US3125502A

US3125502A - scott

Info

Publication number: US3125502A
Authority: US
Publication date: 1964-03-17
Anticipated expiration: 1981-03-17

Description

March 17, 1964 .1.w. sco-r1'. JR., Erm. 3,125,502

STAGED CATALYTIC CONVERSION 0F PETROLEUM DISTILLATES Filed Feb. l. 1961 N I moth HOO n. whiny D A Y zum 0(0' m 9 Nul?) K n f E e: fu/ 0 N L N I INVENTORS JOHN W'. SCOTBJR.

WILL/.4M A. RAATZ United States Patent O 3,125,502 STAGED CATALYTIC CONVERSION OF PETROLEUM DISTILLATES .lohn W. Scott, lr., Ross, and William A. Raatz, San Rafael, Calif., assignors to California Research Corpo-l ration, San Francisco, Calif., a corporation of Delaware Filed Feb. 1, 1961, Ser. No. 87,051 6 Claims. (Cl, 208-59) This invention relates to a process for upgrading petroleum distillates to gasolines of high quality. It is particularly directed to a process wherein the desired product is obtained by subjecting the feed to a staged catalytic conversion operation employing as catalyst in each stage a material incorporating a hydrogenating-dehydrogenating component supported on an active cracking base, said catalyst being relatively free of any tendency to hydrogenate aromatic compounds under the reaction conditions employed.

This application is a continuation-in-part of our copending application Serial No. 557,060, led January 3, 1956, now abandoned.

In converting petroleum distillates such as naphthas, cycle stocks from cracking operations, coker distillates, and gas oils to gasolines, a primary object is, of course, to obtain a product of high octane value. However, if the process is to be attractive from the commercial standpoint, it is also necessary to obtain the desired product in high yields and with little loss to the less valuable, normally gaseous products. It is, therefore, a particular object of this invention to provide a conversion process which affords such yields and which is also highly eilicient in the sense of providing relatively long catalyst onstream periods at good conversion levels before regeneration is required. The nature of still other objects of the invention will be apparent from a consideration of the descriptive portion to follow.

The present invention is based on the discovery that the foregoing and other objects are attained by the practice of a process wherein the petroleum distillate feed stock, along with added hydrogen, is rst passed at elevated pressures and at temperatures between about 650 F. and 950 F. through a first isomerization-cracking zone (zone I) provided with a catalyst comprised of a hydrogenating-dehydrogenating component disposed on an active cracking support, said catalyst being relatively free of any tendency to hydrogenate aromatic compounds under the reaction conditions employed. The ellluent from the first zone, which may have been joined with the whole or a portion of the whole normally liquid effluent from a second zone, or a fraction distilled therefrom, is then treated in a separation zone by conventional methods such as partial condensation, ash distillation, and fractional distillation to separate and recover hydrogen-rich and other normally gaseous products, any desired gasoline or other liquid product fraction, and a bottoms fraction to be subjected to further processing. If desired, a light gasoline boiling below about 180-200 F., and a 200 F. initial, 400 F. end point (ASTM D-86) for conventional reforming may be separated from the bottoms fraction at this point. Alternatively, if desired, a gasoline boiling up to 375 F. i 50 F. and an intermediate aromatic fraction boiling between 375 F. i 50 F. and 425 F. i 25 F. may be separated from bottoms at this point, with the intermediate fraction being recycled to one or the other of the isomerization-cracking zones, and the bottoms recycled, in whole or in part, to the zone not handling the intermediate cut. In any case, at least a substantial portion of the liquid effluent from the separation zone is recycled, together with hydrogen, to either or both the first isomerization-cracking zone (zone I), or the second isomerization-cracking zone (zone II), the

3,125,502 Patented Mar. 17, 1964 ice latter also provided with a catalyst relatively free of aromatic hydrogenation characteristics, comprising a hydrogenating-dehydrogenating component disposed on an active cracking support, at elevated pressures substantially the sarne as in zone I, and at temperatures somewhat lower in range than those employed in zone I, the zone II temperatures ranging from about 550 F. to about 850 F.

