|Publication number||US3172839 A|
|Publication date||Mar 9, 1965|
|Filing date||Dec 4, 1961|
|Priority date||Dec 4, 1961|
|Publication number||US 3172839 A, US 3172839A, US-A-3172839, US3172839 A, US3172839A|
|Inventors||Robert H. Kozlowsld|
|Export Citation||BiBTeX, EndNote, RefMan|
|Patent Citations (7), Referenced by (8), Classifications (10)|
|External Links: USPTO, USPTO Assignment, Espacenet|
United States Patent O CATALYTEC CONVERSION PROCESS FOR 'II-1E PRODUCTION F LOW LUMENOSHTY FUELS Robert H. Kozlowski, Berkeley, Calif., assigner to California Research Corporation, San Francisco, Caiif., a
corporation of Delaware Filed Dee. 4, 1961, Ser. No. 156,669 4 Claims. (Cl. 20S-68) INTRODUCTION This invention relates to a hydrocarbon conversion process, and, more particularly, to a process for the catalytic conversion of petroleum distillates to produce fuels characterized by high luminometer numbers, high smoke points, and low freezing points.
PRIOR ART It is well known that high smoke points and low freeze points are desirable characteristics of middle distillate boiling range fuels, particularly jet fuels. Only recently it has been discovered that another desirable characteristic of such fuels that can be critical to proper engine operation is a high luminometer number, indicating a fuel that burns with a low luminosity. Two different fuels can burn equally hot and yet one can be more luminous than the other. The luminosity of the more luminous flame is caused by incandescence of molecular fragments in the ame; these glowing fragments Vgive off radiant heat, which increases the temperature of the surrounding engine parts without -adding engine power. Luminosity of a fuel is determined by a luminometer number (LN) rating, according to how little radiant heat the fuel gives of during burning. The higher luminometer numbers indicate less radiant heat given olf land a more desirable fuel. Most jet fuels to date have luminometer numbers of 4510 65.
Heretofore, in producing jet fuels by hydrocracking a feed of a particular boiling range, the luminometer number .of the jet -fuel products has been essentially xed by crude source and feed boiling range, a little improvement in the luminometer number has been obtainable by varying the process operation. Heretofore, in such an operation, it has Abeen possible, by various process techniques, to lower the freezing point and raise the smoke point of the jet fuel product to Vsome extent; however, there has been a continuing need for process `improvement in these respects.
OBJEFTS In view of the foregoing, itis an object of the present invention to provide a method for operating a process for hydrocracking particular hydrocarbon distillate stocks to produce jet fuels which will result in a higher luminometer number jet fuel product than heretofore obtainable in such a process with the same feeds, and which will also result in reducing and therefore improving the freezing point of the jet fuel product.
It is a further object of the present invention to provide a method for operating a hydrocracking process -for producing jet fuels which will -result in raising the jet fuel product smoke point, particularly in those situations in which the jet fuel product obtained when hydrocracking a particular stock does not quite lmeet smoke point specifications because of the `character of the feed.
DRAWING This invention will be more clearly understood, and further .objects and advantages thereof Will be apparent, from the following description when read in connection with the accompanying drawing. The drawing is a `diagrammatic illustration of an embodiment of process units and iiow paths suitable for carrying out the process of the invention.
STATEMENT OF INVENTION In accordance with the present invention, there is provided a hydrocracking process for producing jet fuelsof high luminometer number which comprises contacting a hydrocarbon feed selected from the group consisting of hydrocarbon distillates boiling above -about 350 F. and hydrocarbon residua boiling above about 1050 F. in a hydrocracking zone in the presence of at least 1000 standard cubic feet (scf.) of hydrogen per barrel of said feed with a catalyst comprising a hydrogenatingdehydrogenating component disposed .on an active cracking support at a temperature of about from 400 to 900 F., a pressure of at least 500 p.s.i.g., and a liquid hourly space velocity (LHSV) of about from 0.1 to 15.0, Withdrawing from said zone at least a normally gaseous fraction, at least one naphtha fractiomand a jet fuel fraction boiling in the range of about from 320 Vto 550 F. and recycling to said zone at least 5 volume percent of the portion of said jet fuel fraction boiling below the initial boiling point of said feed.
