US3185639A - Hydrocarbon conversion process - Google Patents

Description

N. J. PATERSON Filed April 6, 1964 HYDROCARBON CONVERS ION PROCESS May 25, 1965 United States Patent O A v 3,185,639 v HYDROCARBON CONVERSION PROCESS Norman J. Paterson, San Rafael, Calif., assignor to California Research Corporation, San Francisco, Calif., a corporation of Delaware Filed Apr. 6, 1964, Ser. No. 358,151 1 Claim. (Cl. 208-68) INTRODUCTION This application is a continuation-in-part of application Serial No. 184,646, filed March 30, 1962, now abandoned.

This invention relates to a process for the catalytic conversion of petroleum hydrocarbons, including both gas oil portions thereof and residual portions thereof, to gasoline and middle distillate fractions. More particularly, the invention relates to an integrated renery process wherein nonresidual and residual hydrocarbon stocks boiling essentially above the gasoline range are converted to fuel values, including high -octane gasoline and middle distillate fractions, in an exceptionally high liquid yield.

DEFINITIONS Because throughout this specification numerous terms will be used to characterize various hydrocarbon charge stocks and fractions thereof, various conversion products, and various characteristics of the aforesaid stocks, fractions and products, such terms will first be defined in order to facilitate understanding of the subsequent description.

In order to avoid thermal cracking, the distillation of petroleum fractions boiling above about 750 F. conventionally is carried out under vacuum. However, throughout the following description the boiling .temperatures given refer to the boiling point at atmospheric pressure, i.e., for uniformity with the boiling points referring to atmospheric pressure distillations, the boiling points referring to vacuum distillations have been corrected to the corresponding boiling points at atmospheric pressure.

The process of the present invention is capable of cnverting an entire petroleum crude hydrocarbon feed to fuel values in an exceptionally high liquid yield, and is particularly effective in so converting a parainic crude, for example, a Mid-Continent, Middle East, Canadian, West Texas and East Texas crude, although it also is applicable to oils recovered from shale, gilsonite, tar sands and the like. The term crude refers to the crude petroleum as recovered from an oil well, after separation therefrom of constituents that are gaseous under recovery conditions at the well, for example constituents such as components which are gaseous at recovery conditions but which are condensed to liquid products.

Upon distillation, various fractions-may be separated from the crude. A fraction which may boil from about C5 to about 180 F. is termed a light gasoline (also known as light naphtha). A fraction which boils from about 180 to about 400 F. is termed a heavy gasoline (also known as heavy naphtha), and is conventionally used as a reformer charge stock.

After the light gasoline and heavy gasoline have been distilled from the crude, the entire remaining portion of the crude and is called a whole topped crude.V

The fractions distilling olf after the gasolines are called distillates or gas oils. It is well known that the term gas oil is a broad, general term that covers a variety of stocks. When used in the following description, the term, unless further modified, includes any fraction distilled from petroleum which has an initial boiling point of at least about 350 F., a 50% point of at least about 475 F. and an end point of at least about 600 F., and boiling substantially continuously between the initial boil- ICC ing point and the end point. The portion of the crude oil which is not distilled is` considered to be a residual stock or residuum. The exact boiling range of a gas oil therefore will be determined by the initial distillation temperature (initial boiling point), the 50% point, and by the temperature at which distillation is cut 0E (end point).

In practice, petroleum distillations have been made under vacuum up to temperatures as high as 1l00 to 1200 F. (corrected to atmospheric pressure). Accordingly, in the broad sense a gas oil is a petroleum fraction which boils substantially continuously between two temperatures that establish a range falling Within from about 3507 to about 1l00 to l200 F., the 50% point being at least about 475 F. Thus, a gas oil could boil over the entire range of about 375 4to 1200 F., or it could boil over a narrower range, for example 500 to 900 F.

The gas oils can be further roughly subdivided by overlapping boiling ranges. Thus, a light gas oil boils between about 375 and 650 F. A medium gas oil boils between about 600 and 750 F. A heavy gas oil boils between about 600 and 900 F. A gas oil boiling between about 800 to 1200 F. is sometimes designated as a vacuum gas oil. It must be understood, however, that a gas oil can overlap the foregoing ranges. It might even span several ranges; for example, it may include both light and medium gas oils.

As heretofore mentioned, a residual stock or residuum is any portion which is not distilled. Therefore, any portion, regardless of its initial boiling point, which includes all the heavy bottoms such as tars, asphalts, etc., is a residual stock. For example, a residual stock can be the portion of the crude remaining undistilled at 1200 F., or, if distillation has not been carried to such a high temperature, it may be the same portion plus a gas oil portion that has not been distilled off.

If desired, the residual portions and/or whole topped crude can be deasphalted by conventional means.

Cycle stocks or cycle oils refer to product fractions from catalytic cracking units and hydrocracking units which boil above the gasoline boiling range, usually between about 400 and about 850 F. A light cycle oil is a cycle oil boiling generally from about 400 to 650 F. A heavy cycle oil is a cycle oil boiling from about 650 to about 850 F. As in the vcase with the gas oils, it will be recognized that the term cycle oil is not an ex tremely precise term, and the boiling ranges given are subject to some variations.

