between about and 400 p.p.m. of nitrogen.

8. A process as defined in claim 1 wherein product oil having a boiling point above that of butane is recycled to give an overall conversion to butane and lighter products of substantially 100 percent.

9. A process for the manufacture of propane-butane mixtures which comprises subjecting a mineral oil fraction boiling between about 120 and 500 F. to catalytic hydrocracking in the vapor phase at pressures above 400 p.s.i.g. and at a temperature between about 600 and 750 F. and a space velocity between about 1 and 5 in the presence of added hydrogen and a hydrocracking catalyst comprising a minor proportion of a Group VIII metal hydrogenation component deposited upon a crystalline, alumino-silicate molecular sieve zeolite cracking base having a SiO /Al O mole-ratio between about 3 and 12, wherein a substantal proportion of the zeolitic ion-exchange capacity is satisfied by hydrogen ions and/or polyvalent metal ions, and maintaining an adjusted combination of hydrocracking temperature, pressure and space velocity within the above ranges so as to give at least about a 40 volume-percent yield of total propane and butane per pass, based on feed.

10. A process as defined in claim 9 wherein said Group VIII metal is a noble metal.

11. A process as defined in claim 9 wherein said zeolite cracking base is an alumino-silicate molecular sieve of the Y crystal type, having a SiO /Al O mole-ratio between about 3 and 6.

12. A process as defined in claim 9 wherein said hydrocracking is carried out at a temperature defined by the equation:

wherein T is temperature in F. and E is the end-boilingpoint of said mineral oil fraction in F.

13. A process as defined in claim 9 wherein said mineral oil fraction contains between about 5 and 400 p.p.m. of nitrogen.

14. A process as defined in claim 9 wherein product oil having a boiling point above that of butane is recycled to give an overall conversion to butane and lighter products of substantially percent.

References Cited UNITED STATES PATENTS 2,560,433 7/1951 Gilbert et al. 252-459 XR 2,830,880 4/ 1958 Shapleigh et al. 48-214 XR 2,906,700 9/1958 Stine et al. 252-459 XR 3,069,351 12/ 1962 Davis 208-214 XR 3,077,448 2/ 1963 Kardash et al. 208-217 XR 3,144,414 8/1964 Silverman 252-459 XR 3,329,628 7/1967 Gladrow et al. 252--459 XR 3,325,316 6/1967 Mulaskey 252-459 XR 3,325,465 6/1967 Jones 252-459 XR FOREIGN PATENTS 930,512 7/ 1963 Great Britain.

MORRIS 0. WOLK, Primary Examiner R. E. SERWIN, Assistant Examiner US. Cl. X.R. 208-137; 252-459