US3617486A

US3617486A - Hydrocrackfining of hydrocarbon fractions over mixed metal catalysts

Info

Publication number: US3617486A
Application number: US879941A
Authority: US
Inventors: William R Lewis; Ralph B Mason; Glen P Hamner
Original assignee: Exxon Research and Engineering Co
Current assignee: ExxonMobil Technology and Engineering Co
Priority date: 1969-11-25
Filing date: 1969-11-25
Publication date: 1971-11-02
Anticipated expiration: 1988-11-02

Abstract

A process for upgrading low octane heavy naphtha, such as a 350*-430* F. cut from a conventional hydrocracking operation, or normally virgin crude fractions used to prepare premium jet fuel and middle distillate, as well as mixtures of the cracked and virgin fractions. By treating such hydrocarbons, e.g., the 350*430* F. cut, alone or in admixture with other hydrocarbons, with a dual functional catalyst composition comprising a metal hydrogenation component composited with a crystalline aluminosilicate zeolite cracking component, product quality is upgraded. The conditions utilized in the process are less severe than that normally utilized in a hydrocracking operation but more severe than generally employed in hydrocracking operations so as to result in the conversion of the 350*-430* F. hydrocrackate fraction into a high-octane number C5- 350* F. naphtha. This C5350* F. naphtha product is amenable to further improvement by powerforming and/or special recycle. Catalysts which are of special interest in the practice of the ''''hydrocrackfining process'''' comprise the mixed nonnoble metal catalysts supported on crystalline aluminosilicate zeolites having a silica-toalumina mole ratio greater than 3.

Description

United States Patent 72] Inventors William R. Lewis 3,392,106 7719'68 Mason etal. 208/59 Baton Rouge;

Primar Exammer-Delbert E. Gantz Ralph Mason Denham springs; Glen Assistai zt Examiner-G. E. Schmitkons Hamner Baton Rouge Attorneys Pearlman and Stahl and Llewellyn A. Proctor [21] Appl. No. 879,941 [22] Filed Nov. 25, 1969 [45] Patented Nov. 2, 1971 [73] Asslgnee gi r fizfgzg f z? frfiz zzzg gz ggz 2 ABSTRACT: A process for upgrading low octane heavy 640 588 Ma 23p1967 :3 abandone'd naphtha, such as a 350430 F. cut from a conventional confinugflolzim cation Ser & hydrocracking operation, or normally virgin crude fractions 785 115 Dec 1968 :3 abandonea used to prepare premium jet fuel and middle distillate, as well as mixtures of the cracked and virgin fractions. By treating such hydrocarbons, e.g., the 350-430 F. cut, alone or in admixture with other hydrocarbons, with a dual functional catalyst composition comprising a metal hydrogenation com- [54] HYDROCRACKFINING 0F HYDROCARBON ponent composited with a crystalline aluminosilicate zeolite FRACTIONS OVER MlXED METAL CATALYSTS crackrng component, product quality is upgraded. The condi- 26 Claims 1 Drawing trons utilized in the process are lesssevere than that normally utilized in a hydrocracking operation but more severe than [52] US. Cl 208/59, generally employed i hydrocracking operations so as to resu 208/111' 208/136 208/138, 208/141 252/455 Z in the conversion of the 350-430 F. hydrocrackate fraction [51] Int. Cl ..C0lb 33/28, into a highmctane number 350 naphtha This 350= 8 13/04vCl08 37/00 F. naphtha product is amenable to further improvement by [50] Field oi Search 208/59 -f rmi and/or special recycle. Catalysts which are Of special interest in the practice of the hydrocrackfining [56] Reierences Cited process" comprise the mixed nonnoble metal catalysts sup- UNITED STATES PATENTS ported on crystalline aluminosilicate zeolites having a silica- 3,239,447 3/1966 Reeg et al 208/59 to-alumina mole ratio greater than 3.

4 44 3 26 7 a; 5 AKE-UP H 4 F RACTIONATOR FEED FRACTIONATOR 2 a a Q45 2; 28

i. Q f l9 HYDROCRACK-FINER HYDROFINER HYDROCRACKER HYDROCRACK HYDROCRACKFINING F HYDROCARBON FRACTIONS OVER MIXED METAL CATALYSTS This is a continuation-in-part of applications Ser. No. 640,588 filed May 23, 1967, and Ser. No. 785,1 l5 filed Dec. 19, 1968, both now abandoned.

Multistage hydrocracking with prehydrofining of the feed stream is now well known to the art. Such a process is described in US. Pat. No. 3,306,839, issued Feb. 23, 1967, Raoul P. Vaell, inventor.

In the process of this patent, the product stream from the second stage hydrocracker is fractionated into a C -C light gasoline fraction, a C -400 F. gasoline fraction and a bottoms fraction which is recycled to a second-stage hydrocracker. The octane number of the C -400 F. gasoline is seen to vary from about 85 to about 95 (+3 ml. TEL), depending on the operating conditions selected in the second-stage hydrocraclter. It is believed, however, that the octane value of this gasoline fraction is substantially lower than would be theoretically obtainable due to the fact that the heavy ends of the C 400" F. gasoline product are composed of a very low octane number fraction. The inclusion of this heavy end fraction into the gasoline product results in an octane debit in the gasoline pool.

