Desulfurization of asphaltene-containing hydrocarbonaceous black oils

[57] ABSTRACT

[52] U.S. CL....................... 208/86, 208/61, 208/78,

208/80

[51] Int. Cl.............................................ClOg 13/00

[58] Field of Search.................... 208/86, 61, 78, 80

[56] References Cited

UNITED STATES PATENTS

2,002,004 5/1935 Gard..................................... 208/14

Sulfurous, asphaltene-containing black oils are converted into desulfurized, lower-boiling hydrocarbon products. The process involves a combination of solvent deasphalting, thermal cracking and multiple-stage hydrocracking.

Claims, 1 Drawing Figure

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FATENTE08GV27 1973

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DESULFURIZATION OF ASPHALTENE-CONTAINING

HYDROCARBONACEOUS BLACK OILS

APPLICABILITY OF INVENTION

The invention described herein is applicable to a process for the conversion of petroleum crude oil, and the heavier fractions derived therefrom, into desulfurized, lower-boiling hydrocarbon products. More specifically, my invention is directed toward a process for convert- 1 ing atmospheric tower bottoms, vacuum tower bottoms (vacuum residuum), crude oil residuum, topped crude oils, oils extracted from tar sands, etc., all of which are commonly referred to in the art as “black oils,” and which contain a significant quantity of asphaltic material and sulfurous compounds.

Petroleum crude oils, particularly the heavier oils extracted from tar sands, topped or reduced crudes, and vacuum residuum, contain high molecular weight sulfurous compounds in exceedingly large quantities, nitrogenous compounds, high molecular weight organometallic complexes (principally containing nickel and vanadium) and light hydrocarbon-insoluble material. The latter is principally found to be complexed with sulfur and, to a significant extent, with the metallic contaminants. In this regard, “black oils” differ considerably from heavy gas oils which are not so severely contaminated, and which normally do not have as high a boiling range. In the petroleum refining art, a black oil is generally characterized as a heavy hydrocarbonaceous material of which more than about 10.0 percent (by volume) boils above a temperature of 1,050°F., and which has a gravity less than about 20.0 °API. Sulfur concentrations are exceedingly high, being more than 1.0 percent by weight, and often in excess of 3.0 percent by weight. There currently exists an abundant supply of such hydrocarbonaceous material; however, the utilization thereof, as a source of more valuable distillable liquid hydrocarbon products, is virtually precluded by present-day catalytic reaction techniques.

Knowledgeable experts are currently predicting a world-wide “energy” crisis in the not-too-distant future. Those possessing expertise in the field of petroleum exploration, for example, are very much concerned with the ever-dwindling reserve supply of natural gas in comparison to the ever-increasing demand thereof. As a result of legislation being imposed upon the sulfur content of fuel oils burned to meet certain energy requirement, more and more “energy” suppliers are looking to natural gas as a substitute. Several processes are presently being proposed which, it is hoped, will alleviate the forthcoming critical shortage of natural gas. These processes primarily involve the conversion of naphtha fractions, via steam reforming and shift methanation, into a “synthetic” natural gas. However, this in turn creates a shortage of naphtha boiling range material for ultimate utilization as motor fuel, particularly with the advent of the need to produce “clear” gasolines to avoid severe atmospheric pollution as the result of metal-containing motor fuel additives. Likewise, a shortage of kerosene boiling range fractions, principally employed as jet fuels, as well as gas oils, will stem from the necessity to convert such charge stocks into suitable automotive motor fuel. A multitude of factors are, therefore, contributing to the developing energy crisis. Processing technology is required to insure the utilization of virtually 100.0 per10 cent of petroleum crude oil charge stocks. In the petroleum refining art, this is commonly referred to as “the bottom of the barrel.”

The process encompassed by the present invention supplies at least some of the technology required to effect the catalytic conversion of hydrocarbonaceous black oils into distillable hydrocarbons in volumetric yields exceeding 100.0 percent. Specific examples of the charge stocks to which the present scheme is adaptable, include a vacuum tower bottoms having a gravity of 7.1 °API, and containing 4.05 percent by weight of sulfur and 23.7 percent by weight of asphaltenic material; a “topped” crude oil having a gravity of 11.0 °API, and containing 10.1 percent by weight of asphalts and 5.2 percent by weight of sulfur; a vacuum residuum having a gravity of 8.8 “API, and containing 3.0 percent by weight of sulfur and 4,300 ppm. of nitrogen; a vacuum bottoms having a gravity of 5.4 °API, and containing 6.15 percent by weight of sulfur, 233 ppm, by weight of metallic contaminants and 12.8 percent by weight of heptane-insoluble asphaltic material; and, a reduced crude having a gravity of 11.5 “API, and containing 4.2 percent by weight of sulfur, 3,400 ppm. of nitrogen, 166 ppm. of metals and 8.6 percent by weight of asphaltenic material.