PRIOR ART PROBLEMS (l) General Heretofore in a refinery the crude oil has been lpassed to a distillation unit, normally termed a crude column, and the oil has -been fractionated into various cuts, including light gasoline, heavy gasoline, light gas oil, heavy gas oil, and a residual portion boiling so high as to resist vaporization in the crude column. The residual portion has been withdrawn as bottoms and occasionally has been subjected -to further processing, for example distillation under vacuum, thermal cracking, steam stripping or coking, to recover additional gas oil distillate fractions.

Heretofore in refinery operations conducted for the pur-pose of maximizing the conversion of gasoline, yboth the light and the heavy gas oils, but not the aforesaid residual portion, have been used as a feed to a conventional catalytic cracking unit (cat. cracker), although, in some cases, a thermal cracker has been used either alone or in conjunction with the catalytic cracking unit.

In these conventional relinery operations, the gas oil Si fractions recovered from the crude oil are converted in the catalytic cracking unit in varying measure to fractions boiling in the gasoline range. Thus, light gas oils can be converted to gasoline at a yield of approximately 40%, while the yield from heavy gas oils is approximately 60% these and other percentages given hereinrbeing on the basis of volumes of gasoline produced per volume of feed.

(2) Handling of residual portion from crude column As stated above under (1) General, heretofore the residual portion from the crude column has been withdrawn from the system or, occasionally, has been subjected to such further processing steps as vacuum distillation, thermal cracking, steam stripping or coking.

The residual portion has a high asphaltene content, and usually a high metals content. Accordlingly, heretofore this portion could not be sent to a conventional catalytic cracker, which is adversely affected both by asphaltenes, which cause an inordinate amount of coking of the cracking catalyst, and by metals, which foul the catalyst.

In the prior art processing alternatives to which the residual portion has been subjected, that portion has been converted in a large measure to coke or low value black fuel. When it has been sent to a coker, the products are unstable and are poorer feed stocks for subsequent catalytic cracking due to the nonselective nature of the coking process, even though the carbon-hydrogen ratio of the hydrocarbons has been improved by the formation of coke. From the standpoint of straight run residual stocks, coking may be competitive with solvent deasphalting in the preparation of feed stocks for `subsequent catalytic cracking if there is a market for the coke and complete elimination of black fuels and asphalt is desired. Likewise, coking may be competitive with solvent deasphalting since coker gas oils are generally -of lower metal and asphaltene content then deasphalted oil and may be included in catalytic cracker feed even though they are inferior to virgin stocks. Residual processing schemes, such as vacuum distillation, visbreaking and recycle thermal processing, produce a large percentage of black fuel.

In view of the foregoing, and of major importance to the producer of gasoline and middle distillates, it would be desirable if a practical process wereV available whereby the residual portion from the crude distillation column could be much more completely utilized than heretofore in the production of gasoline and middle distillates.

(3) Maximizing yield of high fuel value liquid products per barrel of crude Heretofore, partially because of the handling of the aforesaid residual portion, even very advanced refinery process combinations have produced considerable quantities of coke and/ or low value liquid products, resulting in a somewhat less than satisfactory yield per barrel of crude of high fuel value liquid products. It would -be desirable if a process were provided. by means of which this yield could be significantly increased. l

(4) Production of heavy naphtha and light naphtha As is well known to those familiar with conventional catalytic cracking andrhydrocracking, the major products of such cracking operations are dry gas, butanes, C5+ light naphtha, heavy naphtha and cycle stock (boiling at temperatures higher than about 390 F). In each case, the light naphtha fraction has a relatively high octane number. On the other hand, the heavy naphtha fraction, particularly that which is obtained by cracking in the presence ofV hydrogen, has an octane number that is generally several numbers lower. Accordingly, inorder to produce a finished gasoline having a relatively high octane number, it has been the practice to blend a heavy naphtha fraction with a light naphtha fraction, and with butanes in amounts limited by the maximum permissible vapor pressure. There is, however, a steadily increasing demand for higher octane gasolines (about 95 and higher). Those skilled in the art will readily appreciate that such octane requirements cannot be met by the aforementioned conventional blending operations. Accordingly, the relatively low octane heavy naphtha fraction has been subjected to reforming operations. As the increasing demand for higher octane gasolines must be satisfied by reforming, instead -of Iby blending, there is a greater demand for heavy naphtha fractions that can be reformed and a correspondingly lesser demand for light naphtha fractions that can be used for blending purposes. The light naphtha fraction usually is not subjected to a reforming operation because it produces excessive amounts of dry gas, coke, etc. The yield of gasoline obtained by reforming light naphtha therefore is prohibitively small. Accordingly, it is very desirable that there be provided a process that will produce greater amounts of heavy naphtha `and lesser amounts of light naphtha.

(5 Production of gasoline product with satisfactory road octane number Heretofore, conventional catalytic cracking processes have been used to convert a portion of the crude feed to heavy gasoline, whichy then has been reformed to increase its octane number. However, the resulting reformate has not had a satisfactory road octane number without additional blending with an alkylate produced from isobutane and olens, and sutiicient quantities of alkylate have not been obtainable in the same process to meet the blending requirements of the catalytic reformate. By definition, road octane is A(F1)+B(F-2)+C, where F-l and F-Z are the conventional octane ratings, and where A, B

and C and constants, and where A is about the same order of magnitude as B. Also by definition, sensitivity is (F-1)-(F-2). Accordingly, a high sensitivity adversely affects road octane number, and conversely, a low sensitivity is helpful to road octane number.