It is the object of the present invention to separate the lowoctane heavy naphtha fractions, such as the 350-430 F. cut, from the product stream of the conventional multistage hydrocracking process. This material is then upgraded to a high-octane naphtha by treating it in a hydrocrackfining zone at conditions less severe than in normal hydrocracking operations but more severe than in conventional hydrotreating operations so as to convert this fraction in high yield to a C5-3 50 F. naphtha of improved octane.

It is another object of this invention to blend normally virgin crude fractions suitable for the preparation of jet fuels and middle distillates with the 350-430 F. cracked fraction. Such virgin fractions require hydrofining for the removal of sulfur and nitrogen components that contribute to poor color stability and sediment. Such impurities are easily removed by processing them conjointly with the hydrocracked fraction.

It is a further object of this invention to process these virgin fractions alone in the absence of the hydrocracked fraction. Thus the hydrocrackfining operation acts to hydrofine these fractions under conditions somewhat more severe than are used in hydrofining but less severe than hydrocracking. The present invention is based on the unexpected discovery that the lowoctane number heavy naphtha product, such as the 350430 F. cut from a conventional hydrocracking operation, may be converted to a high octane number C -,350 F. naphtha in high yield by treating the 350-430 F. material at novel conditions over a dual functional catalyst having both a hydrogenation and a cracking functionality. The utilization of novel process conditions and singular dual functional catalysts has given rise to a new process which has been given the descriptive name of hydrocrackfining process."

The feed stream utilized in the hydrocrackfining process of the present invention may comprise the low-octane number 350-430 F. product cut from a commercial hydrocracking operation.

According to another embodiment of this invention, lowgrade, normally virgin crude fractions used to prepare premium jet fuel and middle distillate are added in the proportion of from 10 to 80 percent virgin fraction to from 20 to 90 percent of the hydrocracked fraction to the 350-430 F. fraction from hydrocracking in order to remove nitrogen and sulfur components that contribute to poor color stability and sediment. Alternatively, the virgin fractions may be the sole feed to the hydrocrackfining reactor.

The catalysts utilized in the practice of the present invention may be a bifunctional catalyst comprising a metal hydrogenation component and a crystalline aluminosilicate zeolite cracking component such as has been conventionally used in hydrocracking. However, in order to achieve maximum yields and maximum selectivity to desired high octane components, it is preferred that the catalyst utilized in the practice of the present invention comprises a mixed nonnoble metal component supported on a crystalline aluminosilicate zeolite having a silica-to-alumina mole ratio greater than 3. However, when it is desired to maximize the production of jet fuel and middle distillate of improved color stability, it is preferable to use a Group Vlll noble metal catalyst, such as a metal of the platinum series on a zeolite such as steam hydrogen Y-faujasite.

The mixed nonnoble metal catalyst utilized in one embodiment of the present invention contains as a first component a metal cation selected from Group ll-B and the nonnoble constituents of Group Vlll of the Periodic Table (Handbook of Chemistry and Physics, 38th Edition, Chemical Rubber Publishing Co.). Particularly preferred metal cations used as the first component in the mixed nonnoble metal catalyst include zinc, cobalt and nickel. When noble metals are used, these comprise platinum, palladium, iridium, rhodium, ruthenium, and osmium. The second component of the mixed metal modification of the crystalline aluminosilicate molecular sieve zeolite is a metal or metal compound selected from Groups lV-A, V-B or Vl-B metal or metal compound deposited on said sieve by wet impregnation or other processes known to the art. Particularly preferred for use in the present invention are the following metal modifications of a crystalline aluminosilicate molecular sieve zeolite: zinc molybdate faujasite, zinc tungstate faujasite, zinc stannate faujasite, nickel molybdate faujasite, nickel tungstate faujasite, cobalt molybdate faujasite and cobalt tungstate faujasite, as well as platinum or palladium hydrogen Y-faujasite. Other crystalline aluminosilicate zeolites having a silica-to-alumina mole ratio greater than 3 such as those having the mordenite crystal structure are also desired as molecular sieve bases for the mixed nonnoble metal catalysts used in the present invention.

In a preferred method of preparation, a crystalline aluminosilicate molecular sieve zeolite is cation-exchanged with a Group "-13 cation or a noble or nonnoble metal member of Group VIII in the cationic form. The resulting partially exchanged sieve is then contacted with either an ammoniacal or alkaline solution of the Group lVA, V-B or Vl-B acidic oxides. in a particularly preferred embodiment, the metal cation form of the molecular sieve is contacted with an ammonium polysulfide solution of the sulfides of the Group lV-A, V-B or Vl-B metals. As a less preferred embodiment, it is possible to incorporate the Group lV-A, V-B or VI-B metal into the molecular sieve first, followed by cation exchange with the metal cation from Group "-3 or Vlll. The crystalline aluminosilicate zeolite molecular sieves utilized as the catalyst base are commercially available materials. For example, the crystalline aluminosilicate zeolite having a crystalline structure of faujasite and a silica-to-alumina mole ratio greater than 3 is available from Linde Division of Union Carbide, New York, New York, and is identified as Zeolite Y. Similarly, mordenite, preferably in the synthetic form, is available from the Norton Company of Worcester, Mass.