The paramount difficulty, heretofore encountered with fixed-bed catalytic systems, has been the lack of catalyst stability when processing at those conditions required to convert the sulfurous compounds into hydrogen sulfide and hydrocarbons. At the operating severity required to achieve acceptable desulfurization, the asphaltic material, finally dispersed within the black oil, has the tendancy to flocculate and polymerize, and thus become deposited upon the catalytically active surfaces of the catalyst. Furthermore, the metallic contaminants filter into the internal cavities, or pores of the catalyst and effectively shield active catalytic sites from the material being processed. In addiI tion to fixed-bed, vapor-phase hydrocracking, another attempted approach has been liquid-phase hydrogenation. In this type of process, liquid-phase oil is passed upwardly, in admixture with hydrogen, into a fluidizedfixed bed of subdivided catalyst; although perhaps ef> fective in removing at least a portion of the organometallic complexes, this type process is ineffective with respect to the insoluble asphalts. Since they are finally dispersed within the oil, the probability of effecting simultaneous contact between the catalyst particle, the asphaltic material and the hydrogen required to prevent coke deposition is remote.

An integral part of the present combination process constitutes the removal of asphaltic material prior to effecting the fixed-bed catalytic conversion of the charge stock. However, the separation of the metalcontaining asphaltic pitch is accomplished in a manner which retains a convertible resin concentrate subsequently processed in a manner significantly increasing the volumetric yield of more valuable distillable hydrocarbon products.

OBJECTS AND EMBODIMENTS

A primary object of my invention is to provide a process for effecting the conversion of hydrocarbonaceous black oils. A corollary objective is to afford maximum yields of substantially desulfurized, distillable hydrocarbon products.

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Another object of my invention is to increase the acceptable effective life of catalytic composites utilized in fixed-bed processing of asphaltene-containing, sulfurous petroleum feed stocks.

Therefore, in one embodiment, my invention involves a process for the conversion of a sulfurous, asphaltene-containing hydrocarbonaceous charge stock to produce lower-boiling, desulfurized hydrocarbon products, which process comprises the steps of: (a) deasphalting said charge stock with a selective solvent, in a first solvent extraction zone, at extraction conditions selected to provide a solvent-lean asphaltic pitch and a solvent-rich, deasphalted first liquid phase; (b) deresining at least a portion of said first liquid phase with a selective solvent, in a second solvent extraction zone, at extraction conditions selected to provide a solvent-lean resin concentrate and a solvent-rich second liquid phase; (c) reacting at least a portion of said resin concentrate with hydrogen, in a catalytic first reaction zone, at hydrocracking conditions selected to convert resins into lower-boiling hydrocarbons; (d) further reacting at least a portion of the resulting first reaction zone effluent in a non-catalytic second reaction zone, at thermal cracking conditions selected to produce additional lower-boiling hydrocarbons; (e) reacting at least a portion of the resulting thermally cracked product effluent and at least a portion of said second liquid phase, in a catalytic third reaction zone with hydrogen, at hydrocracking conditions selected to produce additional lower-boiling hydrocarbons; and, (f) recovering said lower-boiling, desulfurized hydrocarbon products from the resulting third reaction zone effluent.

Other objects and embodiments of my invention relate to additional details regarding preferred catalytic ingredients, the concentration of components within the catalytic composites, preferred processing techniques and similar particulars which are hereinafter given in the following, more detailed summary of the present invention and the combination process encompassed thereby. In one such other embodiment, the first reaction zone effluent is separated, in a first separation zone, at substantially the same temperature and pressure, to provide a first principally vaporous phase and a third liquid phase, a portion of said third liquid phase being reacted in said non-catalytic second reaction zone. In another such embodiment, the second reaction zone effluent is separated, in a second separation zone, at substantially the same temperature and at a reduced pressure in the range of from subatmospheric to about 200 psig., to provide a second principally vaporous phase and a heavy resin by-product.

In still another embodiment, the extraction conditions in said second extraction zone include a higher temperature than that in said first extraction zone.

SUMMARY OF INVENTION

The combination process, encompassed by the present inventive concept, utilizes two solvent extraction zones for the purpose of (1) deasphalting the charge stock to remove an asphaltic pitch and (2) deresining the deasphalted oil to recover a resin concentrate, and to provide a deresined, deasphalted oil. A first fixedbed catalytic reaction zone is utilized to process the resin concentrate to produce lower-boiling hydrocarbon products therefrom. At least a portion of the resin concentrate product effluent is subjected to theimal cracking, after which heavy resins are removed as a byproduct and the remainder is combined with the deasphalted, deresined oil resulting from the second solvent extraction zone. This mixture is then subjected to hydrocracking in a second fixed-bed catalytic reaction 5 zone in order to produce additional lower-boiling hydrocarbon products. Desulfurization is effected to some extent in the first fixed-bed catalytic reaction zone, and is completed in the second fixed-bed reaction zone to the extent that the normally liquid hydrocarbon 10 products are substantially sulfur-free. The inventive concept, upon which the present combination process is founded, stems from recognition that black oils, of the type hereinbefore described, contain a convertible resin concentrate in addition to an asphaltic fraction.

