[1]United States Patent Office Patented Dec. 1, 1964 3,159,564 WTEGRAL HYDROFINING-HYDROCRACKING P-ROCESS Arr@old E. Kelley, Orarge, Raoul P. Vaelf, Los Angeles,and Texas V. Inwood, La Habra, Calif., assignors to Union Oil Coni,?any of CaUforWa, Los Angeles, Cali.., a corporation of California Filed Oct. 20, 1961, Ser. No. 146,64!4 12 Claims. (Cl. 208-59) This invention rel,ates to the catalytie hydrogenacion and/or hydrocracking of hydroca.-bons. to prodi-ice e.g., li3rdroc,,enated aromatics, or lower boiling hydrocarbons, boiling for example in the gasoline or jet fuel range. In one aspect, the invention embraces an "inte,@ral" hydrofining-hydrocracking process, the hydrofining treatment being ;,itegrated in such a manner that the bu-lk of the hydrofiner effluent boiling abov-- the gasoline range can be sent directly, a-.id essentially nitrogenfree, to the hydrocracking zone with a mimmtim of intervening cooling, depressuring, washing reheating and repressurin-. The invention is particularl; conceriled with the treatment of high endpoint feeds@@ocks whl:ch are contamin-aled with organic nitrogen arid/or sulftir compounds. Briefly stated, the invention comprises the following essential steps: (1) the initial feedstock is subjected to catalytic hydrofining with added hydro-,en under conditio.ns such that both a liquid phase and a vapor phase are produced, each com@risiig a substantial portion of the initial feed; (2) the li,,luid phase, after stibstantial hydrofining has taken place, is stripped with hydrogen at essentially hydrofining temperat-,lres and pressures to remove ammoilia, hydrogen sulflide and light hydrocarbons; (3) the vapor phase, ei'RLient from the hydrofinp-r, together with the stripping vapors from step (2), are the-@i cooled sufficiently to condense out a substantial portion of the remaining hydrocarbons; (4) the resulting condensate is stripped with hydroaen at essentially hydrofining pressures to remove ammonia, hydrogen sulfide, and dl:ssolved liglit hydrocarbons; (5) the stripped liqliid phase from step (2) and the strinped condensate froi-n step (4) are then subjected to catalytic hydrogenation or hydrocrackirlc., substantiauy in the @absence of catalyst poisons such as anunonia or hydrogen sulfide; and (6) the gasoline, or other desired hydrocarbon fraction synethesized in the hydrogenation zone, is recovered by condensation and fractionation. According to a preferred embodiment of the invention, at least a portion of the stripp'@ng of liquid phase (step 2) is carried out integally during the hydrofining by providiiig a cou7itercurrent flow of hydrogen iipwardly throtigh the lower portion of the hydrofiner to strip out ammonia a,,id light hydrocarbois from the dowriflowing liquid phase, the resulting strippipg vapors being withdrawii from the hydrofiner at an upper point, along viith the vapor phase hydrofiner effluent. This preferred i3.2ode of operation may be carried oil-t either (1) by admitting feed-plus-hydrogen at the top of the hydrofiner, admitting an additional hydrogen stream at the bottom, and removin- vapor phase effltient at a mid-point, or (2) by admitting feed-plus-hydrogen at a m-@d-point to the hydrofiner, admitting an additionnl hydrogen stream at the b at'Loni, and removing vap6r phase e'Lffuent from the top. In either case the liqiiid phase -ravitates downwardly in the lower portion of the reactor, cou nterci-,rrently to the hydrogen admitted at the bottom. The invention also embraces several other modificatio-@is which will become apparent as the description proceeds. A major difficulty in hydrocracking processes is encountered in connection with feedstocks containing organic nitrogen compounds. These nitrogen compotinds are mostly basic in character, and te-@id to, poi-Son the 2 acidic cracking centers of hydrocracki-.qa catalysts. To overcome this difficulty, the feedstock is often subjected to a preliminary catalytir, hydrofining treatment in order to convert the nitrogen compounds to ammonia, and the product is then condensed and washed to remove the ammonia. The remaining nitrogen-free hydrocarbons are tnen subjected to hydrocracking. The principal objection of this pretreatment process is that the facilities required for condensing, washing, reheatingand repressuring the hydrofiner effluent, iare very expensive. 10 It is also known in catalytic hydrocraelcing that the poisoning effect of nitrogen compoands can be minimized byoperatingathi-herhydrocrackin-temperatures. However, this procedtire is objectionable in the case of feedstocks containing substantial amounts of high-boiling 15 components, boiling for example above about 650' F. Mgh-te nperature hydrocracking of feedstocks contal'r@ing such high-boiling constituents leads to rapid deactivation o.@ the catalyst due to coke formation. 20 In the catalytic hydro.