US3468788A - Hydrocracking process - Google Patents

Description

YUnited States Patent O 3,468,788 HY DROCRACKING PROCESS Herbert F. Wilkinson, Fullerton, Calif., assignor to Union Gil Company of California, Los Angeles, Calif., a

corporation of California Filed Aug. 30, 1966, Ser. No. 576,086

Int. Cl. Cg 23/02, 23/00 U.S. Cl. 208-89 12 Claims This invention relates to new methods for the catalytic hydrocracking of high-boiling hydrocarbons to produce therefrom lower-boiling hydrocarbons, and is particularly directed to a flexible hydrocracking process which can be operated for 100% conversion to high octane gasoline, and at other times to produce a minimum of gasoline-boilingrange products, and a maximum of high-quality middle distillate product including for example jet fuel, diesel oils, fuel oils and the like. In broad aspect the invention embraces a series combination of hydrofining followed by hydrocracking. It has been found that in such systems it is preferable to pass total effluent from the hydrofner (including impurities such as ammonia which were formed in the hydrofner) through the hydrocracker at relatively high temperatures when maximum production of highoctane gasoline is desired, but that when maximum production of middle di-stillate is desired, it is preferable to by-pass the hydrofner effluent around the hydrocracker, blending it with effluent therefrom, and separate from the resulting blend the total desired middle distillate product, the hydrocracker being used only for converting recycle oil boiling above the end boiling point of the middle distillate product.

In the gasoline-producing cycle of the process the hydrofining-hydrocracking sequence is integral, in that no separation or other purification of the hydrofner effluent is required prior to hydrocracking. Conversely, in the middle distillate producing cycle of the process, the hydroiining-hydrocracking sequence is non-integral; the effluent from the hydrofner is cooled, water washed, and fractionated before any of it reaches the hydrocracker. As is well known, ammonia and other basic nitrogen compounds have a strong temporary poisoning effect upon hydrocracking catalysts, an effect which can be counteracted to a substantial degree by raising the hydrocracking temperature. Since ammonia is present in the feed to the hydrocracker during the integral gasoline-producing cycle of the process, higher temperatures are employed in the hydrocracker resulting in a relatively high octane, aromatic gasoline product. However, in the non-integral cycle of the process ammonia is substantially absent from the feed to the hydrocracker, and hence substantially lower temperatures can be employed therein, resulting fortuitously in ythe production of a more saturated, paraffinic and naphthenic middle distillate product. Thus, there is a two-fold advantage in the non-integral system for middle distillate production; due to the exclusion of middle distillate from the feed to the hydrocracker there is a minimal conversion of middle distillate to gasoline, with resultant minimum gasoline yields, and due to the low temperatures employed in the hydrocracker the resulting middle distillate product is a source of higher quality jet fuels, diesel fuels, burner oils and the like.

The principal objective of the invention is to provide flexibility of product in hydrocracking processes. In many areas there is a substantial seasonal fluctuation in demand for gasoline and or middle distillate products such as diesel fuels, burner oils, jet fuels and the like. Typically, the demand for gasoline may be at a maximum during the summer months while in winter months gasoline demands may decline substantially, accompanied by a marked increase in demand for diesel oils and fuel oils. It is well ICC known that all of these desired products can be produced by catalytic hydrocracking, but it is not a simple matter to design a single unit possessing the flexibility required to permit rapid and convenient alternation between an -100% gasoline product and a 70-90% middledistillate product, while still maintaining maximum quality of both products. To a substantial extent, the process of this invention achieves all of these objectives, and others which will be apparent from the more detailed description which follows.

For the purposes of this invention, middle distillate is defined as mineral oil distillates boiling within the range of about S50-800 F. This includes primarily jet fuels boiling in the S50-650 F. range, diesel fuels boiling in the 400-800 F. range, and stove or burner oils boiling in the 40G-600 F. range. From the quality standpoint, jet fuels should contain a minimum or normal paraflins in order to meet freezing point specifications, and a minimum of aromatic hydrocarbons in order to producea smokeless, relatively non-luminous flame. Diesel fuels require primarily a low aromatic content in order to meet cetane number specifications. Burner oils and stove oils may be somewhat more aromatic in character, but the normal paraffin and aromatic content must be sufficiently low to meet freezing point specifications. Most of these desired product qualities conflict with the major desired quality in gasolines, namely a high aromatic content. The process of this invention affords not only great flexibility in respect to product boiling range, but also substantial flexibility in respect to product aromaticity, coinciding with the quality desired in products of each respective boiling range.