Therefore, the first solvent extraction zone functions to reject the asphaltic pitch while maintaining the resin concentrate in the solvent-rich deasphalted oil (DAO) phase. The latter is subjected to solvent extraction, preferably at a higher temperature, to recover sepa20 rately the resin concentrate, and to provide a deresined oil (DRO). In this manner, the asphaltic pitch is removed from the charge stock prior to processing over the fixed-bed catalytic composites, and the resin concentrate can be processed at an operating severity re25 quired to produce lower-boiling hydrocarbons without incurring adverse effects with respect to the remainder of the deasphalted oil. By the same token, the deasphalted oil can be processed at conditions conducive to the production of normally liquid hydrocarbons in the 30 absence of adverse effects stemming from the presence of the resinous material. As hereinafter set forth, the resin concentrate will generally be processed at a lower severity level than that imposed on the reaction system processing the deasphalted oil in admixture with other ³³ internally produced streams, the sources of which are hereinafter set forth.

The present combination process utilizes two solvent extraction zones to precipitate an asphaltic concentrate and to provide a resin concentrate for subsequent pro⁴⁰ cessing. It must necessarily be acknowledged that the prior art is replete with a wide spectrum of techniques employed for effecting solvent deasphalting of asphaltene-containing hydrocarbonaceous charge stocks. It is understood, therefore, that no attempt is herein made ⁴⁵ to claim solvent deasphalting other than as it is employed as an integral element of the present combination process. Any suitable solvent deasphalting technique known in the prior art may be employed, several examples of which are to be found in the references ⁵⁰ hereinafter briefly described. In the interest of brevity, no attempt will be made to delineate exhaustively the solvent deasphalting art.

Exemplary of such prior art is U. S. Pat. No. 1,948,296 (Class 208-4) which discloses a process for obtaining a particularly good road-type asphalt product. The separated asphaltic fraction is admixed with a suitable oil (lubricating oil, gas oil, etc.) and subjected to oxidation. For effecting the precipitation of the asphaltic fraction, suitable solvents are described as including light petroleum hydrocarbon mixtures such as naphtha, light petroleum fractions comprising propane, n-butane and isobutane, certain alcohols, ether and mixtures thereof, etc. In U. S. Pat. No. 2,101,308 (Class 208-309) similar solvents are utilized to precipitate only a portion of the asphaltic fraction. A solvent of an altogether different character, for example liquid sulfur dioxide, is utilized to separate the resulting ex3,775,293 tract into an asphalt-containing, non-paraffinoid hydrocarbon product.

U.S. Pat. No. 2,002,004 (Class 208-14) involves a two-stage deasphalting process with intermediate distillation. The second stage completes the precipitation of : asphalts onlv partially effected in the first stage. As noted previously, the described solvents include naphtha, casinghead gasoline, and liquefied normally gaseous hydrocarbons such as ethane, propane, butanes, and mixtures thereof. ¹

U.S. Pat. No. 2,914,457 (Class 208-79) describes a multiple combination process involving fractionation, vacuum distillation, solvent deasphalting, hydrogenation and catalytic reforming. Again, suitable liquid deasphalting solvents include liquefied normally gaseous hydrocarbons such as propane, n-butane, isobutane, as well as ethane, ethylene, propylene, n-butylene, isobutylene, pentane, isopentane and mixtures thereof.

Identical solvents, utilized to precipitate asphaltics, are disclosed in U.S. Pat. No. 2,587,643 (Class 208-309). These are, however utilized in admixture with an organic carbonate.

Conspicuously absent from such prior art is a recognition of the difference between the asphaltic fraction and the convertible resin concentrate. Certainly there is found no awareness that the resin concentrate can be processed in a fixed-bed catalytic reaction system to produce more valuable distillable hydrocarbons of lower sulfur content. The multiple solvent extraction zones of· the present combination process recover a convertible resin concetrate as an essential feature of the present invention, and the prior are is silent with respect to this operating technique.

Although both solvent extraction zones will function at generally the same operating conditions, a preferred technique involves precipitating the asphaltic pitch in the first solvent extraction zone at a lower temperature than is utilized to recover a resin concentrate in the second solvent extraction zone. Suitable extraction conditions include a temperature in the range of about 50°F. to about 600°F., and preferably from about 100°F. to about 400°F.; the pressure will be maintained within the range of about 100 to about 1,000 psig., and preferably from about 200 to about 600 psig. The precise operating conditions will generally depend upon the physical characteristics of the charge stock as well as the selected solvent. In general the temperature and pressure are selected to maintain the solvent extraction operations in liquid phase and, with respect to the first extraction zone, to insure that substantially all the asphaltic pitch is removed in the solvent-lean heavy phase with the resin concentrate being retained in the solvent-rich deasphalted oil phase. Suitable solvents, for utilization in the present combination process, include those hereinbefore described with respect to prior art deasphalting techniques. Thus, it is contemplated that the solvent will be selected from the group of light hydrocarbons such as ethane, methane, propane, butane, isobutane, pentane, isopentane, neopentane, hexane, isohexane, heptane, the monoolefinic counter-parts thereof, etc. Furthermore, the solvent may be a normally liquid naphtha fraction containing hydrocarbons having from about five to about 14 carbon atoms per molecule, and preferably a naph- ₆₅tha fraction having an end boiling point below about 200° F. The solvent-rich normally liquid phase is generally introduced into a suitable solvent recovery system, the design and techniques of which are thoroughly described in the prior art.