By operating in the foregoing fashion, with the bottoms from the tirst isomeriZation-cracking zone being converted in the second zone at similar pressures but at lower average temperatures than those prevailing in the said first zone, a number of highly beneficial results are unexpectedly obtained. First, the loss to normally gaseous products (C3) is well below that encountered when the conversion process is conducted in one zone only at the relatively high temperatures employed in the first zone of the present process, and with the bottoms from such single zone being cycled thereto. Secondly, the gasoline obtained by a practice of the present invention is higher in aromatic content and, therefore, in octane rating than can be obtained by operating the conversion process in one zone only, at the relatively low temperatures employed in the second zone of the present process, and with the bottoms from such single zone being recycled thereto.

A further benefit :obtained by operating in the stated fashion disclosed herein is one which assumes particular importance in the case of nitrogen-containing feed stocks which have been found to be more refractory and limiting on run length at pressures in the lower portion of the preferred range than is the case with stocks which are substantially free of nitrogen-containing components. That is to say, while the latter stocks can be converted in good yields at temperatures as low as 550 F. in some instances, it has been found that the presence of nitrogen compounds may at least partially suppress conversion at all temperatures and that temperatures in excess of 650 F. may be required to attain reasonable conversion levels. Since methane production becomes increasingly objectionable as temperatures increase above about 900 F., this effect of nitrogen compounds can reduce the working temperature range from about G50-900 F. to about G-900 F., or from a span of 250 to one of around This effects a corresponding reduction in on-stream periods, since the whole temperature range is normally traversed during the course of each such on-stream period, the practice being to employ progressively higher temperatures in order to keep the conversion level up as the activity of the catalyst falls olf. Accordingly, it has been found that not only does the use of the two-stage conversion system disclosed above make for improved octane qualities as well as better yields of the desired gasoline product, but it also makes it possible to employ the catalyst over its entire useful temperature range (normally from about 550-600 to 850-950 F.) with stocks of high nitrogen content as well as with those which are substantially nitrogen-free. The latter result is based on the discovery that, even when running a high nitrogen feed, the bottoms fraction from zone I, discussed above, is practically free of nitrogen compounds, which fraction, containing unconverted feed boiling range stock from the first stage of the operation, can thereafter be further converted in the zone Il, or second, reaction stage under the reduced temperature conditions which have been found to give a better product distribution ratio with less loss to C3- components.

The present invention is adapted to be employed in the conversion to premium gasolines of a wide variety of petroleum distillates boiling from about F. to` 1050 F. Representative starting materials from which high octane gasolines can be obtained in good yield include petroleum naphthas of straight run, catalytic or thermal cracked origin, and boiling in the range of from about 175 to 500 F.; cycle oils from thermal or catalytic cracking units, including light stocks boiling between about 380 and 600 F., heavy cycle oils boiling between about 550 and 750 F.; and gas oils boiling within the range of from about 400 to about 1050 F., together with mixtures of one or more of the foregoing stocks or fractions thereof. Said distillates may be of either the essentially nitrogenfree variety (by which it is meant to include those stocks containing less than about 3 ppm. total nitrogen) as well as those which contain nitrogen compounds in amounts equivalent to 3 ppm. nitrogen or more, said latter stocks frequently containing as much as 500 to 1500 p.p.m. nitrogen. However, in processing nitrogen-containing stocks in accordance with this invention, and particularly those containing at least p.p.m. of nitrogen, improved results are obtained by so operating that any hydrogenrich gases removed from the effluent from zone I and recycled to the system are free of ammonia. This can be accomplished, among other methods, either by waterwashing the whole eiiiuent from zone I and thereafter separating out the resulting water layer or by bringing the hydrogen-rich gaseous recycle stream into contact with a material (such as H3PO4 on kieselguhr) capable of removing the ammonia from the gaseous recycle stream. The effect of so operating is to maintain the conversion in zone I at a higher level than would otherwise be the case if the ammonia content of the gaseous recycle stream were allowed to build up in the system, the same also being true of zone II in case part or all of said stream reaches that zone.