-enation of aromatic hydrocarbons, maqy catalysts, a.,id especially the more active ones such as platinum and other noble m--tals, are sensitive to posoning by siilftw compounds. Noble metal hydrocracking catalysts may also be sulfur-sensitive in some 25 cases. It is therefore desirable in many cases to remove sulfur compo,,lnd&, as well as nitrogen, from hydro.genation or hydrocracking feeds:Locks' This may be achieved quite efficiently by catalytic hydrofining, but subjert to the same disadvmtages previously no'Led, involving expensiv.- intestage treatments to remove hydrogen sLIfide. 30 It is there.Lore, the overall objective of this invention to provide a coirrenient and and ecoqomicd method for removing nitrogen and/or sulfur c6m-oounds from hydrogenation or hydrocracking feedstocks. A more specific 35 objective is to obtain all of the substantial benefits of prehydrofini--qg w,:@thout the normally accruing disadvantage of the expensive condensation, washin.-I reheating and repressuring facilities fequired for separate, or noninte.cral hydrofining. A slill further objective is to pro40 vide hydrofining facilities which will effect LI maximum removal of nitrogen and sulfur compounds without overtreatment of the lower boiling portions, thus miiii@zing the overall volume of liydrofiring catalyst and reactor space required. Other objectives will be apparent froiii 45 the ni6re detailed desciiption which follows. The invention may perhaps be, more readily understood with reference to the accompanying drawing, which is a flowsheet illustrating several modifications. in the succeeding description, i@t will be tinderstood that 50 ,he dranving has been simplified by the omission of certain coiiventio-@ial elenents such as valves, pumps, cornpressors, and the like. Wbere heaters or coolers are indic-ated, it will be understood that these are merely symbolic, alid in actual practice riany of these will be combined iiito banks of heat exchangers and fired heaters, according to ste-,ndard engineering practice. Peferri@n.- more particularly to FIGURE 1, the initial feedstock is broli.-ht in via line 2, im'xed with recycle hydrogen from line 4, preheated to incipient hydrofining 60 lc-mperature ii heater 6, and t'iien passed directly into the to@@) of hydrof@ner 8, , which contains an upper bed of hydi-0ii@nin.- catalyst 10 where mixed-phase hydrofining proceeds L,.nder substancally co-iiventidnal, concurrent downflow conditiolis. The ni-lxed-phase effluent from 6 uppcr hydrofinin,@ catalyst bed 10 emerges into an intersi)zce 1-2, where the initial phase separation takes place. Prod,,lct vapors, comprising vapor phase ofiluelit from bed 8 and vapor-phase stripping effluen@t from lower catalyst bed 1-4, a:re withdrawn via line 16 and t-eated as 70 hereinafter de-Qcrib-,d. The liquid phase portion of effluent from bed 8 gravitates downwardly through lower cat-
[2]alyst bed 14, countercurrently to a stream of hydro@.en admitted at the bottom of the reactor via line 13. In the modifcation il'@ustrated, a small optior@al stripping column 20, containing a series of stripping trays 21, is provided below catalyst bed 14, through whicli the liquid portion of feed percolates countercurrently to tb-e hydrog.-n stream and collegts in the bottom, forming a liquid seal. The purpcse of the stripping coli-,mtl is to assure that arrimonia generated in the lower portion of catalyst bed 14 will also b-- removed. For efficient stripping, the hydrogen admitted via line 18 sho-t,,,Id pref,-,-ably b-- preheated to approximately the ten-iperatur(, prevailing in the reactor, as by means of heater 22. The racio o'L hydrogen to oil in catalyst bed 14 and stripping column 20 preferably rayges between about 500 and 15,000 s.c.f. per barred of oil, at reactor pressures between about 400 and 2,000 p.s.i.g. and temperatures betm7een about 400' and 800' F. The stripped heavy oil accumulatiii- in the bottom of stripp,'P.g co'umn 20 is taken off via line 24 anl vaINIC 26, in response to l@'.quid level controller 28, and is transferred directly to hydrocracker 30, via lines and 34 and heat exchanger 36, as will be stibsequently described. Suitab'ie hydrofining catalysts for use in hydrofiner 8 include for example mixtiires of the oxides a-.id/or silliides of cobalt and molybdcnun-i, or of iiiel,@el and tun.-sten, preferably sup-porte-I on a carrier such as alum;na, or alumina contai,.ling a small amount of coprecipitated silica gel. Other suitable catalysts include in @eneral the oxides and/or siilfides of the Group VIB and/or Grouz) VIII metals, preferebly sudported on substantially noncracking adsorbent ox,.de carriers such as aluiiuna, silica, ti,tania, and the like. The hydrofining operation may be conducted either adiabatically or isotherinally, and under the following general copditions: IIYDROFINILNG CONDITIONS Operative Pref@rred Temperature, I F ------------------------- 600-850 Prossul@e, p.s.i.g --------------------------- 400-3, 000 80-@2, 000 LHSV, v./v.fhi ---------------------------- 0. 