FEEDSTOCKS The feedstocks which may be treated herein include in general any mineral oil fraction boiling above the conventional gasoline range, i.e., above about 300 F., and usually above about 400 F., and having an end boiling point up to about l,l00 F. This includes straight run gas oils and heavy naphthas, coker distillate gas oils and heavy naphthas, crude oils, reduced crude oils, cycle oils derived from catalytic or thermal -cracking operations and the like. These fractions may be derived from petroleum crude oils, shale oils, tar sand oils, coal hydrogenation products and the like. Specifically, it is preferred to employ feedstocks boiling between about 400 and 900 F., having an API gravity of about Ztl-35, and containing at least about 20% by volume of acid-soluble components (aromatics plus olefins). Such oils may also contain up to about 5% by weight of sulfur and up to about 2% of nitrogen. When the principle desired product is a jet fuel, it is further preferred that the feed contain' not more than about 30 Volume percent of material boiling below the end-boiling-point of the desired jet fuel product. This is to insure that the jet fuel product will contain a sufficiently high ratio of iso-paraflns/normal paraftins to mee-t freezing point specifications; jet fuel hydrocarbons synthesized by hydrocracking are richer in iso-paraffins than the virgin feed fraction boiling in the desired product range.

DETAILED PROCESS DESCRIPTION Reference is now made to the attached drawing which is a flow sheet illustrating a preferred mode of practicing the invention. Operation of the process will be described first in connection with the total gasoline production cycle. The raw feed is brought in via line 2, and passed through heat exchanger 4 where it is preheated -to e.g., about 400 to 650 F., by exchange against hot reactor effluent. The resulting preheated feedstock in line 6 is then passed into the top of catalytic hydrofner 8, after being mixed with hot make-up and recycle hydrogen injected via lines 10 and 12 and preheater 14. This hot hydrogen completes the preheating of feed to the desired incipient hydrofining temperatures. In hydrofiner 8, hydroning proceeds under -substantially conventional conditions. Suitable hydroning catalysts include for example mixtures of the oxides and/ or sultides of cobalt and molybdenum, or of nickel and tungsten, preferably supported on a carrier such as alumina, or alumina containing a small amount of coprecipitated silica gel. Other suitable catalysts include in general the oxides and/ or suldes of the Group .VI-B and/or Group VIII metals, preferably supported on relatively noncracking adsorbent oxide carriers such as alumina, silica, titania, and the like, having a Cat-A cracking Activity Index below about 25. The hydrofining `operation may be conducted either adiabatically or isothermally, and under the following conditions:

HYD ROFINING CONDITIONS The above conditions are suitably adjusted so as to reduce the nitrogen content of the feed to below about 25 parts per million, and preferably below `about l parts per million.

Hot eiuent from hydroner 8 is withdrawn via line 16 and transferred via line 18 (valve 20 open; valve 22 closed) into the top of catalytic hydrocracker 19, along with recycle oil from line 58 to be described hereinafter. Any desired temperature adjustment of the hydrocracker feed is obtained by means of a second portion of recycle hydrogen injected into line 18 via line 24 and preheater 26.

In hydrocracker 19 the conversion to gasoline boiling range hydrocarbons is carried out under the following general conditions:

HYDROCRACKING CONDITIONS FOR GASOLINE The conditions of temperature and space velocity are adjusted and correlated within the above ranges so `as to achieve about 30-80 volume-percent conversion to 400 F. end-point gasoline per pass. Since the reaction is highly ex-othermic, it is normally desirable to inject cool quench hydrogen by means not shown at one or more levels in the reactor. To achieve hydrocracking under the above conditions, and in the presence of ammonia generated in the hydroner, it is necessary to employ highly active hydrocracking catalysts, to be described hereinafter.