The catalyst employed in the isomerization-cracking zones is one wherein a material having hydrogenatingdehydrogenating activity is deposited or otherwise disposed on an active cracking catalyst support. The cracking component may comprise any one or more of such fluorided or non-iiuorided acidic materials as silicaalumina, alumina-EP3, silica-magnesia, silica-aluminazirconia composites, as well as various acid-treated clays and similar materials. The hydrogenating-dehydrogenating components of the catalyst can be selected from any one or more of the various group VI and group VIII metals, as well as the oxides and suldes thereof, representative materials being the oxides and sultides of molybdenum, tungsten, vanadium, chromium and the like, as well as metals such as iron, nickel, cobalt and platinum and various oxides and suldes thereof. If desired, more than one hydrogenating-dehydrogenating component can be present, and good results have been obtained with catalysts containing composites of two or more of the oxides of molybdenum, cobalt, chromium and zinc, and with mixtures of said oxides with iiuorine. The amount of the hydrogenating-dehydrogenating component present can be varied within relatively wide limits of from about 0.1 to 15%, based on the weight of the entire catalyst. The amount of this material present should be suiicient to provide a reasonable catalyst on-stream period at required conversion levels, but insuliicient to effect substantial saturation of aromatic rings under the reaction conditions employed in the respective isomerization-cracking zones. The latter quality, referred to herein as the se- Verity factor (Sa), can be evaluated by subjecting the catalyst to a standard test employing pseudocumene as a feed stock, the test involving passing the pseudocumene through the catalyst at a liquid hourly space velocity of 2, pressure of 1200 p.s.i.g. and temperature of 800 F., and with 6000 s.c.f. of hydrogen per barrel feed, for a period of ten hours. The product is then analyzed to determine the percent of the product portion which is made up of aromatics (Ap), and then determining Sa by the following equation:

,over periods of 50 to 300 or more hours.

The above equation is derived, as shown below, from the following general relationships:

aromatics hydrogenation index (Ah) aromatics cracking index (Ac) where percent'J aromatics in feed (Af) minus percent aromatics in product boiling below initial boiling point of feed (Ap) Ah: percent aromatics in feed (Af) percent aromatics in product portion boiling below initial boiling point of feed (Axt) percent aromatics in feed (Af) Combining the above equations,

A ,gn= f D and in the case of a pseudocumene feed, A, has a value In order to be satisfactorily employed in the present linvention, the catalyst materials employed in zones I v5 are those containing from about 1 to 12% molybdenum oxide, or a mixture of from l to 12% molybvdenum oxide and from 0.1 to 5% cobalt oxide, or with mixtures of from about 0.5 to 5% each of cobalt oxide and chrominum oxide, the said hydrogenating-dehydrof genating component being deposited on an active cracking support comprising silica-alumina cracking catalyst beads having a silica content of about to 90%. Thus, the molybdenum oxide catalyst can readily be prepared by soaking the beads in a solution of ammonium molybdate, drying the catalyst for 24 hours at 220 F., and then calcining the dried material for 10 hours at 1000 F. If cobalt oxide is also to be present, the calcined beads can then be similarly treated with a solution of a cobalt compound, whereupon the catalyst is again dried and calcined. Under favorable operating conditions the isomerization-cracking catalyst will maintain high activity Catalyst activity can then be restored by a conventional regeneration treatment involving burning otf catalyst contaminants with an O2-containing gas.

While the invention will be described more particularly below in connection with the method of fixed catalyst bed operation wherein the respective catalyst beds are periodically regenerated in situ, the process is also well adapted to be carried out in a moving catalyst bed system or in one of the fluidized catalyst type. In the later methods of operation (wherein the general procedures to be employed are now well established in the art) separate vessels are employed for the respective reaction and regeneration zones. However, since in carrying out the process of this invention the catalyst retains its activity over long periods of time, it is normally preferable, from an economic standpoint, to employ the fixed catalyst bed method of operation or some modifications thereof.