Generally, a catalytically cracked heavy gasoline has a sensitivity of around 11 to 13, and generally the reformate resulting from the reforming of this heavy gasoline has a sensitivity of around 10 to 11.5. Accordingly, to obtain a satisfactory road octane number, the reformate must g be blended lwith alkylate,which has a sensitivity of around -1. Where suticient quantities of alkylate has not been available from the process, the deficiency has had to be made up with alkylate from another source.

In view of the foregoing, it would be desirable if an integrated process were provided which would produce substantial quantities of products having a low sensitivity similar to that of alkylate, to help meet gasoline blending requirements for such low sensitivity materials.

(6) Flexibility in ratio of light gas oil product to gasoline product The objects of the present invention include solutions to the foregoing problems; accordingly, the objects include the following:

(1) It is a general object of this invention to provide an integrated refinery process wherein, in addition to producing gasoline fractions by distillation from the crude hydrocarbon oil, all of the gas oil fractions distilled from that oil as well as the remaining residual portion also can be converted to gasoline fractions of high octane rating.

(2) Another object of the invention is to integrate into a process of this character a particular means for handling said residual portion so that it may be substantially completely utilized in the integrated process for ultimate conversion to gasoline, instead of all or a substantial portion thereof being directed to less valuable uses as has been necessary heretofore.

(3) Another object of the present invention is to provide a process of the foregoing character .with which may be maximized the total liquid yield per barrel of cru'de of high fuel value liquid products.

(4) Another object of the present invention is to provide a process of the foregoing character with which large amounts of heavy naphtha suitable for catalytic reforming may be produced and in which at least a portion of the desired heavy naphtha fraction is produced from a residual portion of a crude petroleum feed stock.

(5) Another object of the present invention is to provide a process capable of producing substantial quantities of products having low sensitivity to help meet gasoline blending requirements for such low sensitivity materials.

(6) Still another object of the present invention is to provide a process with which the ratio of light gas oil product to gasoline product may be varied over a wide range.

STATEMENT OF INVENTION In accordance with a specific embodiment of the present invention, there is provided a process for converting substantially all of a crude hydrocarbon feed to fuel values in an exceptionally high liquid yield, which comprises separating said crude feed into fractions including light gasoline, heavy gasoline, gas oil and residual fractions, converting at least a portion of said gas oil to gasoline in a hydrocracking zone, converting at least a portion of said residual fraction into gasoline and cycle oils in a catalytic cracking zone operating under conditions of low conversion, including `an oil to catalyst contact time below l0 seconds, and converting at least a portion of said cycle oils to gasoline in said hydrocracking zone.

In a more specific embodiment of the present invention, there is provided a process for converting substantially all of a crude hydrocarbon feed to fuel values in an exceptionally high liquid yield, which comprises separating said crude feed into fractions including light gasoline, heavy gasoline, light gas oil, heavy gas oil and residual fractions, recovering said light gasoline as a net product, converting a substantial portion of said light gas oil in a hydrocracking zone to gasoline, converting a substantial portion of said heavy gas oil in said hydrocracking zone to gasoline and cycle oil, converting a substantial portion of said residual fraction in a catalytic cracking zone operating under conditions of low conversion, including an oil to catalyst contact time below l0 seconds, preferably 0.7 to 9 seconds, to gasoline and cycle oil, and converting a substantial portion of said cycle oil from said catalytic cracking zone in said hydrocracking zone to gasoline and cycle oil.

DRAWING The present invention will best be understood, and further objects and advantages thereof will be apparent, from the following detailed description, when read in conjunction with the accompanying drawing, which is a simplified liow diagram illustrating a group of refinery units and interconnecting flow paths suitable for carrying out the process of the invention.

PROCESS UNITS AND OPERATING CONDITIONS, GENERAL The process of the present invention may include, in various combinations, distillation zones, a catalytic cracking zone operating at low conversion, a deasphalting zone, a hydrofining zone and a hydrocracking zone.

Suitable catalysts and operating conditions for these various zones are described immediately below.

DISTILLATION ZONES The necessary distillation zones and operating conditions thereof for the process of the present invention are conventional, from the crude column to the miscellaneous distillation zones operating in conjunction with the various conversion units.

CATALYTIC CRACKING ZONE 'The catalytic cracking zone in the process of the present invention is a catalytic cracking unit operated with a conventional catalytic `cracking catalyst, for example, silica-alumina or silica-magnesia, at a temperture of about 850 `to 1000 F., a pressure of about from 10 to 30 p.s.i.g., a space velocity of about 20 to 50, and an oil to catalyst contact time below 10 seconds, preferably 0.7 to 9 seconds, more preferably 2 to 5 seconds. As will be discussed hereinafter, the catalytic cracking zone operates at a low conversion not only 35% per pass, for example, 30 to 35% per-pass conversion, and, desiraibly, is a transfer line catalytic cracker. Preferably, the transfer line catalytic cracker operates once through with low activity catalyst and at low conversion.