Hot efuent from hydrocracker 19 is withdrawn via line 28 and passed through heat exchanger 4 as previously described to effect partial cooling thereof to temperatures of e.g., 20G-300 F. The partially cooled effluent in line 30 is then mixed with wash water injected via line 32, and passed through secondary product cooler 34 where nal cooling to temperatures of e.g., Sil-150 F. is effected. The cool mixture then ows into high-pressure separator 36, from which puried recycle gas is taken overhead via line 38, and spent wash water containing dissolved ammonia and/ or ammonium suldes is withdrawn via line 40. The high-pressure liquid condensate in separator 36 is then flashed via line 42 into low-pressure separator 44, from which ash gases comprising mainly light hydrocarbons are exhausted via line 46.

The low-pressure condensate in separator 44 is transfel-red via line 48 to fractionating column 50, from which light C4-C6 gasoline is withdrawn overhead via line 52, and a heavy C7+ gasoline product via sidecut line 54. The

light gasoline withdrawn via line 52 is normally rich in iso-paraffns and constitutes excellent gasoline blending stock. The heavy gasoline recovered via line 54 is relatively rich in aromatics, and may also vbe blended directly into finished gasolines in many cases, but in other instances may desirably be subjected to catalytic reforming. The bottoms from column 50 is withdrawn via line 56 and recycled via line 58 (valve 62 open; valves 69 and 64 closed) to hydrocracker feed line 18 as previously described in order to achieve total conversion to gasoline. Normally, during the gasoline producing cycle, fractionation column 68 is not used.

MIDDLE DISTILLATE CYCLE To convert the system described above from gasoline to maximum middle distillate production, valve 20 is closed, and valve 22 opened, thereby diverting the hydroner effluent in line 16 directly through heat exchanger 4 in admixture with the hydrocracker efiuent from line 28. The combined effluents are then water washed, separated and depressured as previously described, with total low-pressure condensate from both reactors Ibeing transferred via line 48 to fractionating column 50. The gasoline products, constituting a minor proportion of the total product, are recovered via lines 52 and 54, as previously described. It should be noted however that the C7+ gasoline fraction recovered via line 54 is normally somewhat less aromatic than the C7+ gasoline produced during the gasoline cycle, and hence is preferably subjected to reforming. v

The bottoms fraction in line 56 is, in this cycle, diverted via line 66 (valve 62 closed; valves 60 and 64 open) to column 68, from which the desired middle distillate product is recovered overhead via line 70. The bottoms fraction from column 68, consisting of material boiling above the end-boiling-point of the middle distillate product, is withdrawn via line 72 and recycled to hydrocracker 19 via lines 58 and 18. In this manner, total conversion of the feed to middle distillate and gasoline products is achieved, with the C4+ gasoline product `boiling below the desired middle distillate product amounting to no more than about 10-40 volume-percent of the total. It should be noted however that a substantial proportion of this gasoline product, normally about 40-60% thereof, is synthesized in the hydroner as a result of the decomposition of organic sulfur and nitrogen compounds, and that this gasoline is of substantially higher quality than the gasoline produced in hydrocracker 19 under middle-distillate hydrocracking conditions. Hence, by minimizing the conversion per pass to gasoline in hydrocracker 19, preferably to less than about 20 volume-percent based on feed to the hydrocracker, the total gasoline product recovered during the middle distillate production cycle is not substantially lower in quality than that of the total gasoline product produced during the gasoline cycle.

Due to the absence of ammonia and nitrogen compounds in hydrocracker 19 during the middle-distillate hydrocracking cycle, distinctly different conditions are maintained therein, summarized as follows:

HYDROCRACKING CONDITIONS FOR MIDDLE DISTILLATE PRODUCTION Broad Preferred range range Temperature, F- 400-750 450-650 Pressure, p.s.i.g 500-5, 000 G-3, 000 LHSV, V./V./'hr l 0. 5-20 1-10 Hn/Oil Ratio, M s.c.f./b 0. 5-20 1-10 version rates is presented. From the standpoint of maintaining maximum yields and quality of middle distillate product, and minimizing gasoline synthesis in the hydrocracker, it is preferable to operate at low per-pass conversions, e.g., 30-50 volume-percent (to products boiling below the end-point of the middle distillate product), entailing space velocities in the range of about 2-20, thus minimizing over-cracking to gasoline. This operation may tend to place a greater load on heat-exchanger 4, cooler 34, and the product fractionation equipment due to the higher oil recycle rates and resultant larger total volume of liquid products per hour owing into separator 36. However, heat-exchanger 4 and cooler 34 can handle substantially larger volume rates of product during the middle distillate cycle because of the lower temperature of the hydrocracker eflluent in line 28. Hence, substantially different throughput rates can be maintained in exchanger 4 and cooler 34 in the two cycles of operation, without expensive over-design of these units for either cycle of the operation.