In a preferred embodiment of the present invention, the hydrocarbon feed boils in the range of .about from 500 to l050 F., the feed is contacted in the presence of about at least 1000, preferably 1500 to 30,000, and still more preferably 1500 to 6000 s.c.f. of hydrogen per barrel of feed at a temperature of from about 500 to 800 F., a pressure of about from `800 to 3000 p.s.i.g., and an LHSV of about from 0.2 .to 3.0, and there is recycled to said zone aboutfrom 5 to 95 volume percent, and more preferably about 10 to 50 volume percent, of the portion of said jet fuel fraction boiling below the initial boiling point of said feed.
An especially effective method of operating the process of the present invention to obtain a final jet fuel product of even higher luminometer number and higher smoke point is to hydrogenate the aromatics in the jet fuel porv tion to be recycled prior to returning it to the hydrocracking zone.
It is advantageous in operating the process of the resent invention to return to the hydrocracking zone a hydrogen-rich yrecycle stream obtained from the efiiuent from said zone, along with the make-up hydrogen necessary to replace that consumed in the process, which operates with a net consumption of hydrogen.
The feed stocks employed in the process of the present invention preferably boil over a range of at least F. within the aforesaid boiling ranges; suitable feed stocks include those heavy distillates normally defined as heavy straight run gas oils and heavy cracked kcycle oils, as well as conventional FCC feeds and portions thereof. Cracked stocks may be obtained from thermal or catalytic cracking of various stocks, including those obtained from petroleum, gilsonite, shale and coal tar. Residual feeds may include Minas paratlnic residua, which may boil above 1l00 F., and other parainic-type residua boiling above about 1050 P.
The most preferred feeds to the present process are those yhaving an initial boiling point of at least 500 F., particularly when the feeds are to be hydrocracked in the presence of a non-acidic or only weakly acidic catalyst. With such feeds the middle distillate products, including jet fuels, tend to have more superior properties, in that they are more naphthenic, less aromatic (therefore having higher smoke points), and lower in normal parailins (therefore having lower freeze points), than products from feeds having Aa lower initial boiling point. Feeds having lower initial boiling points tend to produce excessive quantities of non-synthetic products having a high aromatics content and therefore an unacceptably low smoke point, although the freeze point may be satisfactory. More of the product in the jet fuel boiling range is satisfactory for jet fuel purposes when the initial boiling point of the feed is at least 500 F. than when it is lower.
NITROGEN CONTENT F FEED While the invention can be practiced with utility in connection with hydrocarbon feeds to the hydrocracking zone which contain relatively large quantities of nitrogen, the operation becomes much more economical with stocks containing less than 200 parts per million (p.p.m.), preferably less than 100 ppm., and much more preferably less than p.p.m. of total nitrogen. A reduction in feed nitrogen level generally permits the hydrocracking reaction to be conducted at lower temperatures than with feeds containing relatively large amounts of nitrogen compounds. Therefore, in the case of feeds which are not inherently low in nitrogen, acceptable levels can be reached by hydroflning the feed prior to passing it into the hydrocracking zone.
In general, the effect of a total nitrogen content in excess of l0 p.p.m. in the hydrocracking step is a reduction in catalyst activity which is reflected in reduced operational eiciency and poorer product distribution. As the nitrogen content increases above the specified maximum, higher reaction temperatures are necessary to maintain an economic per-pass conversion level. These higher reaction temperatures cause a disproportionate increase in the amount of product converted to gases and carbonaceous residues deposited on the catalyst surface and thus further decrease catalyst activity. Such further decrease in catalyst activity must be compensated for by resort to still higher operating temperatures if acceptable conversion is to be maintained; therefore, nitrogen causes the on-stream life of the catalyst to be shortened because unacceptable operating temperatures are reached sooner when nitrogen is present.
The effect of nitrogen upon the hydrocracking reaction is marked at low operating temperatures, for example, 500 to 800 F., but minor at higher operating temperatures, for example, above 850 F, Accordingly, where a feed to be processed in accordance with the present invention contains more than l0 p.p.m. total nitrogen, and particularly where an acidic-type hydrocracking catalyst is used in the hydrocracking zone, denitrification of the feed will be desirable, because the process, at least for substantial portions of the ori-stream period, desirably is conducted at those low temperatures within the ranges disclosed herein that are sufficient to maintain desired per-pass conversions. Where a nonacidic-type hydrocracking catalyst is used in the hydrocracking zone, denitritication tends to take place in that zone along with hydrocracking, as discussed hereinafter, and there is less need for a previous hydroiining step.