Petroleum residua are the highest boiling portions of the crude petroleum and represent the most diilicult of all petroleum fractions to convert into high quality prodnets. The presence of numerous poisons and contaminants greatly complicates the basic problem which is one of hydrogen deficiency. Total coke production in residnum catalytic cracking can be assumed to be the sum of the carbon residue content of the residuum and the carbon yield from the cracking of a carbon residue-free oil. Petroleum residua can be separated by techniques known in the art into three broad classifications-asphaltenes, resins and oi-ls. The asphaltenes are relatively poor in hydro-gen and rich in sulfur, nitrogen and metallic constituents. Removal of the asphaltenes with their low hydrogen contents and high concentrations of heterocyclic atoms (sulfur, nitrogen, vanadium, nickel and copper) leaves oil and resin fractions which are greatly improved in quality. Although the asphaltenes constitute only about 20% of a l000 F.+ residuum from Middle East crude, their effect on the reaction rates in catalytic cracking is disproportionately large. When present, the asphaltenes apparently control the rapid reaction, possibly by selective adsorption. Since removal of the asphaltenes causes a substantial reduct-ion in the viscosity of the oil being processed, it is also possible that a dilfusional improvement exists. A decrease in viscosity of .the residuum feed would be expected to increase the rate of diffusion both through the bulk liquid and within the pores of the cracking catalyst.

The effect of metallic constituents in the residuum feed vto the catalytic cracker is to gradually build up these constituents on the catalyst, decreasing its ability to convert oil `and increasing the tendency to form gas and coke. The type of meta-l constituent, concentration, and activity of the starting catalytic cracking catalyst all affect the selectivity of the catalyst ,as measured by the carbon-producing factor. Thus, the coke yields in residuum catalytic cracking are dependent on the carbon residue of the feed and the selectivity of the catalyst; and the manner in which these two factors contribute to the total carbon yield is particularly important in residuum catalytic cracking.

It has been observed that the coke contribution by the carbon residue of the feed is independent of the carbonproducing factor of the catalyst. The catalyst selectivity is markedly affected by catalyst age, and selectivity be comes poor as the rate of inventory replacement increases because .a shorter time for metal deactivation is provided. Conversely, if the rate of inventory replacement is too low, selectivity decreases and the product distribution from catalytic cracking residuum is adversely affected. The optimum level of catalyst replacement depends on the metal content of the feed stock and upon the conditions of temperature and steam partial pressure in the cracking unit.

In order to keep within coke burning and throughput limitations on a residuum catalytic cracker, it is desirable to feed a residuum having a ramsbottom coke of 5 to 20 weight percent, preferably to 10 weight percent. For parainic crudes, this will represent a true boiling cut point of the crude of 750 to 1000 F. and a residuum yield to crude of 30 to 8%. In the longer residuum, the effect of the lower cut point will be to decrease the feed partial pressure, thus increasing vaporization and reduce the carbon residue contribution to coke yield. On the shorter residum, the introduction of steam into the reactor will have a marked effect on coke yield. To maintain a constant carbon-producing factor fo-r the catalyst, replacement rates of the order of 0.5 to 2.5 pounds per barrel ,are desirable, depending, of course, on the metal content of the feed stock. It will be recognized by those skilled in the art that many crude oils, as they are Produced, contain varying quantities of suspended or occluded salts such as sodium and magnesium chlorides. It is desirable that these be removed as completely as possible in well-known desalting equipment prior to the crude distillation unit. These materials are strong catalyst poisons and, in addition, contribute to severe corrosion and coking problems in the processing units.

To obtain maximum decarbonization of the residuum `and maximum yieldsV of liquid products suitable for additional processing, it is desirable that the catalytic cracking operations on the residuum be carried out at low conversion in a transfer line-type reactor at high temperature rand high space velocities. The present invention is particularly useful in connection with the production of gas oils for feed to subsequent cracking operations such as hydrocrack-ing for the production of motor gasoline or middle distillates. Transfer line catalytic cracking at low conversions can recover from residual fractions gasoline `and gas oils utilizing a iuidiz'ed catalyst such as a low activity or partially-spent silica-alumina, silica-magnesia or natural aluminum silicate material. The low activity catalyst effects substantial dealkylation or alkylated polycyclic components and subsequent condensation of deyalkylated polycyclics to highly carbonaceous materials. An enhanced yield of distillate stocks is thereby obtained, and thedealkylated polycyclics are removed onto the catalyst as asphaltenes and other highly carbonaceous oilinsoluble materials in addition to the coke which is formed in theprocess.

The low conversion catalytic process may be carried out las a separate catalytic decarbonization process for production of a maximum amount of liquid distillate hydrocarbons to be used as la feed stock for a conventional catalytic cracking or hydrocracking process. Ori the other hand, it may be carried out `as the rst step in a multiple-stage catalytic process utilizing both a catalytic cracking unit and a hyd-roerackingV unit. The desired low conversion cracking can be obtained in a transfer line reactor which is .operated at high space velocity. The lowV conversion .catalytic cracking can be carried lo-ut in a iluidized systemV either in dense or dispersed phase or in a nonfluidized catalytic system which may be either Liked bed or moving bed type. The low conversion cracking operations :and recovery of high yields of distillate stocksV therefrom `are particularly applicable to residua from crude petroleum and which contain .a substantial amount of metallic contaminants and carbon residue, both of which .are known to cause catalyst deactivation on the usual cracking catalyst.