The terms, “deasphalting” and “deresining,” as employed in this specification and the appended claims, connote the rejection of an asphaltic pitch and, subsequently, a resin concentrate. The precise chemical and physical nature of these two fractions is largely dependent upon the origin of the crude oil and the conditions utilized in the extraction zones, the latter including the nature of the selective solvent. As currently practiced, solvent deasphalting generally refers to a one-stage precipitation operation as applied to an asphalt-containing residuum, whereas deresining, as in the present process, refers to a similar treatment performed on an es15 sentially asphalt-free residuum. Regardless of their precise nature, deasphalting and deresining apply to the rejection of two contiguous bottoms fractions.

In accordance with the present invention, the deresining operation involves the use of a greater solvent/oil 20 volumetric ratio and a higher temperature. For example, in propane deasphalting, a solvent/oil ratio ^of about 6.0:1.0 and a temperatuere of 100°F. to 160°F. are typical. For the deresining operation, a propane/oil ratio of about 10.0:1.0 and a temperature of about 25 140°F. to about 180°F. are typical.

The asphaltic pitch, obtained from the deasphalting operation will exhibit an average molecular weight in the range of 3,000 to 6,000, or higher, and will contain from 75.0 percent to 90.0 percent of the metals present 30 in the fresh feed charge stock. The sulfur content will be approximately twice that of the fresh charge. The molecular weight of the resin concentrate will be lower, and in the range of about 1,000 to about 4,000; it will contain only a minor portion of the virgin metal con³⁵ taminants, and have a sulfur content about one and one-half times greater than the feed stock.

before describing my invention further, and especially with respect to the embodiment illustrated in the accompanying drawing, several definitions are believed ⁴⁰ necessary in order that a clear understanding be obtained. In the present specification and appended claims, a “pressure substantially the same as,” of a “temperature substantially the same as,” is intended to connote the pressure or temperature on a downstream i ⁴⁵ vessel, allowing only for the normal pressure drop due to fluid flow through the system, and the normal temperature loss due to the transfer of material from one zone to another. For example, where the pressure at the inlet to the first catalytic reaction zone is about 3,000 psig., and the effluent temperature is about 750°F,, the first separation zone will function at substantially the same pressure and temperature of about 2,925 psig. and 750°F. Similarly, the utilization of the phrase “at a least a portion of, when referring to ei⁵⁵ ther a principally vaporous phase, or a principally liquid phase, is intended to encompass both an aliquot portion as well as a select fraction. Thus, at least a portion of a hydrogen-rich principally vaporous phase is recycled to a catalytic reaction zone, following the re⁶⁰ moval of hydrogen sulfide therefrom, while at least a portion of the liquid phase (in this case an aliquot portion) may be recycled to a catalytic reaction zone to combine with the fresh feed charge stock thereto.

As previously set forth, the combination process of the present invention utilizes two hydrocracking reaction zones. In most instances, the catalytic composites disposed within the two reaction zones will be of differ50

3,775,293 ent physical and chemical characteristics; it is understood, however, that they may be identical. Regardless, the catalytic composites comprise metallic components selected, from the metals of Groups VI-B and VIII of the Periodic Table, and compounds thereof. Thus, in accordance with the Periodic Table of The Elements, E. H. Sargent and Co., 1964, suitable metallic components are those selected from the group consisting of chromium, molybdenum, tungsten, iron, ruthenium, osmium, cobalt, rhodium, iridium, nickel, palladium and platinum. Additionally, recent investigations have indicated that catalytic composites, for utilization with excessively high-sulfur content feed stocks, are improved through the incorporation of a zinc, tin and/or bismuth component. Throughout the present specification and the appended claims, the use of the term “component,” when referring to the catalytically active metal, is intended to connote the existence of the metal within the catalytic composite either in some combined form, or in the elemental state. Regardless, the stated concentration of the metallic component is computed on the basis of the elemental metal. While neither the precise composition, nor the method of manufacturing the catalytic composites, is considered essential to my invention, certain aspects are preferred. For example, since the charge stock to the present process is of a high-boiling nature, it is preferred that the metallic components of the catalyst possess the propensity for effecting hydrocracking while simultaneously promoting the conversion of sulfurous compounds into hydrogen sulfide and hydrocarbons. The concentration of the catalytically active metallic component, or components, is primarily dependent upon the particular metal as well as the physical and/or chemical characteristics of the feed stock. For example, the metallic components of Group VI-B are generally present in an amount within the range of about 4.0 percent to about 30.0 percent by weight, the iron-group metals in an amount within the range of about 0.2 percent to about 10.0 percent by weight, whereas the moble metals of Group VIII are preferably present in an amount within the range of about 0.1 percent by weight, all of which are calculated as if these components existed within the catalytic composite in the elemental state. When a zinc, tin and/or bismuth component is utilized, the same will be present in an amount of about 0.1 percent to about 5.0 percent by weight.

