US3696022A

US3696022A - Swing-bed guard chamber in hydrogenerating and hydrorefining coke-forming hydrocarbon charge stock

Info

Publication number: US3696022A
Application number: US58443A
Authority: US
Inventors: Leroi E Hutchings
Original assignee: Universal Oil Products Co
Current assignee: Honeywell UOP LLC; Universal Oil Products Co
Priority date: 1970-07-27
Filing date: 1970-07-27
Publication date: 1972-10-03
Anticipated expiration: 1989-10-03

Abstract

A PROCESS FOR HYDROREFINING SULFUROUS HYDROCARBON DISTILLATES CONTAINING MONO-OLEFINIC, DI-OLEFINIC AND AREMATIC HYDROCARBONS. PARTIULARLY DIRECTED TOWARD THE STABILIZATION OF A PYROLYSIS NAPHTHA, THE PROCESS MAKES USE OF A MULTIPLE-STAGE SYSTEM. EASE OF OPERATION, INCREASED CATALYST STABILITY THROUGHOUT THE PROCESS AND VARIOUS ECONOMIC ADVANTAGES ARE OBTAINED THROUGH THE UTILIZATION OF A SWING-BED REACTION ZONE SYSTEM WHICH CONSTITUTES THE INITIAL PORTION OF THE PROCESS. IN COMBINATION WITH THE

FIRST PRINCIPAL REACTION ZONE, FOLLOWING THE SWING-BED SECTION, EACH OF THE SMALLER GUARD CHAMBER REACTORS IN THE LATTER CONTAIN FROM 5.0% TO ABOUT 25.0% OF THE TOTAL CATALYST.

Description

Oct. 3, 1972 SWING-BED GUARD CHAMB'ER L E. HUTCHINGS IN HYDROGENATING `AND HYDROREFINING A COKE-FORMING HYDROCARBON CHARGE STOCK Filed July 27, 1970 A TTORNEYS Unted States Patent O 3,696,022 SWING-BED GUARD CHAMBER IN HYDRO- GENERATING AND HYDROREFINING COKE- FORMING HYDROCARBON CHARGE STOCK Leroi E. Hutchings, Mount Prospect, Ill., assignor to Universal Oil Products Company, Des Plaines, Ill. Filed July 27, 1970, Ser. No. 58,443 Int. Cl. Cg 23/00 U.S. Cl. 208--57 9 Claims ABSTRACT OF THE DISCLOSURE APPLICABILITY OF INVENTION In the present specification and appended claims, the terms hydrocarbons, hydrocarbon fractions, hydrocarbon distillates, and hydrocarbon mixtures are utilized interchangeably to connote synonmously mixtures of hydrocarbons resulting from diverse conversion processes. Such processes include the catalytic and/ or thermal cracking or petroleum, the latter often referred to as pyrolysis, the destructive distillation of wood or coal, shale oil retorting, etc. The resulting hydrocarbon distillate fractions contain impurities which must necessarily be removed before the fractions are suitable for their intended use, or, when removed, enhance the value of the fraction with respect to further processing. Contaminating influences, in the form of these impurities, include sulfurous compounds, nitrogenous compounds and often oxygenated compounds.

In addition to the aforementioned contaminating inuences, these hydrocarbon fractions contain appreciable quantities of unsaturated compounds, including monoolefinic, di-olefinic (including `conjugated di-olefns) and aromatic hydrocarbons. It is the mono-olefinic and diolenic hydrocarbons, particularly butadiene and pentadienes, which induce the polymer and coke-forming tendencies of the hydrocarbon distillate, as do the vinyl aromatics, and which, when the distillate is subjected to hydrotreating for the purpose of removing the contaminating influences, effect the formation of polymers as Well as other carbonaceous material. The formation of olefin polymers, and deposition of coke, appears to be a result of the necessity for effecting the hydrotreating process at f elevated temperatures above about 500 F. in order to convert sulfurous compounds into hydrogen sulfide and hydrocarbons. Not only are the hydrotreating catalytic composites affected, but various heaters and other appurtenances of the conversion zone experience heavy coking which appears as a formation of solid, highly carbonaceous material principally resulting from the polymerization of conjugated di-olefns, and styrenes and olens contained in the fresh feed charge to the unit.

