US3132090A - Hydrocracking process with regulation of the aromatic content of the product - Google Patents

Description

May 5, 1964 P. F. HELFREY ETAL 3,132,090

HYDROCRACKING PROCESS WITH REGULATION OF THE AROMATIC CQNTENT OF THE PRODUCT Filed Jan. 23, 1962 HYDROCRACKING PROCESS WITH REGULA- THGN OF THE ARQMATIC CONTENT F THE PRODUCT Paul F. Helfrey, Montehello, Nichoias L. Kay, Fullerton, Berna] Peralta, Anaheim, and Cioyd P. Reeg, Grange, Calif assignors to Union Oil Company of California, .Los Angeles, (ialiti, a corporation of California Filed Jan. 23, 1.962, Ser. No. 168,058 1% Claims. (Cl. 208-8 I Thisinvention relates to a catalytic hydrocracking process affording a maximum degree of flexibility in reference to variety and 'quality of products obtainable therefrom. More specifically, the process is designed to produce from hydrocarbon feedstocks a relatively aromatic product boiling in the gasoline range, and/or relatively nonaromatic products boiling in the jet fuel-diesel fuel range. In broad aspects, the principal operative features of the process comprise contacting the hydrocarbon feedstock With a group VIII noble metal hydrocracking catalyst at pressures below about 2,500 p.s.i.g. and temperatures between about 400and 750 F., and adjusting the hydrogen sulfide concentration in the reaction mixture upwardly when the major desired product is high octane gasoline, and downwardly when the major desired product is a highly parafiinic jet fuel and/ or diesel fuel. In a preferred aspect of the invention, two separate hydrocracking stages are employed, the first operating in the presence of hydrogen sulfide and nitrogen compounds and at relatively high temperatures to produce high octane gasoline, and the second operating with a group VIII noble metal hydrocracking catalyst, and substantially in the absence of nitrogen compounds, to produce either high octane gasoline when the hydrogen sulfide concentration is relatively high, or highly saturated jet fuels and/or diesel fuels at lower hydrogen sulfide concentrations. The process thus affords maximum flexibility, permitting the refiner to meet changing market demands for the various products, while minimizing the quantity of low octane gasoline produced which must be subjected to a subsequent severe reforming step to obtain the desired octane balance in the refinery.

A principal object of this inventionis to provide an integrated hydrocracking process designed mainly for the production of gasoline, but which can be easily regulated to produce a high quality jet fuel boiling for example in immediately reflected in a significant change in produuct aromaticity even without change in hydrocracking temperature. This sensitivity does not appear to be displayed at temperatures above about 750 F., while at pressures above about 2,500 p.s.i.g., the magnitude of the effect is substantially decreased. It is found also that this reversible sensitivity to hydrogen sulfide concentration is not displayed in the same order of magnitude by other hydrocracking catalysts such as those wherein the hydrogenating component is nickel. It is found also that variations in hydrogen sulfide concentration within the range above about 0.5 millimole, or in the range below about 0.01 millilmole per mole of hydrogen, bringabout relatively insignificant changes in product aromaticity. The critical concentration range for practical purposes hence appears to lie between about 0.01 and 0.5 millimole. While we do not wish to be found by any theoretical explanation for this observed sensitivity to hydrogen sulfide concentration, it would appear to involve in some degree a change in the group VIII noble metal hydrogenating component from the free metal to a sulfide state, and vice versa. But we do not exclude the possibility that other operative factors may be involved.

Another critical feature of the process resides in the use of an initial hydrocarbon feedstock which is substantially aromatic in character. This includes coker distillate gas oils, cycle oils derived from catalytic or thermal cracking operations, as well as aromatic straight-run gas oils. These feedstocks 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 1,000 F., having an API gravity of about 20-35, and containing at least about 20% by volume of aromatic hydrocarbons. Such oils may also contain from about 0.1% to 5% by weight of sulfur and from about 0.01% to 2% by weight of nitrogen. Aromatic feedstocks of this character are required inasmuch as the low temperatures and relatively high pressures required do not thermodynamically favor the synthesis of aromatics from non-aromatics, and hence the aromatics appearing in the product are primarily unhydrogenated fragments of high boiling arothe 350-550 F. range, and/or a high quality diesel fuel boiling for example in the 400-800" F. range. A further objective is to minimize the total reforming capacity required in any given refinery to produce the desired quantity of high octane gasoline. Another object is to provide a flexible hydrocracking process which will permit the refiner to shift rapidly and easily from jet fuel to gasoline products as his market may require.

tion which follows.

The invention rests basically upon our discovery that, within the temperature range of about 400750 F., group VIII noble metal hydrocracking catalysts are remarkably sensitive to hydrogen sulfide concentration in the reaction mixture, in respect to product aromaticity. Moreover, this sensitivity is reversible, and is such that variations in hydrogen sulfide concentration, within the range of about 0 to 0.5 millimole per mole of hydrogen, are substantially Other objects will be apparent from the more detailed descripinatics initially present in the feed. If non-aromatic feedstocks were employed, the products obtained under the conditions of pressure and temperature employed herein would be almost entirely paralfinic and/or naphthenic, regardless of the hydrogen sulfide concentration in the reaction mixture.