The porous carrier material, with which the catalytically active metallic components are combined, is a refractory inorganic oxide of the character thoroughly described in the literature. When of the amorphous type, alumina, or alumina in combination with about 10.0 percent to about 90.0 percent by weight of silica is preferred. It is often appropriate to utilize a carrier material comprising a crystalline aluminosilicate, or zeolitic molecular sieve. In most instances, such a carrier material will be utilized in processing the deasphalted oil in the second catalytic reaction zone. The zeolitic material includes mordenite, faujasite, Type A or Type U molecular sieves, etc. These may be employed in a substantially pure state; however, it is contemplated that the zeolitic material may be included within an amorphous matrix such as silica, alumina, and mixtures of alumina and silica.

It is further contemplated that a halogen component may be combined with the other components of the catalytic composite. Although the precise form of the chemistry of association of the halogen components with the carrier material and metallic components is not accurately known, it is customary in the art to refer to the halogen component as being combined with the carrier material or with the other ingredients of the catalyst. The halogen may be either fluorine, chlorine, iodine, bromine, or mixtures thereof, with fluorine and chlorine being particularly preferred. The quantity of halogen is such that the final catalytic composite contains about 0.1 percent to about 3.5 percent by weight, and preferably from about 0.5 percent to about 1.5 percent by weight, calculated on the basis of the elemental halogen.

The metallic components may be incorporated within the catalytic composite in any suitable manner including co-precipitation or cogellation with the carrier material, ion-exchange or impregnation of the carrier material. Following the incorporation of the metallic components, the catalyst is dried and subjected to a high temperature calcination or oxidation technique at a temperature of about 750°F. to about l,300°F. When a crystalline aluminosilicate is utilized as part of the carrier material, the upper limit for the calcination technique is preferably about l,000°F.

With respect to the catalyst utilized in the catalytic first reaction zone, a preferred composite is of the character described in U.S. Pat. No. 3,640,817 (Class 208-59). Briefly, this catalyst consists of a carrier material of alumina and silica containing from about 5.0 percent to about 30.0 percent by weight of boron phosphate, and has more than about 50.0 percent of its macropore volume consisting of pores having nominal diameters greater than about 1,000 Angstroms.

Prior to its utilization for the desulfurization/hydrocracking of hydrocarbons, the dried and calcined catalytic composite may be subjected to a substantially water-free reduction technique. Substantially pure and dry hydrogen (less than about 30.0 volumetric ppm. of water) is employed as the reducing agent. The calcined composite is contacted at a temperature of about 800°F. to about 1,200°F. and for a period of about 0.5 to about 10 hours. This reduction technique may be performed in situ prior to introducing the charge stock.

Additional improvements are generally obtained when the reduced composite is subjected to presulfiding for the purpose of incorporating therewith from about 0.05 percent to about 0.5 percent by weight of sulfur, on an elemental basis. The presulfiding treatment is effected in the presence of hydrogen and a suitable sulfur-containing compound such as hydrogen sulfide, a low molecular weight mercaptan, various organic sulfides, carbon disulfide, etc. One technique involves treating the reduced catalyst with a sulfiding gas, such as a mixture of hydrogen and hydrogen sulfide having about 10 moles of hydrogen per mole of hydrogen sulfide, and at conditions selected to effect the desired incorporation of sulfur. Presulfiding may also be effected in situ by way of charging a relatively low boiling hydrocarbon feed containing sulfurous compounds.

As hereinbefore set forth, the present invention utilizes two fixed-bed catalytic reaction zones and a noncatalytic thermal cracking zone. The resin concentrate is processed in a catalytic first reaction zone, and the product effluent therefrom is separated in a hot separator at substantially the same pressure. The principal function served by the hot separator is to separate the mixed-phase product effluent into a vapor phase rich in