As noted above, feed stocks containing more than about l0 ppm. total nitrogen preferably are subjected to a pretreating operation that is relatively selective for the removal of nitrogen compounds. The desired low nitrogen levels may be reached, for example, by intimately contacting the feed stocks with various acidic media, such as H2804 or th other liquid acids, or, in the case of feeds that are comparatively low in nitrogen compounds, with such solid acidic materials as acid ion exchange resins and the like. However, it is preferred to carry out deni-trication by catalytic hydrogenation (hydroining) of the feed. This entails contacting the feed at temperatures of from about 400 to 900 F., preferably from about 500 to 800 F., pressures of at least 300 p.s.i.g., liquid hourly space velocities or from about 0.3 to 5.0, along with at least 500 scf, of hydrogen per barrel of feed, with a sulfur-resistant hydrogena-l tion catalyst. Any of the known sulfactive hydrogenation catalysts may be used in the hydroning pretreatment. The preferred catalysts have as their main active ingredient one or more oxides or suldes of the transition metals, such as cobalt, molybdenum, nickel and tungsten. These various materials may be used in a variety of combinations with or without such stabilizers and promoters as the oxides and carbonates as K, Ag, Be, Cu, Mg, Ca, Sr, Ba, Ce, Cr, Th, Si, Al and Zr. These various catalysts may be unsupported or disposed on various conventional supporting materials, for example charcoal, fullers earth, kieselguhr, silica gel, alumina, bauxite and magnesia. While any of the noted classes of conventional sulfactive hydrogenation catalysts may be employed, it has been found that particularly desirable catalysts are: (l) a molybdenum oxide catalyst promoted by a minor amount of cobalt oxide and supported upon an activated alumina, (2) tungsten sulfide on activated alumina, and (3) a molybdenum sulfide catalyst promoted by a minor amount of nickel sulfide supported on activated alumina. The catalyst may be in the form of fragments or formed pieces such as pellets, extrudates and cast pieces of any suitable form or shape.
An effective hydroiining catalyst comprises cobalt irnpregnated on a coprecipitated molybdena-alumina (e.g., prepared in accordance with the disclosures of U.S. Patent 2,432,286 to Claussen et al., or U.S. Patent 2,697,006 to Sieg), combined with cobalt oxide, the final catalyst having a metals content equivalent to about 2% cobalt and 7% molybdenum.
Operable hydrolinng conditions are temperatures of 700 to 800 F., pressures of 200 to 3000 p.s.i.g., space velocities of 0.5 to 3.0, and 1000 to 15,000 s.c.f. of hydrogen per barrel of hydrocarbon feed.
OPERATING CONDITIONS IN HYDROCRACKING ZONE The hydrocarbon feed and hydrogen are contacted in the hydrocracking Zone at pressures of at least 500 p.s.i.g., preferably about from -800 to 3000 p.s,i.g. The contacting temperature is about from 400 to 900 F., preferably 500 to 800 F. The operating temperature during the on-stream period preferably is maintained at as low a value `as possible consistent with maintaining adequate per-pass conversions as catalyst fouling progresses. While those skilled in the art will realize that the desired initial and terminal temperatures will be intiuenced by various factors including character of feed and catalyst, generally speaking it will be desirable to operate the process with an initial on-stream temperature of about from 500 to 650 F., with a progressive increase to about 750 to 800 F., to maintain substantially constant conversion of at least 25 volume percent, preferably 35 to 90 volume percent per pass, of the hydrocarbon feed to products boiling below the initial boiling point of that feed. Higher conversions generally are possible with more acidic catalysts; however, such catalysts give a different product distribution than less acidic catalysts, as discussed below.
HYDROCRACKING CATALYST, GENERAL The catalyst employed in the hydrocracking zone comprises a material having hydrogenating-dehydrogenating activity, and either: (l) deposited or otherwise disposed on an active acid cracking catalyst support, or (2) unsupported, or deposited or otherwise disposed on an active cracking support that is either non-acidic or only weakly acidic.