Having selected or prepared the feed stock to be employed as well as the catalyst, the present process is effected in the general manner described below wherein reference is made to the figure of the appended drawing which is a simplified iiow scheme of a refinery unit suitable for use in practicing the invention. The feed enters the system through line 10 and, along with at least 1500 s.c.f. of hydrogen per barrel of feed (preferably 3000 to 30,000 s.c.f. of hydrogen) provided through lines 11 and 12, is passed through the catalyst in zone I, rcactor 13 at a liquid hourly space velocity (LHSV) of from about 0.2 to 15 (preferably 0.3 to 5), a pressure above 600 p.s.i.g. (preferably 1000 to 3000 p.s.i.g.), and a temperature of from about 650 to 950 F. The lower portion of this temperature range is normally employed in the first part of a given on-stream period, with the temperature being raised thereafter as required to maintain conversion at reasonable levels, the latter normally being regarded as about to 60% per pass in zone I. In this connection, conversion is arbitrarily expressed in terms of the amount of feed converted to product boiling below the initial boiling point of the feed. The temperatures employed herein refer to average temperatures as measured throughout the catalyst bed.

The effluent from the zone I reactor is passed through line 14 into a high pressure gas-liquid separator 15 from which a hydrogen-rich recycle stream (normally containing from about 80 to 95% hydrogen) is taken overhead through line 16 for return to zone I via

lines

12 and 10, along with added hydrogen from line 11 to make up for that consumed in zone I. Said hydrogen consumption in each zone ranges from about 1000 to 2000 s.c.f. per barrel of feed converted in the particular zone, and said conversions usually fall in the range of from about 10 to 60% in zone I and 25 to 75% in zone II, based on total feed to each zone, including any recycle.

When feeding nitrogen-containing stocks, it is desirable to provide an ammonia-free, hydrogen-rich recycle stream to line 16. This can be effected either by water washing the whole effluent in line 14 or by treating the said recycle stream per se.

When the ammonia is to be removed by treatment of the whole eilluent, the latter is admixed with an amount of water from line 17 sufficient to dissolve the ammonia readily. Good results in this respect are normally obtained by injecting from about 1 to 100 lbs. of Water per barrel of said eilluent, calculated as liquid, the amount employed being generally related to the amount of nitrogen present in the total feed to the isomerization-cracking zone. Thus, approximately 10 lbs. of water per barrel of feed can be employed with good effect in the case of total feeds (fresh plus any recycle) containing from about 15 to 25 p.p.m. of nitrogen. For most applications, a preferred range is from about 5 to 50 lbs. of water per barrel of total feed. The term water as here employed includes both fresh water as well as acidified water streams, the latter having a relatively greater capacity for removing nitrogen-containing materials from the hydrocarbon phase. The resulting water-efiiuent mixture (normally after cooling) is passed into the high pressure gas-liquid separator l5 from which a hydrogen-rich recycle stream, now substantially free of ammonia, is taken overhead in line 16 for return to the reaction zone along with additional hydrogen, while the water containing dissolved ammonia is Withdrawn as bottoms through line 18.

When separation of ammonia in the effluent stream from the isomerization-cracking zone is not effected by the addition of water or acidiiied water, as indicated above, the ammonia is separated by passing the hydrogenrich recycle gas stream in line 16 through line 20 to a contacting zone 2l where the ammonia in the gaseous stream taken overhead from the high pressure gas-liquid separator is removed in known fashion by treating the gas with water, acidied water; or other acidic medium, preferably an organic or mineral acid which may be adsorbed or otherwise supported on an inert porous support such, for example, as kieselguhr.

The remaining hydrocarbon phase from separator 15 is then fractionated, preferably after being passed through line 22 to a low pressure vapor-liquid separator 23 to remove C4 and normally gaseous components, including hydrogen sulfide, as taken overhead through line 24. The remaining, normally liquid material is then passed via line 25 to` a distillation column 26 from which one or more product fractions are recovered (as through line 27) while either an intermediate or a total bottoms fraction boiling above the end point of the desired gasoline product is passed via line 28, along with added hydrogen or hydrogen-rich gas from line 29, through the catalyst in zone II reaction vessel 30. If desired, a portion of the bottoms in line 28 can be diverted to diesel oil or the like, by passage out of the system via line 39.

In zone II the reaction conditions fall within the ranges described above for the zone I reactor, except that here the average temperature throughout the catalyst bed falls in a lower range of from about 550 to 850 F., i25". Similarly, as in the case of the zone I catalyst, temperatures are gradually raised within the indicated range to maintain catalyst activity, as reflected by good conversion levels as the on-stream period progresses. Not only are the overall operating pressure ranges for the zone I and zone II reactors substantially the same, but the specific operating pressures for each reactor during process operation are maintained at substantially the same values at any given time.