It is further desirable to utilize as catalyst bases crystalline aluminosilicate zeolites which have been treated so as to substantially increase their silica-to-alumina mole ratios over that normally obtained in their synthesis. A suitable procedure for preparing such materials is disclosed in Ser. No. 552,91 1, filed May 25, 1966, by Eberly et al., and that disclosure is incorporated herein by reference.

It is desirable to introduce the metal cation component into the molecular sieve so as to replace 60 to 98 percent, preferably 70 to percent, by weight of the alkali (sodium) ions originally present. Similarly, it is desirable that from about 3 to about 25 weight percent, based on the molecular sieve, of the Group lV-A, V-B or Vl-B metal anion be introduced into the molecular sieve, preferably from about 5 to 15 weight percent. The alkali or alkaline metal content of the final mixed cation containing molecular sieve should be less than about 10 weight percent, preferably less than about 5 weight percent.

The process conditions for the hydrocrackfining process of the present invention are summarized below in table I.

TABLE I Hydrocrackfining Process Conditions In general, the above process conditions will result in a conversion of only from 5 to 7.5 percent of the feed-heavy naphtha. It is essential that the temperature, pressure, feed rate and gas rate be adjusted to each other so that not more than l0 percent conversion to light products occurs. The resulting product obtained from the hydrocrackfining operation may be utilized as a single gasoline component or may be further fractionated into light naphtha (C -180 F.), powerformer feed (l80-350 F.) and middle distillate (300-550 F.) fractions. The light gasoline fractions obtained from the hydrocrackfining operation exhibit improved research and motor octane levels, and superior volatility characteristics over the feed-heavy naphtha. The octane improvement observed is attributed to:

l. the formation of a higher octane C,250 F. fraction,

2. the retention of single ring aromatics, and

3. effective olefin saturation and sulfur/nitrogen removal.

'The production of a l8035'0 F. naphtha fraction is amenable to further upgrading by hydroforming. Where less gasoline is required and less jet fuel is desired, the reforming feed end point can be reduced to 300 F. The 300-550 P. fraction produced by hydrocrackfining operation makes is feasible to make a satisfactory jet fuel. through improved denitrogenation and desulfurization of this fraction.

Reference is now made to the attached FIGURE which is a flowsheet illustrating a preferred embodiment of the invention in conjunction with a conventional hydrofining-hydrocracking operation. lnthe 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.

The initialfeedstock is introduced into the process via line 2. Typical feedstocks include coker distillate gas oils, .cycle oils derived from catalytic or thermal cracking operations, aromatic straight run gas oils, etc., any of which may be derived from petroleum crude oils, shale oils, tar sand oils, or the like. The feed stream is mixed with recycle and makeup hydrogen from line 4 (makeup hydrogen being introduced via line 44). The feedstock and mixed hydrogen are introduced into hydroflner 6 by means of line 3. Hydrofining conditions utilized in hydrofmer 6 are conventional conditions known to the art and may, for example, include a temperature in the range from about 600 to about 850 F., a pressure in the range of from about 500 to 3,000 p.s.i.g., a space velocity of 0.3 to V/V/Hr. and-hydrogen-to-oil ratio of about 500 to 20,000 SCF/B. Catalysts which may be utilized in hydrofiner 6 include 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. The hydroflning operation may be conducted adiabatically or isothermally.

The total hydrofined product from hydrofiner 6 is withdrawn via line 7 and introduced into first-stage hydrocracker 10 by means of line 8. Additional hydrogen may be mixed with the hydrofiner product via line 9.

The catalysts employed in the first hydrocracker 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 inorganic oxides such as silica-alumina, silica-magnesia, silica-zirconia, etc. Acidic metal phosphates such as aluminum phosphate may also be used. The especially preferred cracking bases comprise partially dehydrated, zeolitic, crystalline molecular sieves, e.g.. of the X" or Y crystal types having relatively uniform pore diameters of about 8 to 14 A. and comprising silica, alumina and one or more exchangeable zeolitic cations. An especially preferred zeolite has the crystal structure of faujasite and has a silica-to-alumina mole ratio greater than about 3.

The foregoing cracking bases are compounded, as by impregnation, with from about 0.05 to 25 percent (based on free metal) of a Group Vl-B or Group VIII metal promoter, e.g., an oxide or sulfide of chromium, tungsten, cobalt, nickel, or the corresponding free metals, or any combination thereof. Preferably, the Group VIII noble metals are employed, in amounts between about 0.05 percent and 2 percent by weight e.g., platinum, palladium, rhodium, or iridium.

The process conditions in the first-stage hydrocracker 10 are suitably adjusted so as to provide about 2060 percent conversion to gasoline per pass, while at the same time permitting relatively long runs between regenerations, e.g., from about 4 to 18 months. The specific selection of operating conditions depends largely on the nature of the feedstock, pres sures in the high range normally being used for highly aromatic feeds, or feeds with high end points. Conventional operating conditions now known to the art may be utilized in the first-stage hydrocracker.