Where the catalyst comprises a non-acidic or only weakly acidic support, it tends to produce greater amounts of products boiling in the middle distillate boiling range than does a catalyst comprising an acidic support.
ACIDIC HYDROCRACKIN G CATALYSTS Where the support is acidic, the cracking component may comprise any one or more of such acidic materials as silica-alumina, silica-magnesia, silica-alumina-zirconia composites, alumina-b'oria, fluorided composites, and the like, as well as various acid-treated clays and similar materials. Preferred catalysts will comprise silica-alumina supports having silica content in the range of from about to 99 percent by weight. The hydrogenatingdehydrogenating components of the catalyst can be selected from any one or more 'of the various Groups V, Vl, VI and VIII metals, as well as the oxides, suliides and selenides thereof, alone or together with promoters or stabilizers that may have by themselves small catalytic effect, representative materials being the oxides, suliides and selenides of molybdenum, tungsten, vanadium, chromium and th like, as well as of metals such as iron, nickel, cobalt, platinum and palladium. 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 uorides. The amount of the hydrogenating-dehydrogenating component present can be varied within relatively wide limits 'of from about 0.5 to 30% based on the weight of the entire catalyst.
Exemplary acidic-type catalysts having satisfactory characteristics as aforesaid include those containing (a) about l to 12% molybdenum oxide, (b) a mixture of from 1 to 12% molybdenum oxide and from 0.1 to 10% cobalt oxide or nickel oxide, (c) mixtures of from about 0.5 to 10% each of cobalt oxide and chromium oxide,
(d) 0.1 to v10% nickel, nickel oxide or nickel sulfide, A'
(e) 0.1 to 10% cobalt, cobalt oxide or cobalt sulfide, (f) mixtures of from 0.1 to 10% each of nickel and cobalt, as metal, oxide or sulfide, (g) 0.1 to 5% platinum or palladium, in each case the said hydrogenating-dehydrogenating component being deposited on an active cracking support comprising silica-alumina beads having a silica content of about 70 to 99%. Thus, the molybdenum oxide catalyst can be prepared readily 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 l000 F. lf 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 hydrocracking catalyst will maintain high activity over periods of to 300 or more hours. The activity of the used catalyst can then be increased, if desired, by a conventional regeneration treatment involving burning od catalyst contaminants with an oxygen-containing gas.
NON-ACIDIC HYDROCRACKING CATALYSTS When a non-acidic, or only weakly acidic, hydrocracking catalyst is used in the hydrocracking step of the present process, the catalyst should be one that is capable of converting the feed at a per-pass conversion of at least 10 to 50 volume percent of said feed, under operating conditions in the hydrocracking zone, in large part to reaction products in the synthetic middle distillate boiling range, i.e., products boiling not only in the middle distillate boiling range but also below the initial boiling point of the feed.
A suitable catalys-t comprises a hydrogenat-ing-dehydrogenating component alone or on `a support comprising at least one metal, metal oxide, metal sulfide, metal selenide or combination thereof, preferably oxides or sulides of metals of Groups Vl and Vlll ofthe Periodic Table. T-he most preferred .catalyst will comprise combinations of suliides `of cobalt and/ or nickel with sulfides of molybdenum and/ or tungsten.
The catalyst generally will comprise the aforesaid hydrogenating-dehydrogenating 'component disposed on a support that is substantially non-acidic or -a-t the most only weakly acidic. Examplary supports include silica, charcoal, kieselguhr, titania, zirconia, bauxite and alumina, with alumina being an especially preferred support. While alumina sometimes is considered to be weakly acidic, its acidity is so lou.l compared with silica-alumina, for example, that it may be considered for purposes of the present process to `be non-acidic, particularly in view of the markedly different product distribution it provides as compared with silica-alumina support. For purposes of the present process, the support particularly cannot be an acidic-mixed oxide, for example, silica-magnesia, aluminasboria` or silica-alumina.
An outstanding catalyst composite fulfilling the aforesaid requirements -bolth as to hydrogenating-dehydrogenating component and support, is the sulded catalyst comprising 4 to l0 weight percent nickel, as metal, and 15.5 to 30 weight percent molybdenum, as metal, on a substantially non-acidic 'base consisting essentially of alumina.