The effluent from reactor 30 passes through line 31 to a high pressure gas-liquid separator 32 from which a hydrogen-rich gaseous recycle stream is recovered through line 29 and returned to the zone Il reactor along with make-up hydrogen from line 11 as supplied through line 33. The nature of this gaseous recycle stream is essentially the same as that described above in connection with the stream in line 16, except that here the stream is essentially free of ammonia, since the bottoms product supplied from column 26 to the zone II reactor is also essentially free of nitrogen compounds.

The liquid material from gas-liquid separator 32 if; passed through line 34 to a low pressure gas-liquid sepa.- rator 35 from which C4- products are recovered through line 36, while the remaining liquid product is then either distilled so as to separate out a gasoline fraction and any other product or recycle desired fractions (with the bottoms from said separation being recycled to the zone Ii.' reactor), or else the liquid fraction from separator 35 is returned in whole or in part as a recycle stream to the zone I reactor; in any event at least a substantial pro` portion of the liquid effluent from the zone II reactor is returned to at least one of the zone I and zone II reactors. As illustrative of the first of these alternatives, the liquid stream from separator 35 is passed via

lines

37, 58 and 25 to the distillation column 26 (which thus serves as a common distillation column for the product from both reaction zones), while in the other alternative any portion of the stream in line 37 not so directed to column 26 is brought back directly to zone I via line 10.

Overhead fractions from distillation column 26 other than gasoline are not shown on the figure of the appended drawing, which is diagrammatic and intended for illustration of a simple means of practicing the process of the invention. The necessary changes and additions of side streams on column 26 and extensions of lines which must be made in order to practice the alternative operations mentioned hereinbefore, including separation of a cutafter gasoline to be reformed, or recycle of an aromatic-rich cut to one or the other of the isomerization-cracking zones, will be obvious to those skilled in the art. Equally while it is essential to the invention that at least a gasoline product fraction and a bottoms fraction be recovered in the process, and that at least a portion of said bottoms be recycled to zones I, II or both, any portion of said bottoms not recycled may be further processed by any other suitable means, including catalytic cracking.

In the ligure of the appended drawing, various required pumps, heat exchangers, valves, and other items of control equipment have been omitted in the interest of simplicity and clarity of expression, the placing of such equipment being evident to those skilled in the art once the general procedure has been outlined. Equally, the

distillation column 26 is symbolic, and may be replaced by additional columns for additional product or recycle streams if sharply cut streams are desired, and in the general case it is well known that the total number of such columns in the separation zone including the lirst, is one less than the number of sharply cut streams desired. Further, it may be desired to cut sharply between the C4s and C5s in dash drums 1S and 32, in which case flash drums 23 and 35 can be replaced by distillation columns for full fractionation, i.e., stabilizers.

As will be seen from the data presented in Table l below, which is illustrative of the results which can be obtained by a practice of the present invention, ultimati' yields (C5-H of finished gasoline are extremely high and normally range from about 95 to 105 volume percent in terms of the volume of fresh feed supplied to the isomerization-cracking Zone. From this it is obvious that the process is an extremely eflicient one wherein substantially all of the feed is converted to the desired gasoline fractions of high octane value with but relatively small losses to less valuable, normally gaseous components.

To illustrate the advantages of the present invention, reference may be had to specific feed stocks, namely, a raw catalytic cycle oil of California origin which is relatively high in nitrogen content, and the same stock after the nitrogen content has been reduced by a conventional hydroiining treatment involving passing the stock, at a temperature of 770 F. and a pressure of 800 p.s.i.g. through a catalyst containing molybdenum oxide (9% Mo) and cobalt oxide (2% Co) deposited on alumina in the presence of 3000 s.c.f. of hydrogen per barrel. The specifications of the respective stocks are as follows:

With the above stocks in mind, and comparing the results obtained by the conversion thereof by a practice of prior art methods with those of the present invention, the following general observations can be made. Firstly, when employing the raw feed in the presence of a suitable catalyst, it is not possible to reach commercially feasible conversion levels even at relatively low throughput rates, except as the temperature is raised to about 775 F. to 800 F. Since it is desired to terminate the run at temperatures of not higher than about 900 F. to 950 F., this imposes a serious limitation on the possible length of the on-stream period prior to regeneration. With a practice of the present invention, however, wherein the raw feed is processed in the rst zone at temperatures above 650 F., accepting relatively low conversion levels in zone I and with the bottoms from said zone boiling above the endpoint of the desired gasoline fraction (e.g., 350 F.) being passed to a second conversion zone operated at substantially lower temperature and relatively higher conversion levels, it is possible to obtain equally good yields of the desired gasoline product (the C5350o F. yield approaching or even exceeding 100 volume percent, based on fresh feed), while extending the temperature span of the catalyst, as sequentially employed in the respective Zones over a broad range of 200 or more degrees. Further, the quality of the product gasoline, as indicated by octane number, is normally 2b somewhat better than that obtained by prior art practices.

When the economics of the operation justify the added expense of the hydrofining step, it has been found that the one-stage method of the prior art can be employed without particular disadvantage as to temperature span during each on-stream portion of the operating cycle. However, the octane number of the product so produced is normally several numbers below that of the gasoline of similar boiling point range produced by the two-stage method of the present invention. This result is one which is thought to be attributable to the fact that in the stage I reactor, where the feed is initially subjected to relatively high temperatures, the undesired aromatic hydrogenation reactions are suppressed while disproportionation reactions are correspondingly favored, the latter reactions being those whereby desired lower boiling aromatics such as benzene, toluene, xylene, ethyl toluene, pseudocumene, etc., for example, are formed by disproportionation of higher boiling feed aromatics with the simultaneous production of still higher molecular weight homologous alkyl substituted benzenes such as pentaand hexamethyl benzenes and the corresponding ethyl substituted compounds. The highly substituted aromatics so produced in the first stage are thereafter saturated, with accompanying ring fracture, into isoparafiins of exceptionally high octane rating as the bottoms stream containing said highly substituted aromatics passes into the second reactor stage where the relatively low temperatures employed favor said aromatic ring hydrogenation reactions. Alternatively, if desired, the lower boiling aromatics may be removed before the bottoms stream is passed to the low temperature second reactor stage, so that they are conserved.

We claim:

1. A process for upgrading a petroleum distillate to a gasoline of relatively high octane rating, which comprises passing said distillate along with at least 1500 s.c.f. of hydrogen per barrel of feed through a catalyst relatively free of aromatics hydrogenation characteristics under the reaction conditions employed incorporating a hydrogenating-dehydrogenating component disposed on an active cracking support in a first conversion zone at pressures during the on-stream period in the range of 600 to 3000 p.s.i.g. and at temperatures of from about 650 to 950 F. sufficiently high to eifect a conversion of the feed of from 10 to 60% per pass; separating the effluent from said zone into at least a hydrogenrich gaseous fraction, a gasoline product fraction boiling below 350 F. and a bottoms fraction boiling above 350 F.; feeding said bottoms fraction along with at least 1500 s.c.f. of hydrogen per barrel thereof through a catalyst relatively free of aromatics hydrogenation characteristics under the reaction conditions employed incorporating a hydrogenating-dehydrogenating component disposed on an active cracking support in a second con- Version zone at pressures throughout the on-stream period in the range of 600 to 3000 p.s.i.g. and at instantaneous pressures during the on-stream period substantially the same as in said first conversion zone and at temperatures substantially lower than those in said first conversion zone, said second zone temperature ranging from about 550 to 850 F. and being sufficiently high to elfect a conversion of said bottoms feed to lower boiling products of from 25 to 75% per pass; and recycling at least a substantial portion of the liquid effluent from said second conversion zone to at least one of said conversion zones; said process being characterized by a net consumption of from about 1000 to 2000 s.c.f. of hydrogen per barrel of feed converted in each of said zones.

2. The process of claim l, wherein a nitrogen-containing petroleum distillate is employed as feed to said first zone and wherein at least a portion of said hydrogenrich gaseous fraction recovered from said eflluent from mamada 9 said first zone is freed of ammonia and is recycled to said first zone.