If it is desired to minimize the fractionation load on columns 50 and `68 and/or to stay below the maximum design capacity of exchanger 4 and cooler 34, hydrocracker 19 may be operated at relatively high conversions per pass of e.g., 40-80 volume-percent, resulting in lower recycle oil rates and lower space velocities of e.g., 0.2-5 in the hydrocracker. The choice of these two modes of operation depends upon a great many economic considerations, but to achieve the primary objective of the process during the lmiddle distillate cycle, i.e., of minimizing the production of low quality gasoline, it is generally desirable to operate hydrocracker 19 at the minimum conversion per pass consistent with the design capacity of exchanger A4 and cooler 34. Under these conditions, the conversion per pass to gasoline in the hydrocracker can be maintained at below about 5-10 volume percent.

HYDROCRACKING CATALYSTS Suitable catalysts for use in hydrocracker 19 comprise in general any refractory, solid cracking base having a cracking activity in excess of that corresponding to a Cat-A Activity Index of about 40, upon which is distributed a minor proportion of a Group VIII metal or metal sulfide hydrogenating component. Operative cracking bases include for example mixtures of two or more refractory oxides such as silica-alumina, silica-magnesia, silica-zirconia, alumina-boria, siliea-titania, silica-zirconiatitania, acid treated clays and the like. Acidic metal phosphate gels such as aluminum phosphate may also he used. The preferred cracking bases comprise crystalline, siliceous zeolites, sometimes referred to in the art as molecular sieves, composed usually of silica, alumina and one or more exchangeable cations such as hydrogen, magnesium, rare earth metals, or other polyvalent metal ions. These zeolites are further characterized by crystal pores of relatively uniform diameter between about 4 and 14 angstroms. Suitable zeolites include for example the synthetic molecular sieves A, L, S, T, X and Y, and natural zeolites such as chabazite, mordenite, etc. It is preferred to employ zeolites having a relatively high SiO2/Al203 mole ratio, between about 3.0 and l2, and even more preferably between about 4 and 8. Specifically preferred zeolites are those of the Y and L crystal types.

The naturally occurring molecular sieve zeolites are usually 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 ion-exchanged out with a polyvalent 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:

(nmz z mo (2) Mixed polyvalent metal-hydrogen zeolites may be prepared by ion-exchanging first with an ammonium salt, then partially back-exchanging with a polyvalent 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 U.S. 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 preferred cracking bases are those which are at least about 10%, and preferably at least 20%, metalcation-deficient, based on the initial ion-exchange capacity. A specifically desirable and stable class of zeolites are those wherein at least about 20% of the ion-exchange capacity is satisfied by hydrogen ions, and at least about 10% by polyvalent metal ions such as magnesium, calcium, zinc, rare earth metals, etc.

The essential active metals employed herein as hydrogenation components are those of Group VIII, 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.

The amount of hydrogenating metal in the catalyst can vary within wide ranges. Broadly speaking, any amount between about 0.05% 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 U.S. Patent No. 3,200,083.

Following addition of the hydrogenating metal, the re sulting catalyst powder is then filtered off, dried, pelleted with added lubricants, binders, or the like if desired, calcined at temperatures of e.g., 700-l,200 F. in order to activate the catalyst and decompose zeolitic ammonium ions. The 'foregoing catalysts may be employed in undiluted form, or the powdered catalyst may be mixed and copelleted with other relatively less active adjuvants, diluents or binders such as activated alumina, silica gel, coprecipitated silica-alumina cogel, magnesia, activated clays and the like in proportions ranging between about 5% and 50% by weight. These adjuvants may be ernployed as such, or they may contain a minor proportion of an added hydrogenating metal, e.g., a Group VI-B and/ or Group VIII metal.