DEASPHALTENG ZONE Where a deasphalting zone is used, it may be any conventional deasphalting zone operating under conventional deasphalting conditions. Suitable deasphalting solvents in accordance with the practice of this invention can include normally gaseous hydrocarbons such as ethane, ethylene, propane, propylene, n-butane, isobutane, n-butylene, isobutylene, pentane, isopentane, pentylenes, and subsequent light hydrocarbons up to and including eight carbon atoms. In addition, the deasphalting solvents may include mixtures of the individual light hydrocarbons referred to above lso that the resulting mixture can have 1an apparent density and properlties for the purpose of the invention similar to those of a CVO? carbon atom paramn-hydrocarbon mixture. Likewise, the light hydrocarbons, either used alone or in mixtures thereof, may be used yalong with minor .amounts Yof additive materials to improve the deasphalting operation or otherwise increase the yield and quality of the deasphalted oil and/or the recovered asphalt. The deasphalting operation is carried out in suitable deasphalting equipment such 4as a oountercurrent tower equipment with baffi-es or rotating disk conatctors. The deasphalting operation is carried out at any suitable deasphalting temperature and pressure, the temperature and pressure lbeing adjusted so as to maintainthe deasphalting solvent in the liquid phase during the deasphalting operation. The deasphalting temperature in the Irange of 150 to 325 F., usually not more than about 50 F. below the critical temperature of the deasphalting solvent. Pressures in the range of 300 to 800 p.s.i.g. are employed depending on the composition of the defasphalting solvent and lto a minor extent upon the composition of the residua undergoing deasphalting. Generally, deasphalting solvent to residuuin volume ratio in the range of 10:2 is employed. The deasphalter may be operated isotherm-ally or under a temperature gradient.

HYDROFINING ZONE While the invention can be practiced with utility in connection with feeds to the hydrocracking zone containing relatively large quantities of nitrogen, the operation becomes much more economical with stocks containing less than 200 ppm., preferably below p.-p.m., `and much more preferably, below 10 ppm., of nitrogen. A reduction in feed nitrogen level permits the hydrocnacking reaction to be conduced at lower temperatures than with feeds containing relatively large .amounts of nitrogen compounds. Therefore, in the case of hydrocracking zone feed stocks which lare not inherently low in nitrogen, acceptable levels can be reached lby pretreating the feed to the catalytic cracking unit by a nitrogen compound extraction process, or by contacting either tlie catalytic cracking unit feed or, preferably, the particular feed to the hydrocracking process, with hydrogen in the presence of a suitable catalyst at elevated temperatures and pressures to remove nitrogen compounds therefrom. A particularly effective catalyst for removing nitrogen by hydrogenation is one wherein la coprecipitated molybdena-.alumina material (e.g., one prepared in accordance with the disclosure of US. Patent 2,432,286 to Claussen et al. or U.S. Patent 2,697,006 to Sieg) is combined with cobalt oxide, the nal catalyst having a metals content equivalent to about 2% cobalt `and 7% molybdena. Representative processing conditions for removing nitrogen with this catalyst are Ian LHSV of 1 to 3, 700 to 800 F., 200 to 2500 psig. an-d 1000 to 15,000 SCF. of of hydrogen per barrel of lfeed stock.

When nitrogen removal is effected by hydrofining, the resulting elliuent is treated, in accordance with methods presently known in the art, so as to remove ammonia [and some hydrogen .sulfide which may be present. A preferred removal method involves injecting Water into the total elfluent from the hydrofining unit and then passing the resulting mixture into a high pressure separator operating under such conditions yof temperature and pressure (for example, 100 F. and 950 p.s.i.g.) that a gaseous overhead is removed that is predominantly hydrogen, but which normally contains some hydrogen suliide and light hydrocarbons. This overhead (following :a clean-up treatment to remove any nitrogen and sulfur-containing compounds, if desired) can be recycled to the hydrofining unit along with make-up hydrogen. Two liquid phases are formed in the separator, an upper hydrocarbon phase and .a lower aqueous phase which contains essentially all of the ammonia present and some hydrogen sulfide in the form of ammonium sulfide. The aqueous phase is removed from the system and discarded.

The hydrocarbon layer is then preferably passed into .a stripper or `distillation column from which lany remaining hydrogen sulfide, ammonia :and Water are removed overhead. The stream may `also be freed of any light yhydrocarbon 4fractions (boiling in the gasoline range or below) formed as la'result of hydrocracking reactions taking place over the hydrotining catalyst.

lHYDROCRACKING ZONE The portion 4of the denitr-iiied feed to be hydrocracked, along with `from :about 1500 to 30,000, and preferably from Iabout 3000 to 15,000, `standard cubic feet (S.C.F.) of hydrogen per barrel of total reaction feed, is passed into the hydrocracking zone at ya liquid Ahourly space velocity (LI-ISV) of .from about 0.2 to 15, and preferably from about 0.4 to 3.0, .and intimately contacted with the catalyst; Y