5- 10 1-5 H2/oil ratio, s.c.f./b ------------------------ @@n,0-15,0.oo 1, ooo-lo, ooo The above conditions are suitably adjtisted so as to reduce the nitrogen content of the feed to below abou, 25 parts per million, and preferably below about 10. parts per millioi. The vapor phase which istalceii off from liydrofiner 9 via line 16 comprises hydrogen saturated witli hydrocarbons of all boilin@ range, and also containin- ammonia and hydrogen@sulfide. The light hydroca@rbons, boiling in and b-.1o,,v ,he gasoli@ie ranl,e, comdris-. niostly hydrocarbon frai,,rnenls derived from the d--composition of sulfur and nitrogen compounds in hydrofiner 8, sijice very little true hydrocracking occtirs therei-,I. It is desired to recover at miniiiium expense, and in an ammoniafree condition, the bulk of the hydrocarbons boiling above the gasoline rang,- for treatment in hydrocracker 30. For ,this purpose, the total vapor phase i-,l line 16 is transferred via cooler 33 to a sinall s--parator-stril)ping column 40 ' At reactor prcssures, it is norrially found that the bu.'k of these heavier hydrocarbons can be condensed out of the vapor phase by cooling the same to about 400'600' F. It is undesirable to cool the vapor any more than is reqtiired to produce a substantial desired amount of liquid condensate; generally it is unnecessary to go below abolat 300' F. It will be understood that stripping column 40 is operated at substantially the same pressure prevailing in the hydrofiner, or within about loo p.s.i.g. thereof. The cooled mixed-phase material entering the top of separator-stripper 40 senarates in the upp-,r portion thereof the vapor phase @oing overhead via line 42, and the li4,iid ,Orfio-n @ravit-,ttin@ downwardly over stripping trays 44, countercurrently to a risiiig stream of preheated 3,159,564 4 hydrogen admitted near the bottom thereof via line 46, thereby stripping light hydrocarbons and ammonia therefrom. Hydrogen rates similar to those p@- escribed for stripper 20 may be employed, and the stripping hydrogen is preheated in heater 48 to about the same temperature as the enteriii.@ mixedphase effluent from cooler 38. The liquid phase accumulat-ing,.n the bottom of strir)pcr 44 Nvill of course contain some high-boiling hydrocarbor@s, btit viill cor@sist primarily of light gas oil boiliig in 1( ',he 40'u' to 600' F. range. This liquid phase is withdr-aivti via I;r-te @O and valve 52 in response to liquid lev--l con-troller 54, and is then transferred via line 56 I to h),drocraclcer 30, where it is stibjected to bydrocrackin.- under ore of two alternate modifications hereinafter 15 describ-.d. The vl@,ror phise overhead from s,,r'.-pper 44 comprises mostly gasol:ne-boiliigraiige hydrocarbons, buL will stil@l conta:l,-i a corisiderable proportior@ of lighl -.,is oil hydrocarbo@is. it is also codtaminated with ammorit 20 Fiid/or hyerogen stilfide. To effect recolery and piriflcit@'on of the v,,iriolis componon',s, the total vapor phpse in lin-. 42 is partially coolf-,d and condr-nsed in cooler 58, then mixed with wash water from line 60, and transferred via final cor@,Jenser 62 to h;,-h Fressui-e separator 64. The 25 p,,,,,rpose of the w-,ish witer is lo remove water-soltible imr@llritA:es, ir@citicling ariiii-io,.Lia -,ind hydrogen sulfide and sal'ts ther,-ol'. T@lori-ii,,Llly about 15 pouncls of viater pcr barrcl of ];quid hydrocarbon is suffcie,.it 'Lor ',his purpose. It i3 preferable to inject the wash water prior to the final 6- 0 coo,in- i-,i condens-.r 62, i-@l order to avoid condensation ol i soi:d salts in t@ie trar@sfer ]4nes, as m;gbt occur if final cooliil "@ to, e.g., 200' F. were' carr,'.ed o'ut before th-- iaject@'on of water. A three-pliase separation '@akes place in high presstire 35 sepai7,Ltor 6,1. Spent wash water coiitai-@iin.- dissolved inipurit"cs is wilhdrawn via line 66, and purif.;ed recycle hye'ro.-en is wi,hc'traw@i via line 63 for reuse in the hydrofiner a.Tid strij)@,)L@ig coumns associated therewith, as previousl@y described. T@@.c.Uid hydrocarbons in se2arator 64 io ai-e withdraw,.i via line 70 and flashed irto lo,@v pressure se-ar,,Ltor 72, from wh,'@el, I;ght hydrocarbon @ases are exhaus@ed via Iiiie 74. The liqu;d bydrocarbon fraction in scriarator 72 is transterred via Iiiie 76 to fractionating cotumi 78, fror@i v,,hioll the d-@sired gasoline product is talen overhead v-@a 80. 45 As will becom@ anp,@irent frorii the sticceedir,,g descrir,- tion, the hydrocarboq fraelion adnlitted io separator 72 v@@a line 70 al@.o ;-@icluciles the hydrocarbo-@i product from Lydrocracker 30. He,.,,ce, 'ractio@iating column 78 serves to reco-,,cr botb the minor p-,-Orortion ol' -asol;ne syn50 thes,'.Zed in tl,@e hydro-fiiicr, and the major porlion syntiiesized i-@i h3,drocrackcr 30. If desircd, a side-cut j-,t fuel or die@.,-l fr,-,ction may be reecvered fro@n th-, column via 1,"iie Th-@ bottoms ftiction from co'@umn 73, compris'ed @ -rom hydrocrackcr ing unconn,ei-t -a., oi'- hydrocarbors f 55 30, a@nd -emaining gas'oil hydrocarbons froin hydrofiler 8, is taken off via line 84 and blended with tiip- liqu;d phase hyd:70finer efflue-@it in lir@e 32, and the ii-i,:xture is tlle-@l transferred via line 3@l ard exch,,in,,Cr 36 to hydroerneke.