While as noted above, the initial feedstock may contain nitrogen compounds, it is important to note that in, the particular contacting stage of the process in which the hydrogen sulfide concentration is to be varied periodically in order to change the aromaticity of the product,nitrogen should be substantially absent, i.e.', below about 25 parts per million by weight, based on hydrocarbon feedstock. The presence of nitrogen compounds militates against a flexible operation (embracing in one cycle the production of highly saturated products), because of the relatively high temperatures, usually above about 700 R, which are required in order to overcome the poisoning effect of the nitrogen compounds. And, as noted above, at these high'temperatures, variations in hydrogen sulfide concentration are relatively insignificant with respect to Patented May 5, 1964 product aromaticity, the product always being substantially aromatic in character.

It is further to be noted that the hydrocracking catalyst employed in the zone where flexibility of product aromaticity is desired, should preferably be one comprising a very active cracking base. This is required in order to obtain the desired cracking activity at temperatures below 750 F. In general, the cat-A cracking activity of the cracking component should be at least about 25 and preferably greater than about 35. Catalysts of this nature will be described more in detail hereinafter.

The process of this invention may be operated either in a single stage or in plural stages of hydrocracking. Raw feedstocks may be employed in many instances, but in most cases it is preferable to employ a hydrofining pretreatment to effect at least partial desulfurization, denitrogenation, stabilization, etc. Where the feedstock contains substantial quantities of nitrogen compounds, it is normally preferable to employ two. stages of hydrocracking, and still more preferable a preliminary hydrofining treatment ahead of the first hydrocracking stage. The hydrofining treatment in this instance may desirably be of the integral type, i.e., wherein the entire hydrofiner effluent is passed directly through the first hydrocracking stage without intervening condensat on or purification.

Since in these multi-stage operations, the first stage feed will normally contain substantial quantities of sulfur and/or nitrogen compounds in the form of hydrogen sulfide and ammonia, the first stage will usually be operated exclusively for gasoline production, since the hydrocarbon product is inherently substantially aromatic. The feed to the second hydrocracking stage is primarily the unconverted oil from the first stage, and is substantially free of nitrogen compounds and sulfur compounds. The second stage may hence be operated with any desired concentration of hydrogen sulfide present. The desired hydrogen sulfide concentration can be maintained for example by blending the feed with a sulfur-containing feed, varying the proportion of hydrogen sulfide-containing recycle gas employed therein, simply adding hydrogen sulfide, or any equivalent method. Any product oil from the second hydrocracking stage which is not converted to the desired boiling range, is normally recycled back to that stage.

Reference is now made to the attached drawing, which or heating, depending upon or isothermally, and under the following general conditions:

HYDROFINING 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 10 parts per million.

The total hydrofined product from hydrofiner 8 'is withdrawn via line 10 and transferred via heat exchanger 12 to first-stage hydrocracker 14, without intervening condensation or separation of products. Heat exchanger 12 is for the purpose of suitably adjusting the temperature of feed to hydrocracker 14; this may require either cooling the respective hydrofining and hydrocracking temperatures employed. Inasmuch as first-stage hydrocracker 14 and hydrofiner 8 are preferably operated at essentially the same pressure, it is entirely feasible to enclose both contacting zones within a single reactor, using appropriate temperature control means.

The catalyst employed in reactor 14 may consist of any desired combination of a refractory cracking base with a suitable hydrogenating component. Suitable cracking bases include for example mixtures of two or more difficultly reducible oxides such as silica-alumina, silicamagnesia, silica-Zirconia, alumina-boria, silica-titania, silica-zirconia-titania, acid treated clays and the like. Acidic metal phosphates such as aluminum phosphate may also be used. The preferred cracking bases comprise composites of silica and alumina containing about 50-90% silica; coprecipitated composites of silica, titania, and zirconia containing between 5% and 75% of each component; partially dehydrated, zeolitic, crystalline molecular sieves, e.g., of the X or Y crystal types, having 'i relatively uniform pore diameters of about 8 to 14 is a flow sheet illustrating the invention in one of its multi-stage adaptations. In the succeeding description, it will be understood that the drawing has been simplified by the omission of certain conventional elements such as valves, pumps, compressors, and the like. Where heaters or coolers are indicated, it will be understood that these are merely symbolic, and in actual practice many of these will be combined into banks of heat exchangers and fired heaters, according to standard engineering practice. The product fractionating equipment is merely illustrative of a system providing for maximum flexibility in handling differentfeedstocks and products; in actual practice, specific desired product distributions would require modifications in the fractionating equipment for maximum economy.