The contacting step is conducted under a pressure of at least 500 p.s.i.g., and preferably from about 800 to 3000 p.s.i.g. The temperature is preferably maintained in the range of from about 400 to 750 F. because at temperatures above about 750 to 800 F. the amount of gasoline product lost to the less desirable C3 and lighter materials rapidly increases, thus lowering the motor fuel yield. For example, it has been found that the amount of methane produced at 800 F. per unit of converted product is approximately sixteen times as great as that formed at 700 F., and four times as great as that produced at 750 F. At higher temperatures, the situation becomes much worse. Accordingly, resort is normally had to temperatures above about 750 F. up to about 850 F. only in the last stages of the catalyst on-stream period when it is desired to maintain relatively high activity at the expense of higher light gas losses or, in the case when the relatively high nitrogen-containing feeds are processed. Further, operations at temperatures above about 750 F. and at low or moderate pressures induce a relatively rapidv decrease in the activity of the catalyst as reflected by reduced per-pass conversion levels. Thus, when operating at 875 F. and at a relatively low pressure, for example 1500 p.s.i.g., on a hydroiined light cycle oil feed, regeneration of a cobalt-molybdenum on silicaalumina catalyst is required in most instances after onstream periods of one day or less, which contrasts with on-stream periods of 100 to 300 or more hours at good activity as temperatures are maintained below about 825 F. With operation at 800 F. and higher with the same and similar feeds, but with nickel sulfide on silica-alumina, regeneration is required after on-stream periods of a few hundred hours or less, compared with operation below 700 F., with which can be obtained on-stream periods of several thousand hours without regeneration. In the present process, it is recommended that the reaction be conducted at an initial on-stream temperaturefrom about 550 to 650 F., with a progressive increase to about 750 to 850 F. so as to maintain catalyst activity at a controlled level. This initial and terminal temperatures will vary, with character of feed and catalyst, within the overall range specified above.

The catalyst employed in the hydrocracking zone is one wherein a material having hydrogenating-dehydrogenating activity is deposited or otherwise disposed on an active cracking catalyst support.- The cracking component may comprise any one or more of such acidic materials as silica-alumina, silica-magnesia, silica-alumina-zirconia composites, alumina-boria, iluorided composites, and the like, 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 Groups VI, VII and VIII metals, as well as the oxides and sulfides thereof, alone or together with promoters and stabilizers that may have by themselves small catalytic effect, representative materials being the oxides and suldes of molybdenum, tungsten, vanadium, chromium and the like, as well as of metals such as iron, nickel, cobalt and platinum. If desired, more than one hydrogenating-dehydrogenating component can be present, and good results may be obtained with catalysts containing composites of two or more of the oxides of molybdenum, cobalt, chromium, tin and zinc, and with mixtures of said oxides with uorine. The amount of the hydrogenating-dehydrogenating component present can be varied with relatively wide limits of from about 0.5 to 30% based on the weight of the entire catalyst.

Exemplary 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, (c) mixtures of from about 0.5 to 10% each of cobalt oxide and chromium oxide, (d) 0.1 to 10% nickel, nickel oxide or nickel sulfide, (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, 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 l() 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. Nickel sulfide and/or cobalt sulfide are especially suitable. The entire preferred catalyst composite and conditions of Scott U.S. Patent 2,944,006 will be especially suitable in the process of the present invention. Under favorable operating conditions, the hydrocracking catalyst will maintain high activity over periods of 50 to 300 or more hours. The activity of the used catalyst can then be increased, if desired, by a conventional regeneration treatment involving burning off catalyst contaminants with an oxygen-containing gas.

The light gasoline produced in the hydrocracking zone has a very low sensitivity, and is well suited for blending with catalytic reformate having an unsatisfactory road octane number, to produce a blend of higher road octane number than the reformate.

DETAILED DESCRIPTION A best mode for carrying out the process of the present invention may be determined by reference to the appended drawing which is a diagrammatic illustration of a group of interrelated reiinery units and flow paths suitable for use in practicing the process. For purposes of clarity and because their location and use will be readily apparent to those skilled in the art, various pieces of conventional equipment, such as heaters and pumps, have been omitted from the drawing. The followingdetailed description will indicate how the process of the present invention may be operated to maximize gasoline production. With this description as a guide, those skilled in the art will understand that the operation easily may be va- 1 1 ried to maximize middle distillate production, mainly by simply tak-ing as products the middle distillate range materials, such as light cycle oils, from the various process units, rather than further processing them to produce gasoline, and by adjusting the recycle cut points to the various process units as desired.

A crude oil feed is supplied through line 1 to crude column 2. Desirably, the crude oil feed is a paraftinic crude, for example, a Mid-Continent, Middle East, Canadian, West Texas or East Texas crude. In crude column 2 the crude is stripped of a light gasoline fraction which is removed as a product through line 3, a heavy gasoline fraction which is passed through line 4 to catalytic reformer 5, a light gas oil fraction which is passed through line to hydrofining zone 11, and a heavy gas oil fraction, which may boil from about 550 to 800 F., which may be passed through line 12 to hydroning zone 11. These light and heavy gas oil fractions need not be thus passed to hydroning zone 11 Where they are sufficiently low in nitrogen content to be passed directly to hydrocracking zone 13. They may be passed directly to hydrocracking zone 13, by lines not shown, if they have a nitrogen content below about 500 ppm.; however, if they have a nitrogen content above about ppm., it is preferable that they iirst be hydrofined in zone 11. Metal contaminants and coke precursors in the crude oil feed tend to concentrate in the higher boiling portions thereof. Accordingly, crude column 2 desirably is operated to limit to about 800 F. the end point of the heavy straight run gas oil in line 12 going to hydroning zone 11; by operating in this manner, contamination of the catalyst in hydrocracker 13 by metals and coking of the catalyst in hydroning zone 11 will be at a slower rate than if the heavy straight run gas oil had a higher end point, and an adequate supply of heavy straight run gas oil for feeding to these zones still will be possible. If desired, a portion of the light gas oil from line 10 may be withdrawn through line 10A, for example for use as a No. 2 oil. If desired, a portion of the heavy gas oil from line 12 may be withdrawn through line 12A as a product.