3. The process of claim l, wherein a portion of said bottoms fraction containing lower boiling aromatics is withdrawn instead of being passed to said second conversion zone, in order to conserve said aromatics.

4. A process as in claim 1, wherein the catalyst in each zone has a severity -factor of from about 0 to 5.0.

5. A process as in claim 1, wherein said iirst conversion zone is operated in the lower portion of said temperature range of from about 650 to 950 F. during the rst part of each on-strearn period, and the temperature is raised thereafter as necessary to maintain desired conversion levels within the range of about 10 to 60% pass.

6. A process as in claim 1, wherein said second conversion zone is operated in the lower portion of said temperature range of from about 550 to 850 F. during the first part of each on-strearn period, and the temperature is raised thereafter as necessary to maintain desired conversion levels within the range of about 25 to 75% per pass.

References Cited in the file of this patent UNTTED STATES PATENTS 2,464,539 Voorhies et al. Mar. l5, 1949 2,945,800 Ciapetta et al July 19, 1960 2,945,801 Ciapetta et al. July 19, 1960 3,037,930 Mason June 5, 1962

Claims

1. A PROCESS FOR UPGRADING A PETROLEUM DISTILLATE TO A GASOLINE OF RELATIVELY HIGH OCTANE RATING, WITH COMPRISES PASSING SAID DISTILLATE ALONG WITH AT LEAST 1500 S.C.F OF HYDROGEN PER BARREL OF FEED THROUGH A CATALYST RELATIVELY FREE OF AROMATICS HYDROGENATION CHARACTERISTICS UNDER THE REACTION CONDITIONS EMPLOYED INCORPORATING A HYDROGENATING-DEHYDROGENATING COMPONENT DISPOSED ON AN ACTIVE CRACKING SUPPORT IN A FIRST CONVERSION ZONE AT PRESSURES DURING THE ON-STREAM PERIOD IN THE RANGE OF 600 TO 3000 P.S.I.G. AND AT TEMPERATURES OF FROM ABOUT 650* TO 950*F. SUFFICIENTLY HIGH TO EFFECT A CONVERSION OF THE FEED OF FROM 10 TO 60% PER PASS; SEPARATING THE EFFLUENT FROM SAID ZONE INTO AT LEAST A HYDROGENRICH GASEOUS FRACTION, A GASOLINE PRODUCT FRACTION BOILING BELOW 350*F. AND A BOTTOMS FRACTION BOILING ABOVE 350*F.; FEEDING SAID BOTTOMS FRACTION ALONG WITH AT LEAST 1500 S.C.F. OF HYDROGEN PER BARREL THEREOF THROUGH A CATALYST RELATIVELY FREE OF AROMATICS HYDROGENATION CHARACTERISTICS UNDER THE REACTION CONDITIONS EMPLOYED INCORPORATING A HYDROGENATING-DEHYDROGENATING COMPONENT DISPOSED ON AN ACTIVE CRACKING SUPPORT IN A SECOND CONVERSION ZONE AT PRESSURES THROUGHOUT THE ON-STREAM PERIOD IN THE RANGE OF 600 TO 3000 P.S.I.G. AND AT INSTANTANEOUS PRESSURES DURING THE ON-STREAM PERIOD SUBSTANTIALLY THE SAME AS IN SAID FIRST CONVERSION ZONE AND AT TEMPERATURES SUBSTANTIALLY LOWER THAN THOSE IN SAID FIRST CONVERSION ZONE, SAID SECOND ZONE TEMPERATURE RANGING FROM ABOUT 550* TO 850* F. AND BEING SUFFICIENTLY HIGH TO EFFECT A CONVERSION OF SAID BOTTOMS FEED TO LOWER BOILING PRODUCTS OF FROM 25 TO 75% PER PASS; AND RECYCLING AT LEAST A SUBSTANTIAL PORTION OF THE LIQUID EFFLUENT FROM SAID SECOND CONVERSION ZONE TO AT LEAST ONE OF SAID CONVERSION ZONES; SAID PROCESS BEING CHARACTERIZED BY A NET CONSUMPTION OF FROM ABOUT 1000 TO 2000 S.C.F. OF LHYDROGEN PER BARREL OF FEED CONVERTED IN EACH OF SAID ZONES.