The following example is cited to illustrate more specifically the process as described above, but is not to be construed as limiting in scope:

Example The process as described in connection with the drawing is employed to convert a blend of catalytic cycle oil and straight run gas oil alternately to a gasoline 7 product, and a 75% .IP-5 jet fuel product. The feedstock inspections are:

The hydroiining catalyst employed consists of a presulfided composite of 3.3% NiO and 15% M003 supported on an alumina-silica cogel base containing 5% SiO2. The hydrocracking catalyst is a copelleted composite of (a) 30% 'by weight of activated alumina and (b) 70% by weight of a Y molecular sieve zeolite containing 0.5 wt. percent palladium, and wherein about 30% of the ion-exchange capacity is satisiied by magnesium ions, 10% by sodium ions and 60% by hydrogen ions.

During the gasoline cycle, total effluent from hydroner 8 passes through hydrocracker 19, with the 400 F.-{- bottoms product from column 50 being recycled to the hydrocracker. During the jet fuel cycle, the hydroiiner eiuent by-passes the hydrocracker, so that the sole feed to the hydrocracker is the 550 Fr-lbottoms fraction from column 68, comprising the 550 13.-]- hydroiiner product mixed with 550 R+ unconverted oil from the hydrocracker. The significant start-of-run conditions and results of the process are as follows:

.Tet Gasoline fuel cycle cycle Hydroning Conditions:

Temp., av. bed, F '725 730 Pressure, p.s.i.g 1,500 1,500 HSV 0. 75 1.0 Hg/Oil ratio, M s.c.f./b 8.0 8. Hydrocracking Conditions' Temp., av. bed, F 710 520 Pressure, p.s.i.g. l, 500 1, 500 LHSV 1. 2. 2 HZ/Oil ratio, M s.c.f./b 8.0 8. 0 Vol. percent conversion per pass in hydrocracker to:

400 F., E.P. gasoline and lighter 60 10 550 F., E.P. jet fuel and lighter 40 Total products, gasoline octane numbers, F-l plus 3 ml. TEL:

(J5-C5 99. 5 98. 5 C1, 400 F 90 C7, 350 F 92 87 J et fuel:

Freezing point, F Vol. percent aromatics CFR luminometer number Approximate material balances, ba els of total product per 100 barrels fresh iced:

Butanes 16 5 C5C6 gasoline 29 7 C7, 400 F. gasoline 78 1 16 35o-550 F. jet fuel 80 1 350 F. end-point.

In contrast to the foregoing, if one attempts to maximize jet fuel production using the same contacting sequence employed above for the gasoline cycle (all hydroiner efliuent passing through the hydrocracker, and with recycle thereto of the 550 F.+bottoms from column 68), but with hydrocracking temperatures reduced to obtain the same conversion per pass therein to gasoline, the yield of 350-550 F. jet fuel is less than 65 barrels per 100 barrels of fresh feed, and the aromatic content thereof is more than about 5% The following claims are believed to define the true scope of the invention, which is not limited to the exemplary details described above.

I claim:

1. In a catalytic hydroiining-hytdrocracking system, a

process for converting a mineral oil feedstock boiling above the gasoline range alternately to a relatively aromatic, predominantly gasoline product, and a relatively non-aromatic, predominantly middle distillate product boiling at least partially above the gasoline range, which comprises:

(1) passing said feedstock plus added hydrogen through a catalytic hydroiining Zone at an elevated temperature and pressure to effect decomposition of organic sulfur and nitrogen compounds contained therein;

(2) in a periodic gasoline-producing cycle, passing eliluent from said hydroining zone, Without intervening puriiication to remove ammonia `and hydrogen sulfide, through a catalytic hydrocracking zone at an elevated temperature and pressure to effect about 30-80 volume percent conversion per pass of said feedstock to gasoline-boiling-range material, thereby producing a first gasoline-rich hydrocracker eiiiuent, and recovering a predominantly gasoline product therefrom, while recycling higher boiling material to said hydrocracking zone; and