3,775,293 hydrogen and a principally liquid phase which may contain from about 10.0 mol.% to about 40.0 mol.% of dissolved hydrogen. In a preferred embodiment, the total reaction product effluent from the catalytic first reaction zone is utilized as a heat-exchange medium in order to lower the temperature thereof to a level in the range of about 700°F. to about 800°F. The liquid phase from the hot separator may be recycled, at least in part, to combine with the fresh resin concentrate, thereby serving as a diluent for the heavier constituents thereof. The quantity of the liquid phase diverted in this manner is such that the combined feed ratio to the catalytic first reaction zone, being defined as total volumes of liquid charge per volume of fresh liquid charge, is within the range of about 1.1:1 to about 3.5:1. The remaining portion of the principally liquid phase from the hot separator is introduced into the thermal cracking reaction zone, or coil, at a reduced pressure in the range of about 200 psig. to about 500 psig. and at a temperature of from about 700°F. to about 950°F. As hereinafter indicated in the description of the accompanying drawing, the product effluent from the thermal cracking coil is introduced into a vacuum flash column maintained at about 20 to about 60 mm. Hg., absolute. The principal function of the vacuum flash zone is to concentrate the remaining heavy, metal-containing resins as a byproduct stream while recovering distillable hydrocarbons as a principally vaporous phase. The vaporous phase from the vacuum flash zone, in combination with the vaporous phase recovered from the hot separator and the deasphalted and deresined oil is processed in the catalytic third reaction zone to produce additional desulfurized lower-boiling hydrocarbon products.

With respect to the two catalytic reaction zones, the operating conditions of temperature, pressure, liquid hourly space velocity and hydrogen/hydrocarbon ratio will be within the same ranges. However, a preferred technique dictates operating the catalytic first reaction zone, processing the resin concentrate, at a lower severity than that imposed upon the catalytic third reaction zone. The variance in operating severity levels between the two catalytic reaction zones is readily obtained through the adjustment of the pressure, maximum catalyst bed temperature and liquid hourly space velocity, the higher severity operation will normally be effected at an increased pressure, an increased maximum catalyst bed temperature and at a decreased liquid hourly space velocity, or some combination thereof. The maximum catalyst bed temperature within the catalytic first reaction zone will be at least about 20°F. lower than that maintained within the catalytic third reaction zone, in most instances.

With respect to the operating conditions imposed upon the catalytic reaction zones, they are selected primarily to effect the conversion of sulfurous compounds into hydrogen sulfide and hydrocarbons, while simultaneously inducing hydrocracking reactions to produce lower-boiling hydrocarbon products. As hereinbefore set forth, the operating conditions imposed upon the catalytic third reaction zone will result in a higher operating severity. Suitable ranges for the various variables will generally be the same for both reaction systems. Thus, the pressure will range from about 500 to about 3,500 psig., and preferably from about 500 to about 2,500 psig. The maximum catalyst bed temperature will be within the range of about 600°F. to about 900°F. In view of the fact that the reactions being effected in the catalytic reaction zones are principally exothermic, an increasing temperature gradient will be experienced as the reactants traverse the catalyst bed. Judicious operating techniques dictate that the increasing tempera5 ture gradient be limited to a maximum of about 100°F., and, in order to control the increasing temperature gradient, it is within the scope of the present invention to employ quench streams, either normally liquid, or normally gaseous, introduced at one or more intermediate 10 loci of the catalyst bed. The hydrogen concentration is expressed as scf./Bbl. of charge, and will usually be within the range of about 1,000 to about 30,000. Liquid hourly space velocities, defined as volumes of normally liquid hydrocarbons charged per hour, per vol15 ume of catalyst disposed within the reaction zone, will be from about 0.25 to about 2.50. In addition to the temperature variable, the liquid hourly spaced velocity is conveniently utilized to adjust the operating severity between the two catalytic reaction zones. Thus, the liq20 uid hourly space velocity through the second reaction zone will generally be less than that through the first reaction zone.

That portion of the effluent from the catalytic third reaction zone boiling at a temperature above that de²⁵ sired with respect to the recovered product streams, may be recycled in order to produce additional lowerboiling hydrocarbon products. When this technique is utilized, the combined feed ratio, defined as total volumes of normally liquid charge to the catalytic third re³⁰ action zone, per volume of fresh charge thereto, will be within the range of about 1.1:1 to about 6.0:1.

Other conditions and preferred operating techniques will be given in conjunction with the following description of the present process. In further describing this ³³ process, reference will be made to the accompanying figure which illustrates one specific embodiment. In the drawing, the embodiment is presented by means of a simplified flow diagram in which many details such as pumps, instrumentation and controls, heat-exchange ⁴⁰ and heat-recovery circuits, valving, start-up lines and similar hardware have been omitted as being nonessential to an understanding of the techniques involved. The use of such miscellaneous appurtenances, to modify the process, are well within the purview of ⁴³ one skilled in die art.

The major, vessels integrated within the combination process of the present invention, as illustrated in the drawing, are as follows: the first solvent extraction zone is deasphalting zone 2, while the second solvent extrac⁵⁰ tion zone is deresining zone 5; the catalytic first reaction zone is reactor 9, the non-catalytic, second reaction reaction zone is thermal coil 14 and the catalytic third reaction zone is reactor 20; and, the first separa55 tion zone is hot separator 11, the second separation zone is vacuum column 16 and the third separation zone is cold separator 22. A fourth separation zone is illustrated as fractionator 24 and functions to recover the various desired product fractions. For example, a propane-minus stream may be recovered through line 25, a butane concentrate through line 26, a combined pentane/hexane concentrate in line 27, a naphtha boiling range, heptane-400°F. product in line 28 and 400°F.-650°F. middle-distillate through line 29.