With the crude thus stripped to about 800 F., a residuum comprising about of the original crude feed remains, and is passed through line 14 to solvent deasphalting zone 15, which may be a conventional propane deasphalting zone, and thence through line 16 to transfer line catalytic cracker 17. It is desirable to include in the residual stock in line 16 to transfer line catalytic cracker 17 all portions of the crude feed boiling as low as 800 F.; if the residual stock includes only higher boiling materials, for example only materials boiling above 900 F., there will be a greater percentage of carbon residue in the stock. This residue will appear as increased production of coke on the catalysts of zones 11 and 13, and those units accordingly will have the activity of their catalysts reduced.

Transfer line catalytic cracker 17 is operated at a low conversion not over per pass, for example 30 to 35% per pass, with a low activity silica-alumina catalyst. Catalytic cracker 17 also is operated at a space velocity of about 20 to 50, and an oil to catalyst contact time of below 10 seconds, preferably 0.7 to 9 seconds. More preferably, this oil to catalyst contact time, or residence time in the reactor, is from 2 to 8 seconds. In catalytic cracker 17 residuum is cracked or decarbonized in concurrent ow upwardly with catalyst circulating through line 20, catalyst and oil separator 21, line 22, conventional catalyst regeneration zone 23 where coke is burned from the catalyst, and line 24. Periodically, a portion of the regenerated catalyst in transfer line catalytic cracker 17 may be withdrawn and treated in a conventional demetallation zone, not shown, to remove from the catalyst the accumulation of metals resulting from processing of the feed stock in zone 17; the thus-treated catalyst then may be returned to catalytic'cracker 17. Even the heavy cycle oil, the heaviest product from catalytic cracker 17,

does not contain any significant amount of asphaltenes or metals, both of which are removed from the feed and deposited on the catalyst of this unit. Catalytic cracker 17 does not produce low value residual black fuel from the residualrfeed, but produces only distillate products. The reactor temperature and pressure may be, for example, 1000 F. and 15 p.s.i.g., respectively. The reactor may be operated with a catalyst to oil ratio of about 8.0. Under these conditions the low conversion may be obtained by operating the reactor with a space velocity above about 25.

In catalyst and oil separator 21, the catalyst is separated from the oil product and passed through line 22 to regenerator 23, and the oil product is passed through line 25 to distillation column Z6, where it is separated into various fractions. A C3= and C4: fraction is produced in small amounts and is withdrawn through line 27 for use, for example, in gasoline pressuring. A light gasoline fraction is withdrawn as a product through lines 29 and 3. A heavy gasoline fraction may be passed through lines 35 and 4 to catalytic reformer 5. However, because of the high octane number of this fraction, it may be desired to withdraw all or a portion of it as a product, instead of reforming it, particularly where it is not necessary to hydroiine it to remove such contaminants as sulfur, nitrogen, dioleiins and other gum-forming precursors. Hydrofining will tend to reduce the outane number and necessitate reforming. A light cycle oil fraction. which would be refractory to catalytic cracking, is passed through line 40 to hydroner 11. If desired, a portion of the light cycle oil may be withdrawn through line 40A as a product. A heavy cycle oil may be passed through line 41 to hydroner 11.

Hydrotiner 11 is supplied with a light gas oil through line 10, a heavy straight run gas oil through line 12, a light cycle oil through line 40 and a heavy cycle oil through line 41. Frequently, it will be found that the nitrogen content of the light and heavy gas oil in lines 10 and 12, respectively, is sufficiently low to warrant passing these gas oils directly to hydrocracking zone 13; however, it generally will be found desirable to pass the light and heavy cycle oils in lines 40 and 41, respectively, to hydroiining zone 11, as shown, rather than directly to hydrocracking zone 13, although the latter operation is possible. The effluent from hydroiining zone 11 is passed through line Sti to hydrocracking zone 13, after being freed of ammonia produced in hydrofning zone 11. If desired, other products of the reaction in hydroiining zone 11, for example gasoline produced by virtue of the hydrocracking activity of the hydroiining catalyst, may be Withdrawn from the efliuent from hydroiining zone 11 prior to passage of the remainder of that effluent to hydrocracking Zone 13.

The reaction products from hydrocracking zone 13 are passed through line 511 to distillation column where they `are separated into various fractions.

From distillation column '75, a light gasoline fraction is recovered through lines 81 and 3 as a product. A heavy gasoline fraction is passed through lines 82 and 4 to catalytic reformer 5. A light cycle oil is passed through line 83, and thence through either line 83A as a product `or through lines 83B and 84 to hydrocracker 13, as desired. A heavy cycle oil is passed through line 86 and thence through either line 36A as a product or through lines 86B and `84 to hydhocracker 13 as desired.

In catalytic reformer 5, ya V'heavy straight run gasoline fraction from crude column V2, a heavy gasoline from catalytic cracker 17, and the heavy gasoline fraction `from hydrocracker 13 are reformed under conventional reforming conditions to produce a high octaine reformate which is withdrawn fromV catalytic reformer 5 through line 85.

EXAMPLES The following example illustrates the processing of 50,000 barrels per day of`a Middle East crude oil to produce -a maximum yield of gasoline, using a conventional rerlnery arrangement including delayed coking, catalytic cracking, C4 alkyl-ation and catalytic reforming, compared with the processing of the same daily quantity of the same feed in a refinery arrangement in accordance with the present linvention, also to produce a maximum yield of gasoline.