- "10. Al@err@atively, where stripper 116 is utilized 60 as descri@ed below, it mq-y be preferable 'Lo mix the bottom,@ fraction in I:ne 84 with the stripped condensate in line 56, and clivert'tlie mixture via preheater 110 i-@ito the iiiid-sectioii o4f hydrocrc-,cker 30, as w;ll be describ(-'d below in conn--ct;.on Vi'Lli alternate B - It should be noted 65 fur,@her that the botoms fract@@on from colum@i 73 need not r@eccssarily be recycled to the hydrocracker; if 100% coiive.-sioii ot the feedstock is not desired, aR or a portion of.' the boltom,, fractior,_ in line 94 m,-,y be d,.verted from the ;-3rocess for other uses, e.g., as feed to a ca+lalytic 70 cracking unit. The o,3eration of hydrocracler 20 will no-,v be e-cscribed in reference to the two princ;pal alternate modifications referred to above. A Iteriic!te 4 .-In this modification, the entire feed to the 75 hydrocracl@er eliters the top and flovis downwardly there-
[3]thro,ligh in a ronventional manner. Th,:s is the simplest modification, and may be preferred in cases where the light gas oil feed derived from stripper 40 constittites a very minor portion of the fresh feed to the hydrocracker. To operate in this rjrann,-r, valves 108 ap-d 112 are closed, 5 and the stripped condensate L-i line 56 ;s transferred via open valve C,6 and line 83 to liie 34, Nvhich conta,ns the main feed stream derived from @stripper 20 and reeycle oil line $4. The mixture is blended viith recycle and fresh hydrogen derived j'rom lines 90 -,md 92, and the temllera- 10 ture is s,,iitably adjusted in exehaiiger 36, whirh may either be a heater or a cooler, debending upon the temperature of the feed mixture in line 34, as well as upon the desired hydrocracking temperature, which in turn will depend upon the relative activity of the hydrocrackiig 15 catalyst. NoTmally however since the hydrogen mixed with the feed -'Is relal.;.ve,y cool, some preheating will be required. In any case, the mixture of hydrogen and oil, at suitable inlet hydrocrackiiig temperatures is admitted via line 94 int6 the tGD of hydrocracking reactor 30 and 20 passes downviardly therethrough under suitable hydrocracking conditions here-,nafter pres(@ribed. Hydrocracker 30 may be operated under cilher adiabatic or isothermal conditions, aiid in the latter case, the reactioii being eyothermic, it may be desirable to admit 25 cool quench hydr@ogen at o@ie or more points in the reactor as illustrated via lines 96 and 93. it will be noted that the calalyst in hvdrocracker 30 is divided i-@ito ail tipper bed 100 and a iov;er bed 102, separated by an irterspace 104. This has no pa cular si ance in r 9@O erence to alternate A, and hence i.,i this case a sin uiu 'tary catalyst bed ma3r be Litilized. Effluent from the hydrocrael,er is taken off via lne 106 a--d treated for product recovery as hereidafter described. ,41ter;za,,'e B.-Since the liquid phase hydrofiner ef- 35 fluent.'Lrom strinoer 20 has a cois-@derably higher average moieciilar weight than the stripped conden-sate from stripper 40, it rnay be a@.ivintageous in some cases to take advant-,Lge of the different optimum cracking t(-mperallires for thes@@ two fractions by introducing them seParatelY ',O il@,,o! the reactor in zones vhere opt@'@mum temperatures prevail for the respective frac@'Lion. This, is partictilarly feasible in cases where tl-ie initial feed to the hydroflii-er contains substant;a'l lronortilons, e.g., n-iore than iabout 20% by volume, of hydrocirbons boiliiig in the 400'- 45 10(lo F, range, i,, whch case the stripped conden,ate in li@ne 56 will coniprise asubstantial proportion of the total feed to th-. hydrocracker. Selective treatmelit of th-, tnvo feed frac-ti6ns '@s eifected i-@i this alterna@Le by maintainin.- upper byd-ocracking bed 100 at a,relatively lower aver- '@A age temperature than lower hydrocrackin- bed 102, and fe@d@@ng the hea-%7y fraction of feed into the, top of the reactor while feeding ',he li.-hter fraction into the top of thp- lower hydrocrackipg catalyst bed. The differential temperattires required in this -mode of operat;on can in 5-5 most cases be nial:nt-.ined by allowing hydrocracldng to proceed in reartor 30 under more nearly adiabatic conditions, ivith a ininimum of quenchlng via lines 96 and 98. It ;s cortemplaled ttiat upper hydrocrarking bed IGO niay be maintained ,t an average bed ten-iperalure aoout 20'- 60 100' F. lower than the avera,@e bed tempere@ure in lower hydrocr acking bed 102. Preferably, upper bed 100 is mainwn ed at an average temperature between abolit 500' and 750' F., a-iid lovier bed at 102 at 5@JO'-800' F. To effect selective treatment of the feed fractions in 65 this inanner, valve 06 is closed and valve. 