(3) in an alternating periodic middle distillate-producing cycle,

(a) blending the eiuent from said hydroiining zone with a second relatively gasoline-lean effluent produced from said hydrocracking zone as hereinafter defined,

(b) washing the resulting blend to remove ammonia (and/or hydrogen sulde) therefrom,

(c) fractionating the washed blend to recover a minor gasoline product, a major middle distillate product, and an unconverted oil boiling above said middle distillate product, and

(d) passing said unconverted oil plus hydrogen through said catalytic hydrocracking zone at an elevated pressure and a relatively low temperature controlled and correlated with liquid hourly space velocity to give a substantially lower conversion per pass to gasoline than in step (2), and to provide said hydrocracking zone effluent for step (3) (a) above;

`(4) the catalyst employed in said hydrocracking zone being the same for each of steps (2) and (3).

2. A process as defined in claim 1 wherein the catalyst employed in said hydroning zone is essentially an Iron Group metal oxide or suliide plus a Group VI-B metal oxide or sulfide supported on an adsorbent oxide carrier having a Cat-A cracking Activity Index below about 25, and wherein the catalyst employed in said hydrocracking zone is essentially a Group VIII metal or an oxide or sulfide thereof supported on a solid cracking base having a cracking activity in excess of that corresponding to a Cat-A Activity Index of about 40.

3. A process as defined in claim 2 wherein said hydrocracking catalyst is essentially a minor proportion of a Group VIII noble metal or an oxide or sulfide thereof, supported on a crystalline, zeolitic molecular sieve cracking base wherein the zeolitic cations are predominantly of the class consisting of hydrogen ions and polyvalent metal ions.

4. A process as defined in claim 1 wherein said middle distillate product is a jet fuel boiling mainly in the 350- 550 F. range.

5. A process as defined in claim 1 wherein said middle distillate product is a diesel oil boiling mainly in the 400- 800 F. range.

6. A process as defined in claim 1 wherein said middle distillate product is a stove oil boiling mainly in the 400- 600 F. range.

7. In a catalytic hydroiining-hydrocracking system, a process for converting a mineral oil feedstock boiling above the gasoline range alternately to a relatively aromatic, predominantly gasoline product, and to a relatively non-aromatic predominantly middle distillate product boiling at least partially above the gasoline range, which comprises:

(1) passing said feedstock plus added hydrogen through a catalytic hydroning zone at an elevated temperature and pressure to effect decomposition of organic sulfur and nitrogen compounds contained therein vwithout substantial cracking of hydrocarbons;

(2) in a periodic gasoline-producing cycle,

i(a) passing effluent from said hydroining zone,

without intervening purification to remove ammonia and hydrogen sulfide, through a catalytic hydrocracking zone at an elevated pressure and a temperature between about 600 and 900 F., in contact with a hydrocracking catalyst comprising essentially a minor proportion of a Group VIII noble metal or an oxide or sulfide thereof, supported on a crystalline zeolitic molecular sieve cracking base wherein the zeolitic cations are predominantly of the class consisting of hydrogen ions and polyvalent metal ions to eli'ect about 30-80 volume-percent conversion per pass of said feedstock to gasoline-boiling-range material, thereby producing a iirst, gasoline-rich hydrocracker eflluent,

(b) separating said rst eluent into a gasoline product and an unconverted recycle oil; and

(c) recycling said unconverted recycle oil to said hydrocracking zone; and

(3) in an alternating periodic middle distillate-producing cycle,

(a) blending the eluent from said hydroining zone with Ia second, relatively gasoline-lean efuent produced from said hydrocracking zone as hereinafter defined,

(b) washing the resulting blend to remove ammonia therefrom,

(c) fractionating the washed blend to recover a minor gasoline product, a major middle distillate product, and an unconverted oil boiling above said middle distillate product, and

(d) passing said unconverted oil plus hydrogen through said catalytic hydrocracking zone at an elevated pressure and a relatively low temperature between about 400 and 750 F., said temperature being correlated with liquid hourly space velocity to give less than about 20 volume-percent conversion per pass to gasoline, and to provide said hydrocracking zone eluent for step (3) (a) above;

(4) the catalyst employed in said hydrocracking zone being the same for each of steps I(2) and (3).