The effluent from hydrocracker 10 is withdrawn via line 11 and introduced into separator and scrubber 12. Recycle gasses are collected and returned to the hydrogen line via line 13. Vessel 12 may be equipped with a water inlet and mixing line wherein the gases produced from hydrocracker 10 are scrubbed so as to prevent the excessive buildup of ammonium and hydrogen sulfide gases in the recycle system. The liquid hydrocarbon phase in vessel 12 is removed by line 14 and the liquid hydrocrackate is introduced into fractionator 22 by means of line 21. The liquid stream introduced into fractionator 22 by means of line 21. The liquid stream introduced into fractionator 22 will also contain recycle streams obtained from the second-stage hydrocracker via line 20 and the thirdstage hydrocrackfiner via line 42. Light ends are taken overhead from fractionator 22 through line 23. A C -350 F. naphtha cut is taken as a first side stream through line 24. This fraction has a leaded research octane number in the range of about to and is therefore a prime gasoline blending fraction.

A second side stream cut is obtained in the 350-430 F. range. This material is collected in line 25. The 430 F.+ bottoms are recycled through line 45 to either the second-stage hydrocracker 17 by means of line 15 or a portion thereof may be recycled back through hydrofiner 6 through line 5.

The bottoms from fractionator 22 are fed into the second stage hydrocracker 17 by means of line 16 where they are mixed with recycle hydrogen from line 43 and any makeup hydrogen that may be necessary which enters through line 46. The catalysts used in hydrocracker 17 may be any of the conventional hydrocracking catalysts previously described. Second-stage hydrocracking conditions are also well known in the art. See, in this regard, U.S. Pat. No. 3,306,839 at column 7.

Process conditions in the second-stage hydrocracker 17 are adjusted to provide about 80-40 percent conversion, based on feed. The effluent from the secondstage hydrocracker is collected in line 18 and fed into separator-scrubber l9. Recycle hydrogen is collected and returned to this second-stage hydrocracker through line 43. As was the case for the previous separator-scrubber 12, vessel 19 may be equipped with a water-wash system so as to control the amount of H 5 and/or ammonia present in the recycle gas system. The liquid hydrocarbon product is removed from vessel 19 by means of line 20 and fed into fractionator 22 by means of lines 14 and 21.

The critical aspect of the present invention is the collection of a low-octane heavy naphtha cut in fractionator 22 and using this material as the feed stream for a third-stage reactor, i.e, hydrocrackfiner 29. The heavy naphtha cut, such as the 350-430 F. fraction, may be mixed with additional heavy naphtha material from line 26 and the resulting mixture is fed into hydrocrackfiner 29 through line 27 with additional hydrogen makeup gas which enters via line 28. The conditions for operation of the hydrocrackfiner have been given above in table I. The catalysts utilized in the hydrocrackfiner preferably comprise a mixed nonnoble metal component supported on a hydrogen form of faujasite. The total hydrocrackfinate product is collected in line 30 and fed into separator-scrubber vessel 40. Recycle hydrogen is returned to hydrocrackfiner circuit by means ofline 41. The liquid hydrocarbon product is collected in line 42 and a portion returned to fractionator 22 by means of line 21. Another portion, or even all, is taken by line 46 and passed to fractionator 47 from which products boiling 320 to 550 F. are removed through line 49 and a bottoms fraction boiling 550 F.+ is removed through line 50 which may be recycled to fractionator 22 by line 51 or withdrawn through line 52.

It should be further noted that the hydrocrackfiner catalysts used in reactor 29 generally require sulfiding to enable them to attain a high level of catalytic activity and desirable levels of selectivity. This can be most readily accomplished by pretreating the catalyst with a compound containing sulfur, e.g., H 8 prior to use. Sulfiding can also be accomplished by running the catalyst on a high sulfur feed prior to using it in the hydrocrackfining operation.

The specific embodiments of the present invention may be more clearly understood by reference to the following examples.

EXAMPLE 1 This example demonstrates the preparation of a representative type of nonnoble metal catalyst usable in the hydrocrackfining process.

A charge of 100 grams of molybdic acid was dissolved in 1,000 grams of concentrated ammonium hydroxide solution at 150 F. and the solution was completed by stirring for 2 hours. Thereupon, the solution was evaporated until faintly ammoniacal, reducing the solution volume to 500 ml. To this concentrate, 1,000 grams of water were added and, while stirring at room temperature, nickel faujasite prepared by nickel ion exchange of ammonium faujasite was added in small increments. The suspension was stirred for 24 hours at room temperature and then was filtered, dried and pelleted for use. The final catalyst composition contained 4.6 weight percent nickel and ll.2 weight percent molybdenum trioxide. The faujasite used in this example had a silica-to-alumina mole ratio of 4.8:].

EXAMPLE 2 This example demonstrates the advantages obtained by utilizing hydrocrackfining to upgrade heavy naphthas produced by a hydrocracking operation. The nickel molybdate faujasite catalyst prepared by the procedure of example 1 was utilized to hydrocrackfine a 350-430 F. cut from conventional hydrocracking. The catalyst was sulfided prior to use. The data obtained from the hydrocrackfining operation is summarized below in table ll.

TABLE ll HYDROCRACKFINING OF 350430 F. HYDROCRACKATE RECYCLE OVER NICKEL MOLYBDATE FAUJASITE CATALYST Process Conditions Average Temperature. F. 650 Pressure. p.s.i.g. L000 Feed Rate, V/V/Hr. l Exit Hydrogen Rate, CF/B 2,000

Hydrocarbon Distribution Feed Product C, & L. Wt. 3.3 i-C,, Vol. l3.2 n-C Vol. 3.3 C,-375 F., Vol. 96 50 67.3 375 F.+Vol. 96 50 23.5

The comparative quality of the naphtha portions of the hydrocrackfined product are clearly evident from table III which compares the hydrocrackfinate naphtha product with the original hydrocrackate recycle stock.