In the drawing, the initial feedstock is brought in via line 2, mixed with recycle and makeup hydrogen from line 4, preheated to incipient hydrofining temperature in heater 6, and then passed directly into hydrofiner 8, where catalytic hydrofining proceeds under substantially conventional conditions. Suitable hydrofining catalysts include for example mixtures of the oxides and/ or sulfides 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 sulfides of the group VIB and/ or group VIII metals, preferbaly supported on adsorbent oxide carriers such as alumina, silica, titania, and the like. The hydrofining operation may be conducted either adiabatically angstroms, and comprising silica, alumina and one or more exchangeable zeolitic cations.

A particularly active and useful class of molecular sieve cracking bases are those having a relatively high SiO /Al O ratio, e.g., between about 2.5'and 6.0. The most active forms are those wherein the exchangeable zeolitic cations are hydrogen and/ or a divalent metal such as magnesium, calcium or zinc. In particular, the Y molecular sieves, wherein the SiO /Al O ratio is about 5, are preferred, either in their hydrogen form, or a divalent metal form. Normally, such molecular sieves are prepared first in the sodium or potassium form, and the monovalent metal is ion-exchanged out with a divalent metal, or where the hydrogen form is desired, with an ammonium salt followed by heating to decompose the zeolite ammonium ion and leave a hydrogen ion. It is not necessary to exchange out all of the monovalent metal; the final compositions may contain up to about 6% by weight of Na O, or equivalent amounts of other monovalent metals. Molecular sieves of this nature are described more particularly in Belgian Patents Nos. 577,642, 598,582, 598,683 and 598,682.

As in the case of the X molecular sieves, the Y sieves also contain pores of relatively uniform diameter in the individual crystals. In the case of X sieves, the pore diameters may range between about 6 and 14 A., depending upon the metal ions present, and this is likewise the case in the Y sieves, although the latter usually arefound to have crystal pores of about 9 to 10 A. in diameter.

' The foregoing cracking bases are compounded, as by impregnation, with from about 0.5% to 25% (based on free metal) of a group VIB or group VIII metal promoter, e.g., an oxide or sulfide of chromium, tungsten,

combination thereof. Alternatively, even smaller propor-' tions, between about 0.05% and 2% of the'metals platinum, palladium, rhodium or iridium may be employed.