108 is opened, th--reby diverting stripped condensate in line 56 through preheate r 110 into interspace 104, where it mingles with the doivnfiowing effluent from Lipper bed 100. It is also desirable in this case to mix the. light feed with hydrog-.n 70 via line Ill a.,ld open valve 112. By operating in the manner prescribed, upper hydrocracking bed 100 serves primarily to crack the heavy gas oil feed to light gas oil, while lower bed 102 serves priniarily to convert liglit gas oil coniponeiits to gasoline. Maximum catalyst effl- 75 6 ciency, and efficiency of conversion to gasoline is obtained. In either of the above alternate modes of operation, th@- hydrocracker effluent in line 106 is next passed through a partial condenser 114 and thence into an optional separator-stripp-Ir 116, which functions similarly to stripping column 40. The purpose of separator-stripper 116 is to provide for economical recovery of a substantial proportion of u--qconverted oil from the hydrocracker before it is cooled to final condensation temperatures, thus minirnizing the utilities and facilities required for reheating the recycle oil. Normally, the temperature of the hydrocracker effluent is reduced in cooler 114 to about 400'-700o F., and about 500-10,000 s.c.f. of hydrogen per barrel of liquid condensate is admitted near the bottom of the stripper via prelieater 118 and line 120, to effect countercurrent stripping of light hydrocarbons from the condensate. The condensate accumulating in the bottom of stripper 116 coniprises the heaviest components of the unconverted oil, and is more or less cont-'@nii-ously r@- moved via lipe 122 and valve 124, in response to liquid level controller 126, and recycled to the hydrocracker via lines 129 and 34. The vapor phase from separator-stripper 116 is taken ov--rhc-ad via line 130, and comprises the uncondensed vapor phase effluent from thehydrocraclcer as well as the stripping vapors from stripper 116. This combined vapor phase is then transferred via final condenser 130 into high pressure separator 132, from which hydrogenrich recycle gas is taken off via line 134 and relsed in the process as previously described. The condensed hydrocarbons in separator 132 are withdrawn via line 136 and blended with the liquid hydrocarbons in line 70, and the restilting mixture is flashed into low pressure separator 72 and fractionated in column 78 as previously described. As indicated above, the use of stri_t3per-separator 116 is optional, and in raany cases, es@ecially in small installations, may not be economically justifiable. In these cases, it is preferable S:IMPIY to ornit the stripper, and send the hydrocracking effluent in line 106 directly through final condenser 130 ind high pressure separator 132. VAere the stripper is employed however, it will be apparent that tne recycle bottoms fraction recovered from column 78 will have a relatively lower average molecular weight than the recycle condensate from stripper 116. its boiling range may approximate that of the condensate from stripper. 40; her@ce the feasibility of combining these fractions for selective treatment in lower hydrocrac&in.bed 102 as noted above. Many variations are contemplated from the processir)g scheme described above and in the drawing. For exainple, as shown in FIGURE 2, it is entirely feasible to operate hydrofiner, 8 with the total feed plus hydrogen enterin@, lntersi)ace 12 via line 140, the vapor phase passing upwardly through bed 10 and the liquid phase gra@iitating downwardly through lower bed 14, countercurrently to the rising hydrogen stream admitted via line 18. The total vapor phase effluent is then taken off aL the top of the byarofiner via line 142 and sent to separatorstripper 40 v-la cooler 38. Liquid phase product is taken off- as before, via line 2,4. This modification is advantaaeous in providing for a rnore efficient hydrofining of he vapor phase in the absence of liquid phase. t The process conditions in hydrocracker 30 are suitably adjusted so as to provide about 30-80% conversioii to -asoline or other desired product, per pass. It is-Fleferred ai the same time to adjust the operating conditions, prei'erably temnerature, so as to permit r-- Iatively long ruris between regenerations, i.e., froin about 2-8 months. For these purposes, it @will b-- underslood that pressures in the high ranges will be used in connection with temperatures in the high range, while the lower operative pressures will normally be used in colijunction with the