The aforesaid catalyst combination results ina significantly different product distribution from that obtained with acidic-type hydrocracking catalysts; it does not exhibit the high cracking `activity of lthose catalysts even yat higher temperatures and accordingly the maximum yield of products is in a higher molecular weight range than in the case of acidic-type hydrocracking catalysts. Further, the catalyst combination #tends to give a much wider boiling range spectrum of products than does an acidictype hydrocracking catalyst. Still further, the maximum total yield of synthetic products, i.e., those products boiling `belovv the initial boiling point of the feed, occurs in a molecular weight range adjacent to and immediately below the initial boiling point of the feed, whereas, in the case of an acidic-type hydrocracking catalyst, this maximu-m yield occurs in `a lower boiling range. Clearly, of the multitude of possible compounds ina given feed, many of these compounds must undergo dilferent cracking and other reactions in the presence of the aforesaid non-acidic- `type hydrocracking catalyst than they do in the presence of acidic-type hydrocracking catalysts; otherwise, the substantial difference in yield structure obtained with the two types of catalyst could not be accounted for.
A corollary feature `of the use of the aforesaid nonacidic-type or substantially non-acidic-type hydrocracking catalysts in the process of the present invention is that such a catalyst generally has excellent denitrilication activity, and ywhere nitrogen is lpresent in the :feed to the process, the catalyst efficiently converts it in the reaction zone to ammonia which may be removed `from the reaction zone eluent by conventional procedures such as by water `scrubbing rlhe cracking accomplished in the hydrocracking zone facilitates denitriiication because, upon the breaking of carbon-to-carbon bonds, nitrogen is more easily removed. At higher levels of cracking conversion, nitrogen is more easily removed than at lower levels. At higher levels of cracking conversion, higher pressures are required to prevent rapid fouling and deactivation of the catalyst.
PROCESS OPERATON Referring now to the drawing, there `shown is an exemplary overall proces-s ilow diagram suitalble for carrying out the process of the present invention.
As discussed above, the ffeed supplied tothe hydrocracking zone 19 through line lo includes heavy straight run gas oil and heavy cracked cycle oils. rThe drawing illustrates an embodiment wherein a catalytically cracked stock containing nitrogen is supplied to the hydrocracking zone 19 through line 16. In the embodiment shown, a hydrocarbon `fraction suitable for use as `a catalytic cracking feed stock is passed through line 10` to a catalytic cracking zone 11, wherein it is contacted with a conventional cracking catalyst under conventional cracking conditions.
From catalytic cracking Zone 11, an eflluent is passed tthrough line 12 to first fractionation zone 13 where it is separated into fractions of various boiling ranges. A normally gaseous stream is withdrawn through line 14, and
a naiphtha stream is withdrawn through line either as a product or -for further processing, for example in a reforming zone. A heavy cycle oil is recycled to catalytic cracking zone 11 through iline 17. A tarry bottom-s fraction is removed through line 18.
A hydrocarbon stock boiling in the range of about 350 to l050 F., in this embodiment a cracked stock, is passed through line 16 into hydro'cracking zone 19, where it is hydrocracked and, if it contains nitrogen, may also be at least partially dem'trified. Hydrogen lfor the hydrocracking and any denitrification reactions in zone 19 is supplied to that zone through lines 20 `and 21. Conversion products yfrom zone 19 4are withdrawn through line 22, where they are contacted in high pressure separator 23 with water supplied through line 24 to facilitate nitrogen rernoval. From high pressure separator 23, a hydrogen stream is recycled through line 20. From high pressure separator Z3, conversion products are passed .to low pressure separator 25 from which water is removed through line 26. A gas stream is separated through Lline 27 from the conversion products in separator 25, 'and the remainder of the conversion products are passed through Iline 2% to second fractionation zone 29.
From second fractionation zone 29, a gas stream is withdrawn through line 30, and a naphtha stream is removed from the system through line 3l, either as a product or for further processing, for example in a reforming zone. If desired, a liquid product boiling above the naphtha boiling range and within the feed boiling range may be withdrawn as a product from second fractionation zone 29 through line 32. A bottoms stream may be withdrawn through lines 33 and 34 as a product, for example as a diesel fuel, or all or a portion thereof may be recycled through line 35 to catalytic cracking zone 11 for further processing, and/ or through lines 35 and 45 to hydrocracliing zone 19 for further processing.