without significant amounts of sulfur being present. In the modification illustrated, variations in sulfur COHCBII? tration in hydrocracker 62 are obtained by the alternate use of separate and mixed hydrogen recycle gas systems from hydrocrackers l4 and 62. The recycle gas from separator 24 normally contains a substantial proportion of hydrogen sulfide which was not removed by the previous Water-washing operation. To operate hydrocracker 62 substantially in the absence .of sulfur (separate recycle The oxides and sulfides of other transitional metals may systems) valve 51' is opened and valves 52 and 54 closed, also be used, but to less advantage than the foregoing. thus sending thesour recycle gas from line 26 through In the-case of zeolitic type cracking bases, itis desirline 4, back to'hydrocracker 14, and the sweet recycle able to deposit the hydrogenating metal thereon by'ion gas from separator 68 back to hydrocracker 62 via lines exchange. This can be accomplished by digesting the 70 and 58. To operate with added sulfur, valve 51 is zeolite with an aqueo s s l ion of a siu able compoun closed and valves 52 and 54 opened, thereby diverting sour f the desired metal. wherein t m tal is p es t in a recycle gas from line 26 into lines 55 and 70, where it' cationic form, a n r i g to m e fr e metal, mingles with sweet recycle gas from separator 68. The as described for example in Belgian Patent No. 598,686. mixed gases are then split, one portion flowing to by- A particularly suitable class f hydrocrackingcatalysts drofiner 8 via lines 56 and 4, and the other portion flowis composed of about 75 95% by weight of a coprecipiin to hydrocracker 62 vi li 58, tated hydrocracking'base containing 5.-75% SiO 5-75% In the sweet cycle of operation in hydrocracker 62, it z. and 575% z, and i rp ra ed therein from is normally desirable to adjust the process variables, princiabout 5-25%, based on free metal, of a group VIII metal pally temperature, so as tomaximize jet fuel and/or or metal sulfide,-e.g., nickel or nickel sulfide. I diesel fuel production, and minimize the conversion to The process conditions in hydrocracker 14 are suitgasoline. Specifically, it is preferred to limit the conably adjusted so as to provide about'20-60% conversion version to gasoline to below about 20% per pass by to gasoline per pass, while at the same time permitting volume. To achieve this objective, while obtaining maxirelatively long runs betweenregenerations, i.e., from about mum quality of the jet fuel-diesel fuel products, the hy 2 to 8 months. The specific selection of per n condrogen sulfideconcentration should be maintained at a ditions depends largely on the nature of the feedstock, value below about 0.2, and preferably below about 0.01 Pressures in the high range n r ly being used for millimole per mole of hydrogen, and the other process highlyaromatic feeds, feeds With g en p in s. The conditions are adjusted within the following ranges: range of operative conditions contemplated for reactor a 14 are as follows, assuming the feed thereto contains more SWEET SECOND-STAGE OPERATING than about 25 parts per million of nitrogen: CONDITIONS FIRST-STAGE HYDROCRACKING CONDITIONS Operative Preferred Operative Preferred hesa ragfzz::;:;;::::::;:;::;::;; 0 3333 se jfii v LHSV, 7./v./hr 0.5-15 1-10 Temperature, 0 0 650-800 2/0 '3 -0-f-/b 500-20, 000 2, 000-12, 000 Prcssure,p.s.i.g ioeasoo sou-2,000 Y LHSV,v. .lhr 0,5 10 1-5 f "f 500 2O-O00 000-10 000 As Will be understood by those skilled in the art, the 40 specific selection of operating conditions within these ill depend on several factors, mainly the relative,- The effluent from hydrocracker 14 is withdrawn via -3 r line 16, condensed in condenser 18, then mixed with wash Ega h catalysL and general reflactonness of the W miected Via km 20 mm and the f In the sour cycle of operation in hydrocracker 62, the mlxture is "than transfeiredfio hlgh-prpssilre g g f process variables are normally adjusted so as to obtain Sour recycle hydrogen}? wlthfimwn Vlahne an aque' a maximum conversion to gasoline per pass which is ous Wash Water comamni-g dl.ssqlved i i and some consistent with the desired run length (which in turn de f i hydrogen Sulfide Wlthdrawn i h 2 pends on the rate of catalyst deactivation), and desired 9 hYdmFarbOH phas's m Separator 24 a en i efiiciency' of conversion to gasoline. If the crack per pass hue mm 1 'Pressure sepegator 3 from a i 5 is too high, the catalyst deactivation rate is accelerated, fl gases.compnsmg.meihane at P t an e and a relatively large proportion of feed is converted to hke wlthdmwn Via hue The hquid bydrocar' C C dry' gases and butanes. Satisfactory run lengths siaparator 32 are then transferred lme 36 to (3-12 months) and conversion cfliciencies are normally fracnonimngpolumln 38 t d th obtained at conversions to 400 F. end-point gasoline of Fractlonatmg 9 1S opefae gnmany or s between about 40 and 80% by volume per pass. To purpoLse of recovermg C7+ gasolme an an unconveii achieve these objectives, while obtaining maximum gasogas 9 feed. the second'stage hyqrocracker Llg t line quality, the hydrogen sulfide concentration should gasolme bollmg up to the C6 range 15 normally taken be'rnaintained at a value above about 0 01 and preferoif as overhead via line 40. The C7+ gasoline is withably above about 02 millimole h dr drawn as a side-cut via line 42. Thebottoms from colp61 9 y 0 and the other process conditions are ad usted Within the umn 38 constitutes the prnnary feedstock for the secondfollowing ranges v V stage hydrocracking,-and is withdrawn via line 50 for that f a P 'P Q SOUR SECOND-STAGE OPERATING The second-stage feedstock in line 50 1s then mixed CONDITIONS with recycle and makeup hydrogen from line 58, preheated to incipient hydrocracking temperatures in heater 65 O 60, and passed into second-stage hydrocracker 62. This pm We Preferred feedstock differs considerably from the feed to the firsta a stage hydrocracker, in that it is substantially free of 403 2288 80 6 3133 nitrogen compounds and sulfur compounds. The choice t V: (15-10 5 is thus presented of operating the second stage with or Hz/ 011 .50H0000 200042000 The catalyst used in hydrocracker 62 comprises about ODS-3% by weight of a group VHI noble metal supported on substantially any of the cracking bases previously described for use in hydrocracker 14. Specifically included are the metals, ruthenium, rhodium, palladium, osmium, iridium and platinum, with palladium being preferred. Specifically preferred cracking bases are the high-silica zeolitic molecular sieves, and especially the decationized or divalent metal forms of the Y molecular sieves previously described.

At the conversion levels and conditions prescribed for the second-stage hydrocracker, the run length between regenerations can be adjusted to coincide substantially with the run length in reactor 14, e.g., between about 3 and 12 months. In extended runs such as these, it is normally preferably to maintain substantially constant conversion in each stage by incrementally raising the temperature as the activity of the catalyst declines. The rate of catalyst activity decline in reactors 14 and 62 under the prescribed conditions is such that constant conversion in both reactors can be obtained by raising the respective temperatures between about 0.1 and 3 F. per day, on the average. The average temperature in hydrocracker 62 will normally be about 25-l25 F. lower than the average temperature in hydrocracker 14.