It is clear from inspection of table II] that the hydrocrackfinate naphtha product has a lower boiling range and has a higher octane number based on either the research or motor octane number test than the corresponding 350-43 0 F. portion of the recycle hydrocracker product. Of additional interest in the above-mentioned table is the fact that the conversion to a lower boiling naphtha of higher octane value has been made with very little material loss. For example, in the above run a total of 97 percent recovery has been achieved.

EXAMPLE 3 This example demonstrates that the hydrocrackfinate product is a superior powerformer feed. For the purpose of this demonstration, a l70-375 F. cut from the hydrocrackfining product was utilized. This material, to be utilized in a powerformer, was compared to the 350-430 F. hydrocrackate fraction utilized as the feed to the hydrocrackfining process. The results are contained in table lV below.

The data in table lV indicate that powerforming the l-375 F. cut from the hydrocrackfining operation results in a higher yield of the naphtha product having lower paraffin content, greater aromatic content and lower condensed naphthene content. Production of C hydrocarbons with isobutane predominating and C -l30" F. naphtha are added benefits. The above results clearly indicate that the hydrocrackfining process provides a greater quantity of a higher quality powerformer feed than is obtainable through conventional hydrocracking.

EXAMPLE 4 This example demonstrates the results obtained from hydrocrackfining a heavy cracked naphtha using various catalysts. The feed consisted of a blend of 60 percent heavy catalytic naphtha, percent heavy coker naphtha and 25 15 percent heavy steam cracked naphtha. Once-through hydrocrackfining runs were made at treating temperatures of 650675 F., 500 to l,000 p.s.i.g., feed rates of l to 2 volumes of oil per volume of catalyst per hour, and hydrogen gas rates of 2,000-4,000 SCF per barrel. Feedstock inspections,

product distribution and detailed inspections of the various fractions are given in table V (Page 18) for runs made using 0.5 weight percent Pd-Mg-hydrogen faujasite, an 8.8 weight percent Zn, ".6 weight percent Mo on hydrogen faujasite, a

5.8 weight percent Ni, 9.9 weight percent W on hydrogen faujasite, and a 6.3 weight percent Co, 13.] weight percent Mo on hydrogen faujasite, respectively.

These data show that'the leaded research octane and motor octane are improved approximately 4 and 5 units, respectively, at 91 and 92 volume percent naphtha yield on feed.

tions, these being tabulated with the results shown in table Vl below.

TABLE VT.HYDROCRACKING OF HEAVY NAPHTHA- KEROSENE FEEDS [Pd-H-faujasite catalyst] 1 Mostly Ce hydrocarbons and very minor above 300 F.

When the foregoing results are compared with previous examples, particularly example 4, it is apparent that a hydrocracking step produces drastically different results from a hydrocrackfming step. Thus, hydrocracking of the two different feeds produced tremendous quantities of gas (C and TABLE V.HYDROCRACKFINING BLENDED HEAVY CRACKED NAPHTHAS Catalyst Pd-Mg-H- Zn-Mo-H- Feed faujasite iaujasite Ni-W-H-iamasite Co-Mo-H-faumsite Operating conditions:

Temperature, F 697 703 663 654 6 675 Pressure, p.s.i.g 1,000 1,000 1,000 500 1,000 0 V. v. hr 1 1 1 l 1 2 Product distribution:

0: and liszhter. DercenL. 3.6 3. 3 12.6 e. 2 4.1 2. 5 04, vol. percent 11.7 12. 0 33.0 17. 1 15.4 10. C5-43O F., vol. percent- 91. 4 92. 3 90.0 75. 6 86. 2 89. 2 02. 5 430 F. vol. percent 8.6 1.9 3. 5 3.4 3.5 3.7 Naphtha inspections:

Gravity, API as 4 50.6 48.4 66.7 53.8 54.1 50.3 Octane data:

RON plus 3 cc. TEL 87.2 90.5 91 97.7 94.0 93.8 89- MON plus 3 cc. TEL 77. 8 86.2 86. 2 92.0 87.6 86.8 83. 5 Sulfur. p.p.m 1,000 5 3 18 120 7 Bromine No 27 0. 2 0. 2 0. 2

To show the advantages associated with the step of hydrocrackfining as contrasted with an additional hydrocracking step, the following demonstrations present comparative data showing, on the one hand, the hydrocracking of the heavy naphtha feed of example 4 and, on the other hand, the hydrocracking of a 50-50 weight percent blend of this feed with virgin kerosene over a 0.5 weight percent palladium-on-hydrogen faujasite catalyst, this moderately active catalyst being essentially the equivalent of the Pd-Mghydrogen faujasite catalyst employed in the foregoing example. End-of-run conditions are selected for the demonstra- 'lighter)viz, 5l.4 percent and 73.4 percent, respectively,