EXAMPLE 1 [Maximizing Gasoline] Conventional Present invention Processing arrangement Case A Case B Delayed coking Transfer line catalytic Catalytic cracking cracking C4 alkylation Hydrocracking Catalytic reforming Catalytic reforming Hydrogen manufacture Refinery feed:

34.5 API Middle B/D B/D i-Butane 1j soo :22221212212222122112:

Total 53, 300 50, 000

Products:

Net Ca--EFO1 6, 580 2, 316 Excess lsobutane. 4, 335 Motor gasoline, 101b.

RVP 2 37, 295 51,850 95. 2 96. 0 F-2-i-3 88. 9 90. 3 Road Octane. 95.0 96. 5 ASTM 50% temperature, F 200 218 Light catalytic cycle oil 8, 800 coke- 8 300 l EFO (Equivalent fuel oil) is amount of 10 API Bunker fuel oil that would have equivalent heating value in B.t.u., assuming 1 barrel of said fuel oil has heating value of 6.3M B.t.u.

2 Reid vapor pressure.

3 Tons/ EXAMPLE 2 [Maximizing Middle Distillates] Conventional Present Invention Processing arrangement (Same as (Same as Example 1) Example l) Refinery Feed:

Refinery Input- B/D 34.5 API Middle East crude 50, 000 50,000 i-Butaue l, 000

Total 51,000 50,000

Products:

21, 800 21, 850 95. 2 95. 2 89. 0 90. 3 95.0 96. 0 Middle distillate- 20, 250 32, 500 Middle distillate] gasoline 0. 9 1. 50 65 1 Reid vapor pressure.

p gas oils of low quality; (c) complete elimination of residual fuel oil, including low quality coke; (d) maximizing production of heavy naphtha Ifor upgrading into high octane aromatic reformates; (e) improvement in gasoline pool road octane by substituting isoparaiiinic gasoline for olefinic compounds; and (f) ability to vary middle distillate/gasoline ratio between 0 and 1.50 compared to 0.23 and 0.9 for conventional refinery processlng.

In the process of the present invention, it will be noted that various fractions, for example the heavy cycle oil fraction in line 41, are processed in hydrocracking zone 13, rather than a catalytic cracking zone as in conventional refinery practice, Accordingly, the total liquid product yield in the process of the present invention, because of such additional hydrogen processing, is higher .than when a conventional catalytic cracker is incorporated in the processing scheme. With the process of the present invention, a total liquid yield of over based on the crude oil feed, may be obtained. Further, with the process of .the present invention, the middle distillate to gasoline ratio may be greater than in processing schemes that utilize a conventional catalytic cracking zone, which reduces the possible yield of middle distiillates. Still further, the middle distillate product produced by the process of lthe present invention is of a generally better quality than the middle distillate prodduct produced by the refinery scheme employing a conventional catalytic cracking zone, because most of the middle distillate produced by the process of the present invention is produced in the hydrocracking zone and not in a catalytic cracking zone. It is Well known that light cycle oil from a catalytic cracker is a rather low quality middle distillate, because it has a high aromatics content.

SUMMARY From the foregoing detailed description, it may be seen that the process of the present invention operates in an extremely efficient manner to utilize the entire spectrum of the crude oil feed in making only (liquid products of high fuel values in high liquid yields, with a minimum installation of expensive reiinery equipment. Such an integrated refinery, capable of converting all oi the crude feed -to liquid products of high value with minimum waste and los-ses, long has been a goal desired by reliners.

It will be apparent that the net iiquid product streams shown easily may be `utilized by the refinery as fuels, in .blending operations, or subjected to further processing, as desired. AAny conventional blending operations may be used in connection with the process of the invention. For example, in a conventional gasoline blending zone, not shown in the drawing, the high octane reformate from catalytic reformer 5 may be blended with butanes land wit-h C5+ light naphtha, in suitable .proportions to produce a finished gasoline ot the desired Reid vapor pressure and octane number.

Although only specific modes of operation of the Iprocess of the present invention have been described, numerous variations could be made in those modes Without departing from the spirit of the invention, and all such variations that fall Within the scope of the appended claim are intended to be embraced thereby.

I claim:

In a process for converting substantially all of a crude hydrocarbon -feed to fuel values in high liquid yield, which comprises separating said crude into fractions including a light gasoline fraction, a heavy gasoline fraction, a gas oil fraction and a residual fraction, converting at least a portion of said gas oil fraction to gasoline in a hydrocracking Zone and converting at least a portion of said residual fraction to gasoline and catalytic cycle oil in a catalytic cracking zone operating at low conversion, the improvement which comprises:

(a) passing substantially all of said gas -o-il fraction to said hydrocracking zone;

'l5 1.6 (b) operating said catalytic `cracking zone at an oil References Cited by the Examiner lto catalyst Contact time below 10 seconds; UNITED STATES PATENTS (c) passing at least a portion of said catalytic cycle oil to said hydro-cracking zone; and ggheh x l y. c (d) recovering from said hydrocrac ung zone a g'lso 5 l 3,072,560 1/63 Paterson et al. 8 4 10X line product derived in part lfrom said gas oil fracti-on and in part from said catalytic cycle oil. ALPHONSO D. SULLIVAN, Primary Examiner-

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