The total efiiuent from hydrocracker 62 is withdrawn via line 64, condensed in cooler 66 and transferred to high pressure separator 68, from which recycle hydrogen is withdrawn via line 70 and utilized as previously described. The condensed hydrocarbons in separator 68 are then flashed via line 72 into low pressure separator 74, from which C -C flash gases are withdrawn via line 76. The liquid hydrocarbon product in separator 74 is withdrawn via line 78 and transferred to second-stage product fractionation column 80, wherein it is fractionated into various gasoline, jet fuel and diesel fuel fractions, as may be desired. Light gasoline blend- 7 ing stock is withdrawn as overhead via line 82, gasoline via line 84-, a diesel bottoms fraction via line 86, and a jet fuel side-cut via line 88. The jet fuel sidecut is transferred to a small stripping column 90, from which overhead gasoline hydrocarbons are returned to column 80 via line 92. During the sweet operating cycle, the entire bottoms from stripper 90 may be withdrawn from the system via line 94 and sent to jet fuel blending and storage facilities. During the sour operating cycle, the jet fuel fraction from line 94 is normally diverted via line 96 and recycled to second-stage hydrocracker 62 via line 50. Also, during the sour operation, the diesel fraction withdrawn as bottoms via line 86 is recycled via lines 98 and 50 to hydrocracker 62. Where both the jet fuel and diesel fuel fractions are to be recycled, there is of course no need to recover them separately, and hence both fractions can be recovered as bottoms from column 80 and recycled via lines 86, 98 and 50, thus eliminating the need for stripping column 90. Where maximum jet fuel production is desired during the sweet operation, all or a portion of the diesel fraction in line 86 may be recycled to hydrocracker 62 for conversion to jet fuel. It will be understood that the choice of the various recycle alternatives depends largely upon the desired refinery balance and market demands.

The following'examples are presented to illustrate certain critical variables in the process, as well to illustrate the operation and results of the process as above described in connection with the drawing. These examples should not however be construed as limiting in scope:

Example I This example demonstrates the remarkable flexibility drogen in the reaction mixture. The catalyst was a copelleted mixture of (1) 50% by weight of 100325 mesh activated alumina, the alumina being impregnated with 25% by weight of nickel oxide, and (2) 50% by weight of a powdered, decationized Y molecular sieve loaded by ion-exchange with 0.5% by weight of palladium. Process conditions constant throughout the run were:

Pressure, p.s.i.g 1,500 LHSV 1.5 H /oil ratio, s.c.f./b 8,000

During the initial 450 hours of processing at about 50% conversion to C -400 F. end-point gasoline (temperature, 560575 F.), and with no sulfur added to the feed, the product characteristics were as follows:

400 F.+gas gasoline oil Total aromatics, vol. percent 0.4-0. 6 0 l-O. 5 Octane No.:

F1 +3 ml. TEL., 73.8-75.0 F-l clear 51. 5-54. 0

The foregoing operation was then modified for 8 hours by incorporating 0.5% by weight of sulfur (as thiophene) in the feed, corresponding to about 2.44 millimoles of H 5 per mole of hydrogen. At 620 F. (to maintain 50% conversion to gasoline), the product characteristics were as follows:

C 400 F. 400 F.+gas

gasoline oil Total aromatics, vol. percent 19. 9 22 Octane No.:

F-l +3 ml. TEL 83.4 F-l clear 65. 5

After the 8-hour run with 0.5% sulfur in the feed,

It will thus be apparent that product aromaticity is almost immediately responsive to changes in sulfur concentration. This responsiveness however, is only apparent at temperatures below about 750 F., for when the above run was continued without sulfur until the temperature level reached 725 F. (to maintain the 50% conversion to gasoline with the relatively more deactivated catalyst), the C -400 F. gasoline product contained 31.5% aromatics which is only slighly lower than the aromaticity obtainable at this temperature in the presence of added sulfur. However, the efiiciency of conversion to C -400 F. gasoline was only 42% at 725 F., compared to at 570-625 F. Efiiciency is a measure of the proportion of feed converted which went to the desired product, and in this case is expressed as: I

(volumes of C-;400 F. gasoline, percent of fresh feed) :-(total volume percent conversion of fresh feed) X Example 11 I A run similar to that of Example I was carried out, using as the catalyst a magnesium Y molecular sieve containing about 3% by weight of zeolitic magnesium, and loaded by ion-exchange with 0.5% by weight of palladium (Linde hydrocracking catalyst MB 53 82). During a prerun period this catalyst was heavily sulfided for 8 hours with a feed containing 2.5% sulfur. Sulfur-free feed (7 p.p.m. sulfur) was then fed to the unit for 44 hours at 600 F., and then 0.5% sulfur was added for 8 hours at 641 F. The results were as follows:

Octane No. Total Aromatics, voLpercent F-l +3 ml. F-l clear TEL Sulfur-free feed, 600 F::

, 1 67-400 F. gasOline 0.4 73. 4 52. 5 400 F.+gas'oil 0.22 0.5% sulfur feed, 641 F.:

C1-400 I gasoline 20.9 85.2 68.0 400 F.+gas oil 23. 3

Example III This example illustrates preferredtechniques and results obtainable in practicing the invention in a two-stage modification, substantially as illustrated in the drawing. The

catalyst used in the hydrofining pretreatment is 3% C and 15% M00 on a carrier composed of SiO coprecipitated with 95% A1 0 the catalyst being sulfided before use. The catalyst used in both stages of hydrocracking is similar to that of Example I, being a copelleted mixture of (1) 50% by weight of 100-325 mesh activated. alumina, the alumina being impregnated with 25% by Weight NiO, and (2) 50% by weight of a powdered, decationized Y molecular sieve loaded by ion exchange with 1% by weight of palladium. The initial feed is a blend of coker distillate and thermally cracked gas oils derived from California crude oils. After an initial hydrofining treatment, the total hydrofining effluent is passed to the first stage of hydrocracking where hydrocracking proceeds in the presence of the ammonia and hydrogen sulfide formed during hydrofining. The firstrstage hydrocracking efiluent is water-washed and fractionated to recover gasoline product fractions, and a substantially sulfur-free gas oil which constitutes feed to the second stage of hydrocracking. I

The second hydrocracking stage is operated alternately with a sour recycle hydrogen stream (about 0.5% by volume H 8), and with a sweet recycle hydrogen stream (less than 10 parts per million H 8). During the sour recycle sequence, the reactor efiiuent is condensed and fractionated to recover gasoline product fractions, and the remaining oil boiling above the gasoline range is recycled to the second hydrocracking stage.

During the sweet recycle sequence the reactor effluent is fractionated to recover a light gasoline (C -C and heavier products such as a C7+ gasoline 'for reforming, a jet fuel fraction, and a diesel fuel fraction. Any remaining oil boiling above the end-point of the heaviest product desired is recycled to the second hydrocracking stage.

In .one of the sweet second-stage operations shown be low, the heaviest desired product is a 330-483 F. jet fuel; in another, a C -plus reformer charge stock and a 370- be clear that a jet fuel fraction could also be produced simply by changes in the fractionation employed, without altering the reactor operating conditions or the volume and composition of the recycle oil. 1

, 630 F. diesel fuel are produced. In the latter case, it Will 1 10 The significant conditions and results of the process are as follows.

Initial feedstock:

Boiling range, F 400-857 Gravity, API 22.2 Aromatics, wt. percent 37 Nitrogen, wt. percent 0.345 Sulfur, wt. percent 2.1 Hydrofining conditions:

Temperature, av. bed, F 725 Pressure, p.s.i.g 1,500 LHSV 0.75 H oil ratio, s.c.f/ b 8,000 First-stage hydrocracking conditions:

Temperature, av. bed, F 765 Pressure, p.s.i.g 1,500 LHSV 1.5 l-l /oil ratio, s.c.-f./b 8,000 Conversion per pass to 400 F. E.P. gasoline and lighter, vol. percent 40 Second-stage hydrocarcking conditions:

Temperature, av. bed, F.- 7

During sour operation 650 During sweet operation for jet fuel 565 During sweet operation for diesel fuel 545 Pressure, p.s.i.g 1,500 LHSV During sour operation 1.5 During sweet operation 1.3 'H /oil ratio, s.c'.f./b 8,000 Conversion per pass, vol. percent- To 400 F. E.P. gasoline during sour operation 60 To 483 F. E.P. jet fuel during sweet operation for jet fuel 70 To 630 F. E.P. diesel fuel during sweet operation for diesel fuel 70 Second Stage First Sweet Stage Sour- Gasoline J et Fuel Diesel Case Case Fuel Case Gaasoline Products, Octane Numers:

Light Gasoline (C -C5) F-l +3 1111. TEL 99. 5 99.5 C =plus Gasoline:

87. 2 80. 0 End Point, F 400 400 Second $tage Jet Fuel Product:

BOlllIlg Range, F Aniline-Gravity Product Freezing Point, F Volume Percent Aromatics- CFR Lurm'ncmeter Number- Secotad Stage Diesel Fuel Prod- Boiling Range, F Aniline Point, F Sulfur, wt. Percent Volume Percent Aromatics. Cctane Number Approximate Material Balances, Barrels of combined First and Second Stage Products per Barrels Fresh Feed:

Butanes C5-C6 Gasoline C1-400 F. Gasoline. .TetFuel. Diesel "Fuel" Higher octane numbers here reflect low end-points of the respective gasolines, as compared to the 400 F. end-point gasoline produced during the sour operation. If 400 F. end-point gasolines were produced during these sweet operations, their octane numbers would be lower than that of the gasoline from the sour operation.

Results analogous to those indicated in the foregoing examples are obtained when other hydrocracking catalysts and conditions, other feedstocks and other hydrofining 1 1 conditions within the broad purview of the above disclosure are employed. It is hence not intended to limit the invention to the details of the examples or the drawing, but only broadly as defined in the following claims.