DESCRIPTION OF DRAWING

The accompanying drawing will be described in conjunction with a commercially scaled unit designed to

3,775,293 process 80,000 Bbl./day of vacuum column bottoms. Charge stock analyses indicate a gravity of about 10.1 “API, 3.08 percent by weight of sulfur, 186 weight ppm. of metals, a Conradson Carbon content of 15.8 percent and a heptane-insoluble portion in the amount of 5.2 percent by weight. The desired product slate includes a light naphtha (heptane-275°F.), a heavy naphtha (275°F-380°F.) and a diesel fuel (380°F-650°F.). All the desired fractions are intended to be substantially free from nitrogenous and sulfurous compounds. In the 1 description, the yields, unless otherwise specifically stated, are given in weight percent, and are based upon the vacuum bottoms charge.

The charge stock is introduced, via line 1, into a deasphalting zone 2, wherein it countercurrently contacts a pentane/butane solvent introduced via line 31. The solvent extraction is effected in substantially liquid phase at a pressure of about 400 psig. and a temperature of 245°F., with a solvent/oil volumetric ratio of 3.0:1.0. A solvent-lean asphaltic pitch, in the amount : of about 15.0 percent by weight, having a gravity of —8.1 “API, is withdrawn through line 3 while a solventrich, resin-containing first liquid phase is recovered via line 4.

The first liquid phase countercurrently contacts additional pentane-butane solvent, at a solvent/oil volumetric ratio of 5.0:1.0, introduced into deresining zone 5 by way of line 32. The temperature is 300°F. and the pressure about 400 psig., which produces a resin concentrate in the amount of about 29.0 percent by weight, having a gravity of 1.62 °API, precipitated and withdrawn by way of line 6. The resin concentrate has a metals concentration of 136 ppm. by weight. A solvent-rich second liquid phase is removed through line 7, in an t_____ quently reacted with hydrogen in reactor 20. The deresined oil in line 7 has a gravity of 16.9 “API, and contains 2.0 percent by weight of sulfur and only 3.0 ppm. of metal contaminants. The 29.0 percent by weight of resin concentrate continues through line 6, is admixed with a hydrogen-rich, principally vaporous phase from line 8, and introduced thereby into a catalytic first reaction zone 9.

The hydrogen concentration in reactor 9 is about 5,000 scf./Bbl. and the pressure is maintained at 2,500 psig. A temperature gradient of 100°F. is controlled through the use of a hydrogen quench stream, while the reactants traverse the catalyst bed at a liquid hourly space velocity of 1.0, to result in a maximum catalyst bed temperature of 875°F. Reactor 9 contains a catalyst of 1.89 percent by weight of nickel, 16.0 percent molybdenum, 8.78 percent of boron phosphate, 6.97 percent silica and 66.96 percent by weight of alumina. The first reaction zone product effluent is withdrawn through line 10, and introduced into hot separator 11 at substantially the same pressure and a temperature of about 750°F. A first principally vaporous phase is withdrawn through line 12, and is admixed with the deasphalted and deresined oil in line 7. A heavy, third principally liquid phase is removed via line 13, to be charged to thermal coil 14. A material balance around, hot separator 11 is given in the following Table I:

TABLE I: Hot Separator Balance

Line No.

Gases

Light Naphtha Heavy Naphtha

Diesel Fuel Heavy Oil Resins

TOTALS:

5.5

6.9

14.5

29.5

4.0

0.9 Trace

6.0

14.5

22.0

The figures presented in Table I are inclusive of hydrogen consumption in an amount of 0.5 percent by weight.

The remaining 14.5 percent by weight of resins are introduced into a non-catalytic, second reaction zone 14 (thermal coil) at a pressure of about 200 psig. Thermal reactions therein are carried out at a temperature of 950°F., to produce a thermally cracked product effluent in line 15 which is separated in vacuum column 16 at a pressure of 60 mm. Hg., absolute and a temperature of 800°F. Distillable hydrocarbons, as the second principally vaporous phase, are removed through line 17, to be combined with the derisined oil in line 7, and a metal-containing, heavy resin by-product is withdrawn from the process by way of line 18, in an amount of 4.5 percent by weight. It should be noted that more than 86.0 percent of the resin concentrate, precipitated in deresining zone 5, has been converted into more valuable, distillable hydrocarbon products. The material balance around vacuum column 16 is presented in the following Table II:

TABLE II: Vacuum Column Balance

Line No.D

Gasesθ·5

Light Naphtha°·

Heavy Naphtha1 θ

Diesel Fuel5.5

Heavy Oil'θ·'

Heavy Resins

The deasphalted oil in line 7, in admixture with the the second vaporous phase from vacuum column 16 (line 17) and a heavy oil recycle stream from line 30, the source of which is hereafter set forth, is introduced into catalytic third reaction zone 20 in admixture with ⁴⁰ a hydrogen-rich vaporous phase in line 19.