[4]3,159,564 L 9 lower temperatures. The range of operative con-Liitions taining between 5% and 75% of each component; parcontempla'Led for, hydrocracker 30 are as follows: tially dehydrated, zeolitic, crystalline molecular sieves, HYDROCP,ACKING CONDITIONS e.g., of the "X" of "Y" crystal types, havin.- relatively u n i f o r m p o r e d i a m e t e r s o f a b o u t 8 t o 1 4 A n g s t r o m s , a n d comprising silica, alumina and one or more exchan.- eable Opor,, itive, Prefcrred zeolitic cations. Any of the foregoing cracking bases rray be flrther Temperature, I F ------------------------- 400- 850 500-7,-, promoted by the addition of small amounts, e.@., I to 1'ressure, p.s.i.g --------------------------- 400-3, 000 8oo-Z o@'O 1,HSV, v./v./lir --------------------------- 0.5- 10 0.7-5 10% by weight, of halides such as iluorine, boron triH2/oil ratio, s.e.f./b ------------------------ 5oo--15, ooo 1, ooo-lo, 000 I 0 fiuoride or silicon tetrafluoride. The foregoing crackiiig bases are conipounded, as by Those skilled in the art will readily understand thqt when ran,-es of operating conditions are specified cs above, a large number of determinative factors are involved. Thus, hirhly active catalysts, or f@-es!i catalysts 15 at the begiiining of a run, will be used in copjunction with lower teiiperatures than will less active or partially deactivated catalysts. Also, the lower temperature ranges will normally be Lised in conjunction with feedstocks havin- .1@gh end-points, i.e., above about 750' F. The 20 lower limit of pressure to b-. utilized in a given operalti--n will normally depend tipon the desired run leiigtli. LoNver pressures -enerally result in a more r-tp,'@d deactivation of the catalyst, and hence where e,-tremely Iona run 1--n.-ths are desired, pressures of above about 1,000 25 p.s.i.g. are mandatory. 14-owever, econom@'@Cally feasibie ru@i lengths are noi-mally obtainable Nith most catalysts and fe@dstocl.-s wit-iiin the 600-2,000 p.s.i,g. pressure raii,@e. The hydi-o.-rackir.,.- feedstocks which n-iay be treated @@ 0 hereir. include in general any mineral oil i'raction having an initial boiling point above the conventional .-asoline range, i.e., above about 400' F., and having an endboiling-point of up to about 1,000' F. This includes straight-run gas oils, coker distillate gas oils, deasphalted @95 crude oils, cycle oils derived from catalytic or thermal crackitig operations and the like. These fractions may b-- derived from petroleum crude oils, shale oils, tar sand oils, coal hydrogenation products and the like. Speciflcall3r, it is preferred to employ f.-,-dstocks boiling between -10 about 400' and 900' F., having an API gravity of 15' to 35', and containing at least about 20% by volumn of of acid-soluble components (aromatics+olefins). Such oils may also contain from about 0.1% to 5% of sulfur and from about 0.01% to 2% by weight of iiitrogen. 45 Hydrogenatioil feedstocks niay comprise any of the foregoing hydrocracking feeds, and in addition riany other aromitic concentrates as e.g., naphthalene or methyl naphthalene fractions. It should be noted further that the process of this in- j5o vention is especially amenable to the treatment of feedstocks boiling predominantly in the 550'-900' F. rarge. Where the feedstock is composed mainly of hydrocarbons boiling in the 400'-600' F. range, the alternate two-stage hydrocracking process disclosed in co-pe.iding 55 application Serial No. 137,366, filed September 11, 1961, may be more desirable, assumin.- that the size of the op@-ration justifies the capital inventment in a two-stage proc,-ss. The process of this inventioii is primarily directed to the treatmenl of feedstocks which (1) contain 6( a substantial propo-tion of heavy - niaterials boiling above about 650' F., and (2) contain less than about 30% by volume of li.-ht componen'ts in the 400'-600' F. range. The hydrocracking catalyst to be employed in the hydrocracking unit described above may consist of any de65 sired combination of a refractory cracking base with a suitabl@- hydro.--Inatin,@ compone-@it. Suitable cracking bases include for example mixtures of two or more refractory oxides such as silica-alumina, silica-magnesia, silica-zirconia, alumina-boria, silica-titania, silica-zirconia- 70 titania, acid treated clays and the like. Acidic metal phosphates such as alumintim phosphtte r@iay also be used, The preferred cracking bases comprise composites of silica and alumina containing about 50%-90% silica; coprecipitated compos;tes of silica, titania and zirconia con75 impregnation, with from about 0.5% to 25% (based on free metal) of a Group VIB or Group VIII metal pronioter, e.g., an oxide or sulfide of chromium, tungsten, cobilt@ nickel, br the corresponding free metals, or any combination thereof. Alternatively, even smaller proportions, between about 0.05% and 2% of the metals platinum, palladiurn, rhodiuiii or irridium niay be employed. The oxides and sulfides of other transitional metals may also be used, but to less adv-,intage than the foregoing. A particularly suitable class of hydrocracking catalysts is composed of about 75%-95% by weight of a coprecipita',cd base contai-iiing 5%--'i5% S'02 5%-75% ZrO2, and 5%- 75% TiO,, and incorporated therein from aboul 5%-25%, based on free metal, of a Group VIII metal or melial sulfide, e..