We claim:

1'. In a catalytic hydrocracking process wherein a hydrocarbon feedstock boiling above the gasoline range and containing aromatic hydrocarbons is contacted with a group VIII noble metal-containing hydrocracking catalyst in the presence of added hydrogen at a pressure between about 400 and 2,500 p.s.i.g. and at a temperature between about 400750 F. selected to give a substantial conversion to at least one desired product selected from the class consisting of gasoline, jet fuel and diesel fuel, the improved method for controlling and alternately varying the aromaticity of said desired product which comprises: maintaining alternately (A) a relatively high, continuous concentration of hydrogen sulfide, above about 0.01 millimole thereof per mole of hydrogen in the hydrocracking zone, to produce a relatively aromatic product, and (B) a relatively low continuous partial pressure of hydrogen sulfide, below about 0.2 millimole thereof per mole of hydrogen in the hydrocracking zone, to produce a relatively non-aromatic product.

2. A process as defined in claim 1 wherein said group VIII noble metal is palladium.

3. A process as defined in claim 1 wherein said hydro cracking catalyst comprises a zeolitic, alumino-silicate molecular sieve' of the Y crystal type containing zeolitic cations from the class'consisting of hydrogen and divalent metals.

4. A process as defined in claim 1 wherein the sulfur concentration during said alternate (A) is maintained at above about 0.2 millimole per mole of hydrogen, and during said alternate (B) at below about 0.01 millimole per mole of hydrogen.

5. In a hydrocarbon conversion process wherein a hydrocarbon feedstock containing aromatic hydrocarbons and boiling above the gasoline range is first subjected to a hydrogenating treatment wherein organic sulfur com pounds are decomposed and removed, and wherein essentially sulfurand nitrogen-free hydrocarbon efliuent boiling above the gasoline range from said first hydrogenating treatment is subjected to a subsequent hydrocracking step in contact with a group VIII noble metal-containing hydrocracking catalyst and at a temperature between about 400750 F. to produce at least one product selected from the class consisting of an aromatic gasoline, a nonaromatic jet fuel and a non-aromatic diesel fuel; the improvement which comprises maintaining in said subsequent hydrocracking step a concentration of hydrogen sulfide which is (A) relatively low, less than about 0.2 millimole per mole of hydrogen when the desired product to be recovered therefrom is selected mainly from the jet fuel-diesel fuel class, and (B) relatively high, greater than about 0.01 millimole per mole of hydrogen when the desired product to be recovered is mainly gasoline.

6. A process as defined in claim 5 wherein said group VIII nobel metal is palladium.

7. A process as defined in claim 5 wherein said hydrocracking catalyst comprises a zeolitic, alumino-silicate molecular sieve of the Y crystal type containing zeolitic cations from the class consisting of hydrogen and divalent metals. 7

8. A process as defined in claim 5 wherein the sulfur concentration during said alternate (A) is maintained at above about 0.2 millimole per mole of hydrogen and during said alternate (B) at below about 0.01 millimole per mole of hydrogen.

9. A process as defined in claim 5 wherein said hydrogenating treatment is catalytic hydrofining.

10. A process as defined in claim 5 wherein said bydrogenating treatment is catalytic hydrocracking.

11. A multi-stage hydrocracking process for converting a hydrocarbon feedstock containing aromatic hydro- 12 carbons and boiling above the gasoline range to highoctane gasoline and a desired proportion, from 0% to about by volume, of a non-aromatic product fraction boiling in the jet fuel-diesel fuel range, which comprises:

(A) subjecting said feedstock plus added hydrogen to hydrocracking in a first hydrocracking zone in contact with a hydrocracking catalyst comprising a hydrogenating metal sulfide distributed on a solid cracking base;

(B) maintaining in said first hydrocracking zone a pressure betweenabout 400 and 2,500 p.s.i.g. and

a temperature adjusted to give a converison per pass to 400 F. end-point gasoline of about 20-60% by volume;

(C) fractionating effiuent from said first hydrocrack' ing zone to recover high-octane gasoline and unconverted oil;

(D) subjecting said unconverted oil plus added hydrogen to hydrocracking in a second hydrocracking zone in contact with a hydrocracking catalyst comprising a group VIII noble metal-containing hydro genating component distributed on a solid cracking case;

(B) maintaining in said second hydrocracking zone a pressure between about 400 and 2,500 p.s.i.g., and a temperature between about 400-750 F. and adjusted to give a substantial conversion per pass to lower-boiling hydrocarbons;

(F) treating the efiluent from said second hydrocracking zone to recover at least one product selected from the class consisting of (a) an aromatic gasoline, (b) a non-aromatic jet fuel and (c) a non-aromatic diesel fuel; and

(G) maintaining in said second hydrocracking zone a concentration of hydrogen sulfide which is (a) relatively low, less than about 0,2 millimole per mole of hydrogen when the product recovered in step (F) is selected mainly from the jet fuel-diesel fuel class, and (b) relatively high, greater than about 0.01 millimole per mole of hydrogen when the product recovered in step (F) is mainly gasoline.