Reactor 20 has disposed therein a catalyst of 1.9 percent by weight of nickel, 14.1 percent by weight of molybdenum, 27.3 percent of silica and 56.7 percent alumina. The hydrogen concentration is about 6,000 ⁴⁵ scf./Bbl., and the pressure is maintained at about 2,400 psig. The normally liquid portion of the feed stock traverses the catalyst bed at a liquid hourly space velocity of 0.6, and the maximum catalyst bed temperature is controlled at about 875°F.

The resulting reaction product effluent passes through line 21, at substantially the same pressure, and, after being used as a heat-exchange medium and further cooling, into cold seaparator 22 at a temperature of about 100°F. A hydrogen-rich, third vaporous phase is withdrawn via line 8, and in part recycled thereby to reactor 9; a portion is diverted from line 8 by line 19, as hydrogen recycle to reactor 20. The normally liquid portion of the product effluent is removed by way of line 23 and introduced thereby into fractionator 24 for separation into the various product streams. A butaneminus stream is recovered via line 25, a pentane/hexane concentrate through line 26, the light naphtha fraction via line 27, the heavy naphtha fraction through line 28 and the diesel oil by way of line 29. Heavy oil, boiling beyond the diesel oil end boiling point of 650°F., is recycled in an amount of 20.0 percent by weight, through lines 30,17 and 7, for further conversion in re10

0.5 0.4 1.0 5.5 10.1

4.5 bCUVllu U4U1VI ------_ο------- - ...

amount of 56.0 percent by weight, and is subse- ³⁵ first vaporous phase from hot separator 11 (line 12), * ____ . ., 1 ___________ —1_  ,>nz%,anm ^nlumn lh

<td> 10</td><td> 12</td><td> 13</td>

<td> 1.5</td><td> 1.5</td><td> Trace</td>

<td> 0.3</td><td> 0.3</td><td> Trace</td>

<td> 0.8</td><td> 0.8</td><td> Trace</td>

3,775,293 actor 20. The balance around reactor 20, including the heavy oil recycle, is presented in the following Table III:

<td> TABLE III:</td><td> Reactor 20 Balance</td><td> 5</td>

<td> Line No.</td><td colspan="2"> 7 (Charge) 23 (Product)*</td>

<td> Gases</td><td> 2.0</td><td> 7.5</td>

<td> Light Naphtha</td><td> 0.7</td><td> 3.0</td>

<td> Heavy Naphtha</td><td> 1.8</td><td> 10.5</td>

<td> Diesel Fuel</td><td> 9.5</td><td> 61.5</td>

<td> Heavy Oil</td><td> 31.0</td><td>^2&lt;2 10</td>

<td colspan="2"> PDR** 56.U * Includes total hydrogen consumption of 2.0% by weight * * Propane deresined oil from extraction zone 5</td><td></td>

Overall product yield and component distributions are presented in the following Table IV, and are based jg on a fresh feed charge stock rate of 80,000 Bbl./day.:

TABLE IV: Product Yield and Distribution

<td> Component</td><td> Wt.%</td><td> Vol.%</td><td> Bbl./day</td>

<td> Ammonia</td><td> 0.2</td><td> —</td><td> —</td>

<td> Hydrogen Sulfide</td><td> 2.2</td><td> —</td><td> —</td>

<td> Methane</td><td> 0.2</td><td> ——</td><td> ——</td>

<td> Ethane</td><td> 0.4</td><td> —</td><td> —</td>

<td> Propane</td><td> 0.7</td><td> —</td><td> —</td>

<td> Butanes</td><td> 1.3</td><td> 2.3</td><td> 1,840</td>

<td> Pentanes</td><td> 1.2</td><td> 1.9</td><td> 1,520</td>

<td> Hexanes</td><td> 1.3</td><td> 1.9</td><td> 1,520</td>

<td> Heptane—275°F.</td><td> 3.0</td><td> 4.2</td><td> 3,360</td>

<td> 275°F.-380°F.</td><td> 10.5</td><td> 13.8</td><td> 11,040</td>

<td> 380°F.-650°F.</td><td> 61.5</td><td> 74.1</td><td> 59,280</td>

<td> Asphaltic Pitch</td><td> 15.0</td><td> 13.3</td><td> 10,640</td>

<td> Heavy Resins</td><td> 4.5</td><td> 4.1</td><td> 3,280</td>

<td> TOTALS:</td><td> 102.0</td><td> 115.6</td><td> 92,480</td>

All the normally liquid streams, including the pentane/hexane concentrate, which may be supplied to an isomerization zone to produce high octane isomers, indicate substantially no sulfur- and/or nitrogencontaining compounds. The propane and butanes may _ggbe recovered as a concentrate and employed as the feed stream to a steam reforming unit to produce a methane-rich synthetic natural gas, or as the feed to a dehydrogenation unit to produce olefins for subsequent alkylation to a high octane alkylate motor fuel. ₄₀The foregoing indicates the method of effecting the present combination process and the benefits afforded through the utilization thereof.