-., nickel or nickel sul.fide. The molecular sieve type cracking bases, when compounded with a hydrogenating metal, are particularly Lseful for hydrocrackiig at relatively low temperatures of 400-700' F., and relatively low nressures of 400-1,500 p.s.i 'g. It is prefered to eniploy rqolecular sieves having a relatively hi.-h SiO2/A]203 ratio, e..-., betvieen about 2.5 and 6.0. The most active forms are those wherein the exchangeable zeolitic cations are hydrogen and/or a t)olyvalent (preferably divalent) metal such as r@iagnesium, calcium or zinc; In partictilar, the "Y" mol@ecular sieves, wherein the SiO2/Al2O3 ratio is about 5, are preferred, either in their hydrogen form, or a divalent nietal form, prefertbly magnesium. The Y molecular sieves can be prepared by heating an aqueous sodilm alumino-silic@- te mixture at temperatures between about 25' and 125' C. (preferably 80- 125' C.) until crystals are formed, and separating the crystals from t'--,-motherliquor. Whenacoll@o;dalsilicasolisemployed as the source of silica, the aqueous sodiuni alumino-silicate mixture may have a composition as follows, expressed in terms of mole-ratios: Na2O/SiO2 ------------------------------- 0.2-0.8 SiO,/AI,O - ----------- ------------------- 10-30 H20/Na2O ------------------------------- 25-60 When sodium silicate is used as the s,@lica source, the optimum molar proportions are as l'Ollows: Na2O/SiO2 ------------------------------- 0.6-2.0 SiO2/AI203 ------------------------------- 10-30 1-1@,O/Na2O ------------------------------- 30-90 The "decationized " or hydrogen form o' the Y zeolite may be prepared b@ ionexchanging the allkali metal cations with ammonium iods, or other easily deconiposable cations such as methyl subsittuted quaternary ammonium ions, and then heatin@, to e.g., 300'-400' C., to drive off the ammonia. It is preferred that the degree of de, catiopization, or hydrogen e,-,change, be at least about 40% of the maximurn 'cheoretically possible. The final composition should contain less than about 5% by weight of Na2OThe polyvalent metal (e.g., Mg, Ca, Sr, Ba, Zn, Mn, Ni, Cr, etc.) forms of the Y z-@olite are prepared by conventioral cation exchange iiiethods, @as e.g., by treating the sodium zeolite with an aqueous solution of magnesium chloride, calciur-,i chlorid.-, etc. It is preferred to rei)lace at least about 40% of the moiiovalent met,,il cations witli their equivaient of polyvalent catiojis; the activity of th'e catalyst @-enera'(Iy iiiere;asp-s as the degree of displacement of mbnovalent metals increases.
[5]9 Mixed, hydro.-en-polyvalent metal forms of the Y zeolite are also contemplated. Generalily such mixed forins are prepared by subjecting a polyvalent metal-exchanged zeolite to f-urLTier ioti exchange with anunonium cations which @are later driven off during a thermal activation treatment. Here again, it is prefeited that at least about 40 % of the monovalent metal cations be replaced with any desired combination of hydrogen ions and polyvalent metal cations. The hydrogenating promoter may be incorporated into the molecular sieve type cracking bases by any rielhod which gives a suitably ititimate admixture. Among acceptable methods are (1) cation ekehange 7Lising an aqueous solution of a suitable metal salt wherein 'Lhe metal itself forms the cation; (2) cation exchange using an -aqueous solutio-ii of a suitable metal compound in which the nietal is in the forin of a complex cation with coordination coniplexing agents such as ammoriia, followed by thermal decomposition of the catiordc complex; (3) conventional impregnation with an aqueous solution o@,L a suitable metal salt, followed by drying and thermal deconipogtion of the metal compound. The ion-exch@ange methods (1) and (2) are much to be preferred in that a more uniform and complete subdivision of the metal on thp- zeolite is obtained. Method (1) above is generally employed toi introduce metals of the iron group, while method (2) is generally be,st adapted for the noble metals of Group VIII. When me'Lhod (I) is eniployed to i-@itroduce an iron group metal, it is desirable to cairy o,,it subsequent thermal activation treatmen@,s in a nonoxidizing or reducing atmosphere in o@-der to avoid oxidizing the metal ard:displacing it fror@i the zeolite lattice. BuL in the case of the Group VIII noble metals such precautions are generally unnecessary, and thermal decomposition of the cationic coinplex can be carried out in ait if desired. The ion-exchange of hydrogenating metal onto the zeolite may be carried out by the usual methods de-scribed ,Lbove in connection with ti-@e ionexchange of polyvalent metals geiierally. In fact, the two steps may be combined if desired by using an aqueous solution of a niixture of salts, or a siigle salt in cases where the desired polyvalent metal is also the desired hydrogenating rletal. Generally, however, it.is preferable to carry out first the ion-exchan.