Another important advantage of hydrocrackfining is that nitrogen is selectively removed from the feed, preferentially to sulfur in the hydrocrackfining step. This is sharply contrasted with the results obtained with conventional catalysts, even at EXAMPLE The following data show the improvement made in hydrotreating ps.i.g., poor quality kerosene fraction from California crude with the 350 F. fraction from hydrocracking. In this example a sufficient amount of a blend of 75 percent Los Angeles basin and 25 percent San Joaquin light crude was added through line 26 of the drawing to make a 50/50 blend with the 350450 F. cut from fractionator 22. This blend was hydrocrackfined in reactor 29 at a temperature of 477 F., a pressure of 400 p.s.i.g., a gas rate of 8,000 standard cubic feet per barrel of feed and a feed rate of 1 volume per volume of catalyst per hour in the presence of 0.5 weight percent palladium deposited on hydrogen Y-faujasite. The product was fractionated and the following distribution was obtained Product Distribution C: & Lighter. Wt. k 0.4 c,, Vol. '1, 1.9 C,l80 F., Vol. Z; l.9 180-320" F., Vol. 12.0 320-540 F., Vol. I: 79.5 540 F.+, Vol. 96 6.5

The kerosene-middle distillate fraction boiling 320540 F. obtained in the above fractionation had the following inspections compared with those of the original virgin feed.

Distillate (Virgin Kerosene Feed) Gravity, Al-"l at 60 F. 38.6 37.6 N,, p.p.m. L3 43 Sulfur, p.p.m. 58 2,500 Color. Saybolt 30 -l6 Color Hold, 3 Hrs. 26 lO-ll TR EXAMPLE 6 A poor quality C,-750 F. fraction from Wilmington crude was contacted with a 5.8 weight percent Ni, 9.9 weight percent W on steamed hydrogen Y-faujasite at a temperature of 455 F., a pressure of 400 p.s.i.g., a feed rate of 1 V/V/i-ir. and a gas rate of 4,000 standard cubic feet per barrel of feed. The product was fractionated giving the following product distribution.

Product Distribution The various liquid fractions had the following inspections.

Total Liquid (C,+) Feed Gravity, API 358 33.6 N; p.p.m. I8 320 Sulfur, p.p.m. 6,400 7,600

Powerformer Feed from Hydrotreating (l-32 0 F.)

Vol. "1': Yield l6.6 Gravity, "API at 60 F. 53.0 N,p.p.m. 0.6 Sulfur, p.p.m. I30

Composition, Vol.

Aromatics 7.9 Paraffins 23.3 Naphthenes 66.2 Cond. Naphthenes 2.6

.let Fuel-Kerosene from Hydrotreating (320500 F.)

Vol. Yield 36.2 Gravity, APl at 60' F. 37.5 N,, p.p.m. /2 Sulfur. p.p.m. 3,000 Color, Saybolt 25 Color Hold, Saybolt 23 Freeze Point, F. 78 Luminometer No. 40.8

Recycle Bottoms from Hydrotreating (500 F.+)

Vol. '1: Yield 45.9 Gravity, API at 60' F. 28.! N,. p.p.m. 36 Sulfur, Wt. '37 L07 The above data show that process conditioh wliichgive low conversion to light gases over a zeolite catalyst are adequate to improve poor quality distillates to give hydroforming feed and jet fuel having excellent stability.

EXAMPLE 7 A poor quality kerosene was treated alone in accordance with this invention and compared with the results obtained by conventional hydrofining. The following data were obtained.

Catalyst Pd on H-Y- CoMo or NiW iaujasitc, on alumina, experimental design basis Process conditions:

Temperature, F 430 484 675 Pressure 400 400 i, 300 V./v./hr 1 1 2. 5-3. 6 Gus rate, 5. 8,000 8,000 3,600

, Feed Product Liquid product inspections:

Gravity,API at 60 F 39.9 44.2 37.6 40.6 N2. p-p.m 0.2 0. 5 43 Sulfur, p.p.m 670 117 2, 500 Luminometer No 40. 7 48. 1 37. 8 245 Smoke point 19 20 17 21 Aromatics, vol. percent 13 12 14. 5 Color, Saybolt i I 17 1 21 -16 215 Color hold, Saybolt. 1 11 10 Product distribution, wt. percent:

Trace l Fractionation of the liquid product to remove conversion products would give a 28 to 30 Saybolt color and probably a corresponding improvement in color hold quality, also.

The above data show that satisfactory results are obtained at lower temperature and pressure by the process of the present invention as compared to conventional methods of hydrofining, indicated by the figures on the design basis.

Having described the invention, what is claimed is:

1. In a hydrocracking process wherein a heavy hydrocarbon feed is contacted with a' catalyst, in the presence of added hydrogen, at conditions sufficiently severe to convert at least 20-60 weight percent of the said feed, on a single pass basis, to products of lower boiling range, the combination comprising the added step of separating a product distillate fraction boiling in the range of about 350 F. to about 430 F., and upgrading same by contacting the said fraction at hydrocrackfining conditions, in the presence of added hydrogen, with a catalyst composite comprising a metal hydrogenation component and a crystalline aluminosilicate ieolite cracking component, at low severity hydrocrackfining conditions of temperature ranging from about 400 F. to about 750 F. and pressure ranging from about 200 p.s.i.g. to about 1,500 p.s.i.g., sufficient to convert not more than about weight percent of the said 350430 F. distillate fraction to lower boiling products, and recovering a lower boiling naphtha having a higher octane value than said product distillate fraction.