From second fractionation zone 29, a synthetic middle distillate stock, i.e., a middle distillate product boiling below the boiling range of the feed, more particularly, one boiling in the range of about 320 to 550 F., is Withdrawn through line 36 and at least 5 volume percent thereof is withdrawn through line 37 as a high luminometer number jet fuel product. At least 5 volume percent, and up to 95 volume percent, preferably from 20 to 60 volume percent, of the stock Withdrawn through line 36 is recycled through that line to hydrocracln'ng zone i9.
Desirably, this recycle portion is lirst diverted through line 38 and hydrogenated in hydrogenation zone 39 in the presence of hydrogen supplied to zone 39 through line 40, the resulting hydrogenated stock being returned to line 36 through line 41. Hydrogenation of this stock may be accomplished with particular ease because the stock is a particularly clean one. Additional stock from an external source may be passed through line 42 for hydrogenation in hydrogenation zone 39.
If desired, a hydrogenated jet fuel net product stream may be withdrawn from the system through line 43.
Hydrogenation zone 39 may be a conventional hydrogenation zone operated under conventional hydrogenation conditions with such conventional hydrogenation catalysts as alumina-supported platinum, palladium, nickel, rhodium, or cobalt-molybdate, with platinum and nickel being preferred. Particularly effective results will be obtained when using a platinum catalyst comprising 1.0 weight percent platinum supported on alumina at temperatures of 550 to 650 F., pressures of 250 to 750 p.s.i.g., weight hourly space velocities of 0.1 to 10.0, and hydrogen feed rate of from 1500 to 4000 s.c.f. of hydrogen per barrel of stock contacted in zone 39.
If it is desired to add additional quantities of feed to the system from an external source, but it is not desired to catalytically crack, hydrogenate, or otherwise treat these additional portions first, they may be introduced into the system through line 44.
INCREASED LUMINOMETER NUMBER AND SMOKE POINT, AND LOWER FREEZE POINT, OF I ET FUEL PRODUCT VHEN OPERATING iN ACCORDANCE WITH PRESENT INVENTION Luminometer numbers are in inverse relation to the luminosity of a jet flame in a jet engine. Accordingly, jet flame will have the lowest luminosity when the jet fuel has the higher luminorneter number. It has been found that with increasing concentration of normal parafins in a jet fuel the luminometer number increases more rapidly than with increases in concentration of isoparafns, aromatics or naphthenes; however, this advantageous effect generally is outweighed by the adverse effect that normal paraifins have on jet fuel freeze point. It has been found that aromatics and naphthenes in a jet fuel have an adverse eifect in that the luminometer number of the jet fuel is lowered as the ratio of aromatics and naphthenes to parafns increases. It has been found that isoparains, while not having as great a tendency to increase the luminometer number of jet fuel as do normal parafiins, nevertheless do exert a markedly beneficial effect on luminometer number, as compared with aromatics and naphthenes. This markedly benecial effect can be obtained by increasing the volume percentage of isoparaffins in a jet fuel, and such increase is not accompanied by the deleterious effect on freeze point that is produced with normal paraflins. Accordingly, as among aromatics, naphthenes, normal paratfins and isoparaflins, the latter are the best species for use in improving jet fuel luminometer number without undue deleterious effect on freeze point.
In the operation of the present process, as described above, it has been found that the synthetic middle distillate fraction Withdrawn from second fractionation zone 29 through line 36 has a high content of isoparaflins and accordingly is especially adapted for use partly as a jet fuel product and partly as a recycle stock to hydrocracking zone )i9 to further increase the luminometer number of the net jet fuel product withdrawn through line 37. By recycling at least 5 volume percent of said synthetic middle distillate fraction, the volume percent of isoparaffins in said net jet fuel product is caused to build up, and the luminometer number of that product increases. It has further been found that hydrogenation in hydrogenation zone 39 of the recycle synthetic middle distillate product, by reducing the aromatic content of that product, causes a desirable increase in the smoke point of the ultimate end jet fuel product by converting aromatics to naphthenes. This desirable effect is accompanied by a still further increase in the luminometer number of the ultimate end jet fuel product, because aromatics have a lower luminometer number than do naphthenes.