-e step for forming the hyd@rogen and/or polyvalent metal zeolite, and then perform the ionexchange step providing the hydrogenating metal. 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 12 A., a@-id this is likewise the case in the Y sieves, although the latter usuoy are foi-ind to have crystal pores of about 8 to 10 A. in diameter. Non-cracking hydrogenation catalysts for use herein may comprise any of the previously noted transitional metal hydrogen,tting components, preferably supported on a neutral, adsorbent carrier having an extended surface area of e.g., 50-300 square rneters per gram. Suitable carriers include for example, activated alumina, silica g-- I, activated charcoal, clays and the lilre. Hydrogenati-,ig catalysts comprising a noble metal stich as platinun, palladium or rhodium are e . specially preferrqd since they are highly active and most susceptible to poisoning by sulfur compounds. Hydrogenating conditions are @iii I general similar to the hydrocracking conditions above described ' except that temperatures in the lower range of about 400'-600' F. are generally preferred. The following example is cited to illustrate more co ncretely, exemdlary process conditions and results, but is -@iot intended to'be limiting. Example This example illustrates the results obtainable in a typical operation of the process described in the drawing, with the liquid phase Lrom th-. hydrofiner b--ing treated 10 scriahy in two superimppsed beds of hydrocracking cata'lyst, and the liquid cmdensate from stripper 40 being treated only in the -second of the two beds. The feed is a 400'- 850' F. boiling range coker gas-oil containing 2% of sulfur and 0. 3 % of nitrogen by weight, and having an API gravity of 22'. The hydroiining catalyst is composed of about 4% cobalt sulfide plus 16% of molybdenum sulfide impregnated on a silica-stabilized (5% SiO2) activated aluniina support. The hydrocracking catalyst is a 10 magnesium "Y" molecular sievb containing about 3% by wei.-ht 4of zeolitic magnesium, and loaded with 0.5% by weiaht of pafladium (Linde hydrocracking catalyst MB 5382). The operating conditiolis are as follows: TABLR, I 15 iaydrofiner 8 Hydrocracker 30 O perating Conditions Stripp er U13p er LoNv er 40 U pper LoNv er bed 10 bed 14 b ed bed 20 100 1 02 L HSV ---------------------- 1.5 2. 0 -------- 2.0 2- 0 P ressure 1,550 1,550 1,540 1IB30 1,530 r. P,3,@,t,,u,rse,,,.gl F.,- Start of ruti --------------------- 700 700 500 450 520 nd of rtin ------------------ 750 750 500 670 700 25 1 12/Oil s.c.f./b ---------------- 5,000 3,000 3,000 8,000 8,000 Both beds Convors@'@on to 400 1 F. endpoint g@-soline, vol. percent per pass -------------- -3 -2 -------- 65 30 Average bed teinperatures. On the basis of an operation atiliz@ing 10,000 biarrels per day of initial feed, and with total recyole of unconverted oil to the hydroer,,wker, the approximate yields are as 35 fcllows. TABLE 2 Conversion to gasoline,,tud ligliter ------------------ 100 vol. percent. CI-C3 dry g,,is --------------------------------------- 100 @.c.f./b. feed. Liquid products: Buta,nes ---------------------------------------- t)50 b./d. 40 C5-co ------------------ ------------------------ 2,400 b./d. 07-400' F ----------------- ---------------------- 9,050 b./d. Total liquid products ------------------------------- 12,400 b./d. Chemical H2 consumption -- ------------------------ 2,400 s.c.f./b. feed. - Produ ct ciuality C@- CB C 7-400' F. 45 G ravity, ' API --------- 8 2 49 F esearch Octaue+3 ml. 99 83 Basic nitrogen, p.p.m ------------------------- ------------ 0.2 . In t@he fore,@oing ex I ample, about 30% by volu@me of the 50 total feed to the hydrofiner is recovered as condensate froin separator-stripper 40. If this condensate is blended with the stripped liquid pliase from the hydrofiner, anct the mixture then subjected to hydrocta-@king at 1.0 space velocity, :a-.qd an isothermal bed teinperatiire adjusted to 55 give 65% conversion per pass to gasoline, simil@ar product distribution and yields axe obtained, except for somewhat higher dry gas and butane makes. The run length @is also somewliat decreased as a result of a more rapid rate of catalyst deactivation. 60 Results :analogous to those indicated in the foregoing examlyle are obtained when other hydrocra-@king catalysts and conditions, other feedstoocs and other hydrofinin, conditions within the biroad purview of the,above disclosure are employed. It is hence not intended to limit the 6,5 invention to the details of the example or the drawing, but only broadly as defined in the following claims. We