2. The process of claim 1 wherein the 350-430 F. fraction submitted to hydrocrackfining conditions is admixed with a low grade virgin crude fraction normally used to prepare premium jet fuel and middle distillate, the said virgin crude fraction being added in the proportions of from about 10 to about 80 percent.

3. The process of claim 2 wherein the virgin crude fraction I is added in proportions of from about to about 90 percent.

4. The process of claim 1 wherein the conversion of the 350-430 F. distillate fraction at hydrocrackfining conditions ranges no more than from about 5 to about 7.5 weight percent butane and lighter.

5. The process of claim 2 wherein the hydrocrackfinate is further fractionated into a C -,l 80 F. fraction, a l80375 F.

fraction and a 375-430 F. fraction.

6. The process of claim 1 wherein said hydrogenation component of said hydrocrackfining catalyst comprises a mixture of nonnoble metals and said zeolite component, upon which the mixture is supported, has a crystalline structure with a silica-to-alumina mole ratio greater than 3.

7. The process of claim 6 wherein said hydrogenation component of said hydrocracking catalyst comprises two nonnoble metals, the first selected from Groups "-8 and VIII of the periodic table.

8. The process of claim 1 wherein the hydrocrackfining catalyst consists of a first nonnoble metal selected from the group consisting of zinc, nickel and cobalt, and a second nonnoble metal selected from the group consisting of molybdenum, tungsten and tin.

9. The process of claim 1 wherein said hycrocrackfining catalyst is sulfided prior to contact with said feed.

10..The process of claim 1 wherein said hydrocrackfining catalyst comprises a nickel molybdate faujasite zeolite.

11. The process as defined in claim 1 wherein said hydrocrackfining catalyst comprises a zinc molybdate faujasite zeolite.

12. The process of claim 1 wherein said hydrocrackfining catalyst comprises a nickel tungstate faujasite zeolite.

13. The process of claim 1 wherein said hydrocrackfining catalyst comprises a cobalt molybdate faujasite zeolite.

14. The process of claim 1 wherein said hydrocrackflning catalyst comprises a palladium hydrogenation component on hydrogen faujasite.

15. The process of claim 1 wherein the said hydrocracking catalyst comprises a palladium hydrogenation component on steamed hydrogen faujasite.

12 RTThe process of claim l wherein said hydrocrackfining conditions comprise the following ranges:

Temperature, "F. 400-750 Pressure, p.s.i.g. 200-l .500 LHSV V/V/Hr. 0.2-l0 Hydrogen Rate. SCF/B 500 l0,000

17. in a hydrocracking process, the combination of steps comprising contacting a heavy hydrocarbon feed, in the presence of added hydrogen, in a plurality of stages, with a catalyst at hydrocracking conditions at severities sufficient to convert essentially all of the said feed to products of lower boiling range, distilling, and separating the distillate into products including a fraction boiling within the range of from about 350 F. to about 430 F.,

subjecting the said 350-430 F. fraction to low severity hydrocrackfining conditions wherein not more than l0 weight percent of the fraction is converted to butane and lighter products, by contacting said fraction, in the presence of added hydrogen, with a catalyst comprising a metal hydrogenation componentvand a crystalline aluminosilicate zeolite cracking component, at a temperature ranging from about 400 F. to about 750 F. and a pressure ranging from about 200 p.s.i.g. to about 1,500 p.s.i.g., and

recovering a lower boiling naphtha having a higher octane value than said product distillate fraction.

18. The process of claim 17 wherein two hydrocracking stages are employed, the first operated to obtain from about 20-60 weight percent, and the second operated to obtain from about -40 weight percent conversion, based on feed, essentially complete conversion of the heavy hydrocarbon feed being obtained by recycling uncracked feed to extinction.

19. The process of claim 18 wherein not more than from about 5 to about 7.5 weight percent of the hydrocrackfined fraction is converted to butane and lighter products.

20. The process of claim 18 wherein the 350-430 F. fraction submitted to hydrocrackfining conditions is admixed with a low-grade vir in crude fraction normally used to prepare premium et fue and middle distillate, the said virgin crude fraction being added in the proportions of from about ID to about 80 percent.

21. The process of claim 20 wherein the fraction admixed with the 350-430 F. fraction is a heavy cracked naphtha consisting of a blend of about 60 percent heavy catalytic naphtha, about 15 percent heavy coker naphtha and about 25 percent heavy steam cracked naphtha.

22. The process of claim 21 wherein the hydrocrackfining runs are made at treating temperatures of about 650-675 F., about 500-l,000 p.s.i.g., feed rates of about 1-2 volumes of feed per volume of catalyst per hour, and hydrogen gas rates of about 2,000-4,000 SCF per barrel.

23. The process of claim 22 wherein the hydrocrackfining catalyst employed is selected from the group consisting of Pd- H-faujasite and Pd-Mg-H faujasitel 24. The process of claim 22 wherein the hydrocrackiining catalyst employed is selected from the group consisting of Zn- Mo-H-faujasite.

25. The process of claim 22 wherein the hydrocrackfining catalyst employed is selected from the group consisting of Ni- W-H-faujasite.