TYPES OF OPERATION TO WHICH PROCESS IS ADAPTED While the invention will be described more particularly in connection with the method of fixed hydrocracking catalyst bed operation wherein the hydrocracking catalyst bed may be periodically regenerated in situ, the process is also well adapted to be carried out in a moving catalyst bed, or in a slurry-type reaction system, or in one of the iluidized catalyst type. However, since in carrying out the process of this invention the catalyst retains its activity over long periods of time, it is normally preferably, from an economic standpoint, to employ the xed catalyst bed method of operation or some modification thereof.
The feed may be introduced into the hydrocracking zone as a liquid, vapor, or mixed liquid-vapor phase, depending upon the temperature, pressure, proportions of hydrogen and boiling range of the charge stocks utilized.
EXAMPLES In the following example of a preferred embodiment of the present invention, the operating variables and conversion factors have been determined by correlation of experimental data where necessary. Thus, in the examples, the results that can be expected from a 10,000 barrel per day (b.p.d.) hydrocracking zone feeding a 550 to 850 F. petroleum distillate under the specified operating conditions are shown in Table I. For comparative purposes only, Case I shows a hydrocracking operation wherein none of the 320 to 550 F. synthetic middle distillate product is recycled to the hydrocracking zone. Case II shows an operation in accordance with the present invention wherein 50 volume percent of the portion of the 320 to 550 F. synthetic middle distillate product is recycled to the hydrocracking zone with the remainder of that product being withdrawn from the system as a jet fuel of high luminometer number.
TABLE I Case I Case II Catalyst, Nickel Sulde on Silica-Alumina:
Fresh Feed, 55u-850 F. Hydronned Heavy Catalytic Cycle Oil, b.p.d 6, 000 6, 000 Recycle, 550-850F. Fraction, b.p.d 000 2, 700 Recycle, 320-550F. Fraction, b.p.d 0 1, 300 Total Feed (including recycle), b d 10, 000 10, 000 Temperature, F 600 600 Pressure, p.s.i 1, 200 1, 200 Hydrogen Recycle, s c f [b Fresh Feed- 6, 500 6, 500 Jet Fuel 2, 600 l, 300 Luminometer Number, CRC method 1 50 62 Smoke Point, ASTM, mm 23 27 Isoparain Content, percent 42 50 Freeze Point, F beloi belowg lThe method of the Coordinating Research Council of the American Petroleum Institute.
1. The method of producing jet fuels of high luminometer number, which comprises contacting a hydrocarbon feed selected from the group consisting of hydrocarbon distillates boiling above about 500 F. and hydrocarbon residua boiling above about l050 F. in a hydrocracking zone, in the presence of at least 1000 s.c.f. of hydrogen per barrel of said feed, with a catalyst comprising a hydrogenatingdehydrogenating component associated with an active cracking support at a temperature of about from 400 to 900 F., a pressure of at least 500 p.s.i.g., and an LHSV of about from 0.1 to 15, Withdrawing from said zone at least one normally gaseous fraction, at least one naphtha fraction, and a synthetic jet fuel fraction boiling in the range of about from 320 to 500 F., recycling to said zone at least 5 volume percent of the portion of said synthetic jet fuel fraction, and recovering as a net product at least a portion boiling below about 500 F. of the unrecycled remainder of said synthetic jet fuel fraction.
2. The method as in claim 1, wherein the recycled jet fuel fraction portion is hydrogenated to reduce the aromatic content thereof before it is passed to said hydrocracking zone.
3. The method as in claim l, wherein said distillate feed is a cycle oil from a catalytic cracking unit.
4. The method as in claim 1, wherein said distillate feed is a straight-run petroleum distillate.
References Cited by the Examiner UNITED STATES PATENTS 2,944,005 7/ 60 Scott 208-109 2,965,564 12/60 Kirshenbaum et al 20S-134 3,000,815 9/61 Haney 208-143 3,008,895 11/61 Hansford et al. 208-112 3,012,961 12/61 Weisz 208-15 3,072,560 1/63 Paterson et al. 208-55 3,092,567 6/63 Kozlowski et al. 208-57 ALPHONSO D. SULLIVAN, Primary Examiner.
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|U.S. Classification||208/68, 208/111.35, 208/15, 208/143|
|International Classification||C10G47/00, C10G65/12, C10G69/04|
|Cooperative Classification||C10G47/00, C10G2400/08|