12. A process as defined in claim 11 wherein said feedstock is first subjected to catalytic hydrofining, and the total efihlent from said catalytic hydrofining is then subjected to said first hydrocracking step (A). v 13. A process as defined in claim 11 wherein the catalyst in step (D) comprises palladium distributed on a zeolitic alumino silicate molecular sieve of the Y crystal type containing zeolitic cations from the class consisting of hydrogen and divalent metals.

14. A hydrocracking process for converting a substantially nitrogen-free hydrocarbon feedstock containing aromatic hydrocarbons and boiling above the gasoline range to high-octane gasoline, which comprises subjecting said feedstock plus added hydrogen to catalytic hydrocracking at a pressure between about 400 and 2,500 p.s.i.g., and a temperature between about 500 F. and 750 F. in contact with a group VIII noble metal-containing hydrocracking catalyst, and maintaining during said contacting a concentration of hydrogen sulfide greater than about 0.01 millimole per mole of hydrogen.

15. A process as definedin claim 14 wherein said group VIII noble metal is palladium.

16. A process as defined in claim 14 wherein said hydrocracking catalyst comprises a zeolitic, alumino-silicate molecular sieve of the Y crystal type containing zeolitic cations from the class consisting of hydrogen and divalent metals.

17. A hydrocracking process for converting a substantially nitrogen-free hydrocarbon feedstock containing aromatic hydrocarbons and boiling above the gasoline range to a non-aromatic product selected from the class consisting of jet fuel and diesel fuel, which comprises subjecting said feedstock plus added hydrogen to catalytic hydrocracking at a pressure between about 400 and 2,500 p.s.i.g., and a temperature between about 400 and 750 F., in contact with a group VIII noble metal-v containing hydrocracking catalyst, and maintaining during said contacting a concentration of hydrogen sulfide less than about 0.2 millim01e per mole of hydrogen.

18. A process as defined in claim 17 wherein said group VIII noble metal is palladium.

19. A process as defined in claim 17 wherein said hydrocrackingcatalyst comprises a zeolitic, alumino-silicate cations from 1 the' class consisting of hydrogen and divalent metals.

References Cited in the file of this patent molecular sieve of the Y crystal type containing zeolitic.

UNITED STATES PATENTS Bannerot July 22, 1952 Seubold -Q May 9, 1961 Hansford et a1. Nov. 14, 1961 Ciapetta et a1. J an. 2, 1962 Watkins Mar. 20, 1962 Disclaimer 3,132,090.Paal F. Helfrey, Montebello, Nicholas L. Kay, Fullerton, Bernal Peralta, Anaheim, and Oloycl P. Reeg, Orange, Calif. HYDRO- CRACKING PROCESS WITH REGULATION OF THE ARO- MATIC CONTENT OF THE PRODUCT. Patent dated May 5, 1964;. Disclaimer filed Aug. 13, 1969, by the assignee, Union Oil Company of California.

Hereby enters this disclaimer to claims 14 and 15 of said patent.

[Oyfiee'al Gazette Septembea" 25, 1969.]

Claims (1)

1. IN A CATALYTIC HYDROCRACKING PROCESS WHEREIN A HYDROCARBON FEEDSTOCK BOILING ABOVE THE GASOLINE RANGE AND CONTAINING AROMATIC HYDROCARBONS IS CONTACTED WITH A GROUP VIII NOBLE METAL-CONTAINING HYDROCRACKING CATALYST IN THE PRESENCE OF ADDED HYDROGENAT A PRESSURE BETWEEN ABOUT 400 AND 2,500 P.S.I.G. AND AT A TEMPERATURE BETWEEN ABOUT 400-750*F. SELECTED TO GIVE A SUBSTANTIAL CONVERSION TO AT LEAST ONE DESIRED PRODUCT SELECTED FROM THE CLASS CONSISTING OF GASOLINE, JET FUEL AND DIESEL FUEL, THE IMPROVED METHOD FOR CONTROLLING AND ALTERNATELY VARYING THE AROMATICITY OF SAID DESIRED PRODUCT WHICH COMPRISES: MAINTAINING ALTERNATELY (A) A RELATIVELY HIGH, CONTINUOUS CONCENTRATION OF HYDROGEN SULFIDE, ABOVE ABOUT 0.01 MILLIMOLE THEREOF PER MOLE OF HYDROGEN IN THE HYDROCRACKING ZONE, TO PRODUCE A RELATIVELY AROMATIC PRODUCT, AND (B) A RELATIVELY LOW CONTINOUS PARTIAL PRESSURE OF HYDROGEN SULFIDE, BELOW ABOUT 0.2 MILLIMOLE THEREOF PER MOLE OF HYDROGEN IN THE HYDROCRACKING ZONE, TO PRODUCE A RELATIVELY NON-AROMATIC PRODUCT.

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