26. The process of claim 22 wherein the hydrocrackfining catalyst employed is selected from the group consisting of Co- Mo-H-faujasite.

l I l i

Claims

2. The process of claim 1 wherein the 350*-430* F. fraction submitted to hydrocrackfining conditions is admixed with a low grade virgin crude fraction normally used to prepare premium jet fuel and middle distillate, the said virgin crude fraction being added in the proportions of from about 10 to about 80 percent.
3. The process of claim 2 wherein the virgin crude fraction is added in proportions of from about 20 to about 90 percent.
4. The process of claim 1 wherein the conversion of the 350*-430* F. distillate fraction at hydrocrackfining conditions ranges no more than from about 5 to about 7.5 weight percent butane and lighter.
5. The process of claim 2 wherein the hydrocrackfinate is further fractionated into a C5-180* F. fraction, a 180*-375* F. fraction and a 375*-430* F. fraction.
6. The process of claim 1 wherein said hydrogenation component of said hydrocrackfining catalyst comprises a mixture of nonnoble metals and said zeolite component, upon which the mixture is supported, has a crystalline structure with a silica-to-alumina mole ratio greater than 3.
7. The process of claim 6 wherein said hydrogenation component of said hydrocracking catalyst comprises two nonnoble metals, the first selected from Groups II-B and VIII of the periodic table.
8. The process of claim 1 wherein the hydrocrackfining catalyst consists of a first nonnoble metal selected from the group consisting of zinc, nickel and cobalt, and a second nonnoble metal selected from the group consisting of molybdenum, tungsten and tin.
9. The process of claim 1 wherein said hycrocrackfining catalyst is sulfided prior to contact with said feed.
10. The process of claim 1 wherein said hydrocrackfining catalyst comprises a nickel molybdate faujasite zeolite.
11. The process as defined in claim 1 wherein said hydrocrackfining catalyst comprises a zinc molybdate faujasite zeolite.
12. The process of claim 1 wherein said hydrocrackfining catalyst comprises a nickel tungstate faujasite zeolite.
13. The process of claim 1 wherein said hydrocrackfining catalyst comprises a cobalt molybdate faujasite zeolite.
14. The process of claim 1 wherein said hydrocrackfining catalyst comprises a palladium hydrogenation component on hydroGen faujasite.
15. The process of claim 1 wherein the said hydrocracking catalyst comprises a palladium hydrogenation component on steamed hydrogen faujasite.
16. The process of claim 1 wherein said hydrocrackfining conditions comprise the following ranges: Temperature, *F. 400*-750* Pressure, p.s.i.g. 200-1,500LHSV, V/V/Hr. 0.2-10 Hydrogen Rate, SCF/B 500-10,000
17. In a hydrocracking process, the combination of steps comprising contacting a heavy hydrocarbon feed, in the presence of added hydrogen, in a plurality of stages, with a catalyst at hydrocracking conditions at severities sufficient to convert essentially all of the said feed to products of lower boiling range, distilling, and separating the distillate into products including a fraction boiling within the range of from about 350 * F. to about 430* F., subjecting the said 350*-430* F. fraction to low severity hydrocrackfining conditions wherein not more than 10 weight percent of the fraction is converted to butane and lighter products, by contacting said fraction, in the presence of added hydrogen, with a catalyst comprising a metal hydrogenation component and a crystalline aluminosilicate zeolite cracking component, at a temperature ranging from about 400* F. to about 750* F. and a pressure ranging from about 200 p.s.i.g. to about 1,500 p.s.i.g., and recovering a lower boiling naphtha having a higher octane value than said product distillate fraction.
18. The process of claim 17 wherein two hydrocracking stages are employed, the first operated to obtain from about 20-60 weight percent, and the second operated to obtain from about 80-40 weight percent conversion, based on feed, essentially complete conversion of the heavy hydrocarbon feed being obtained by recycling uncracked feed to extinction.
19. The process of claim 18 wherein not more than from about 5 to about 7.5 weight percent of the hydrocrackfined fraction is converted to butane and lighter products.
20. The process of claim 18 wherein the 350*-430* F. fraction submitted to hydrocrackfining conditions is admixed with a low-grade virgin crude fraction normally used to prepare premium jet fuel and middle distillate, the said virgin crude fraction being added in the proportions of from about 10 to about 80 percent.
21. The process of claim 20 wherein the fraction admixed with the 350*-430* F. fraction is a heavy cracked naphtha consisting of a blend of about 60 percent heavy catalytic naphtha, about 15 percent heavy coker naphtha and about 25 percent heavy steam cracked naphtha.
22. The process of claim 21 wherein the hydrocrackfining runs are made at treating temperatures of about 650*-675* F., about 500*-1,000 p.s.i.g., feed rates of about 1-2 volumes of feed per volume of catalyst per hour, and hydrogen gas rates of about 2, 000-4,000 SCF per barrel.
23. The process of claim 22 wherein the hydrocrackfining catalyst employed is selected from the group consisting of Pd-H-faujasite and Pd-Mg-H faujasite.
24. The process of claim 22 wherein the hydrocrackfining catalyst employed is selected from the group consisting of Zn-Mo-H-faujasite.
25. The process of claim 22 wherein the hydrocrackfining catalyst employed is selected from the group consisting of Ni-W-H-faujasite.
26. The process of claim 22 wherein the hydrocrackfining catalyst employed is selected from the group consisting of Co-Mo-H-faujasite.