8. A process as defined in claim 7 wherein temperatures and space velocity in said hydrocracking step (3) (d) are controlled and correlated to give between about 30- volume-percent conversion per pass to products boiling below the end-point of said middle distillate product, whereby gasoline production is minimized therein.

9. A process as defined in clairn 8 wherein a substan tially higher liquid hourly space velocity and substantially lower temperatures are maintained in said hydrocracking step (3) (d) than in said hydrocracking step (2) (a).

10. A process as defined in claim 1 wherein the temperature and liquid hourly space velocity are correlated to give a conversion per pass to middle distillate and lighter products in step (3)(d) of about 30-50 volume-percent.

11. A process as defined in claim 1 wherein said substantial conversion per pass in step (2) is about 30-80 volume-percent and wherein said substantially lower conversion per pass to gasoline in step (3)'(d) is less than about 20 volume-percent.

12. A process as defined by claim 11 wherein the conversion per pass to middle distillate and lighter products in step (3) (d) is about 30-50 volume-percent.

References Cited UNITED STATES PATENTS 5/ 1964 Helfrey et al. 208-89 1/1968 Kelley 208--89 U.S. C1. X.R. 208-1 11 UNITED STATES PATENT OFFICE CERTIFICATE 0F CORRECTION Patent No. 3,468,788 September 23, 1969 Herbert P. Wilkinson It is certified that error appears in the above identified patent and that said Letters Patent are hereby corrected as shown below:

Column 8, line 30, Cancel "(and/or hydrogen sulfide)".

Signed and sealed this 21st day of April 1970u (SEAL) Attest:

Edward M. Fletcher, Jr.

Attesting Officer Commissioner of Patents WILLIAM E. SCHUYLER, JR.

Claims (1)

1. IN A CATALYTIC HYDROFINING-HYDROCRACKING SYSTEM, A PROCESS FOR CONVERTING A MINERAL OIL FEEDSTOCK BOILING ABOVE THE GASOLINE RANGE ALTERNATELY TO A RELATIVELY AROMATIC, PREDOMINANTLY GASOLINE PRODUCT,, AND A RELATIVELY NON-AROMATIC, PREDOMINANTLY MIDDLE DISTILLATE PRODUCT BOILING AT LEAST PARTIALLY ABOVE THE GASOLINE RANGE, WHICH COMPRISES: (1) PASSING SAID FEEDSTOCK PLUS ADDED HYDROGEN THROUGH A CATALYTIC HYDROFINING ZONE AT AN ELEVATED TEMPERATURE AND PRESSURE TO EFFECT DECOMPOSITION OF ORGANIC SULFUR AND NITROGEN COMPOUNDS CONTAINED THEREIN; (2) IN A PERIODIC GASOLINE-PRODUCING CYCLE, PASSING EFFLUENT FROM SAID HYDROFINING ZONE, WITHOUT INTERVENING PURIFICATION TO REMOVE AMMONIA AND HYDROGEN SULFIDE, THROUGH A CATALYTIC HYDROCRACKING ZONE AT AN ELEVATED TEMPERATURE AND PRESSURE TO EFFECT ABOUT 30-80 VOLUME PERCENT CONVERSION PR PASS OF SAID FEEDSTOCK TO GASOLINE-BOILING-RANGE MATERIAL, THEREBY PRODUCING A FIRST GASOLINE-RICH HYDROCRACKER EFFLUENT, AND RECOVERING A PREDOMINANTLY GASOLINE PRODUCT THEREFROM, WHILE RECYCLING HIGHER BOILING MATERIAL TO SAID HYDROCRACKING ZONE; AND (3) IN AN ALTERNATING PERIODIC MIDDLE DISTILLATE-PRODUCING CYCLE, (A) BLENDING THE EFFLUENT FROM SAID HYDROFINING ZONE WITH A SECOND RELATIVELY GASOLINE-LEAN EFFLUENT PRODUCED FROM SAID HYDROCRACKING ZONE AS HEREINAFTER DEFINED. (B) WASHING THE RESULTING BLEND TO REMOVE AMMONIA (AND/OR HYDROGEN SULFIDE) THEREFROM,

Images