As hereinbefore set forth, coke-forming hydrocarbon distillates are usually those fractions resulting from prior conversion processes such as catalytic or thermal cracking, or destructive distillation. In the interest of brevity, the following discussion will be directed toward hydro- 3,696,022 Patented Oct. 3, 1972 ice carbon distillates resulting from naphtha pyrolysis units generally designed for the production of normally gaseous olefinic material such as ethylene, propylene, butadiene, etc. Pyrolysis naphtha co-product is available in relatively large quantities, but requires a hydrorefining treatment for the purpose of enhancing its possibilities for further usefulness. In many instances, the pyrolysis naphtha c0- product Will not contain excessive quantities of sulfurous and/or nitrogenous compounds, but will consist of detrimental amounts of mono-olefins and di-olens, including conjugated di-olens, to the extent that immediate use of the distillate is prohibited. For example, in a thermal cracking process designed for ethylene production, a full boiling range hydrocarbon distillate is produced which may contain less than 1,000 p.p.m. by weight each of sulfur and nitrogen, but sufiicient quantities of olefinic hydrocarbons to indicate a bromine number of the order of at least about 25.0, and often more, and di-olens in an amount to indicate a diene value of the order of about 20.0, or more.

Pyrolysis reactions are effected in the absence of a catalytic composite, at elevated temperatures and in the presence of a diluent such as superheated stream. Depending upon the physical and/or chemical characteristics of the charge stock, as well as the specific pyrolysis conditions, the product eiuent comprises varying quantities of light oleinic hydrocarbons including ethylene, propylene, `butylene, butadiene, etc., and a pyrolysis naphtha fraction containing pentanes, hexanes and heavier hydrocarbons boiling up to a temperature in the range of about 300 F. to about 400 F., and including aromatic hydrocarbons, mono-olefnic and di-olefinic hydrocarbons, styrenes and sulfurous compounds. Visualized as being exemplary of those pyrolysis naphthas intended for processing, in accordance With the present invention, are: a naphtha having a gravity of about 35.1 API, and containing 1,100 p.p.m. by weight of sulfur, having an end boiling point of 350 F., a bromine number of 43.0 and a diene value of 40.0; a 400 F. (end boiling point) naphtha having a gravity of 43.5 API, and containing 500 p.p.m. by Weight of sulfur, having a bromine number of 74.0 and a diene value of about 80.0; and, a butane/ pentane concentrate having a gravity of 76.7 API containing 500 p.p.m. by Weight of sulfur, and having a lbromine number of about 200.0 and a diene value of about 230.0.

Since such pyrolysis naphtha fractions are severely contaminated, it is a common practice to hydrotreat or saturate the olens and/or di-olens, while destructively converting the sulfurous and nitrogenous compounds into hydrogen sulfide, ammonia and hydrocarbons. As hereinbefore stated, di-olefnic hydrocarbons present particular diiculty in the operation of the hydrotreating facilities by way of catalyst deactivation and extensive equipment fouling. Attempts are often made to improve the on-stream efliciency of the hydrotreating process by either promoting the polymerization of the diolenic hydrocarbons prior to the hydrotreating step, or by utilizing operating techniques which tend to minimize or inhibit polymer formation. However, none of these approaches are suiiiciently successful in overcoming the difficulty, and to the extent that the process enjoys an acceptable degree of catalyst stability.

Of further significance is the lack of selectivity in prior art processes. For example, the hydrogenation of the conjugated di-oleiinic hydrocarbons and styrenes may not cease with the conversion thereof to saturates and alkyl benzenes, but will continue to complete saturation of the mono-olens in the charge stock and the aromatic nuclei in the styrene compounds. Such non-selectivity obviously results in a decrease of desirable products in the thus-treated product efuent. Candor compels recognition of the fact that usch prior art processes have recently been improved. Since pyrolysis naphtha fractions tend to form gums and other olefin polymer material upon standing, one such improvement involves removal of this material, butadienes and pentadienes from the fresh feed charge stock in a fractionation column; this technique is commonly referred to in the art as rerunning and depentanizing the fresh feed. Additionally, another principal improvement involves the use of a multiple-stage system in which the first stage functions at a temperature below that at which desulfurization is effected, or below about 500 F., while the second stage, in order to effect desulfurization, functions at temperatures exceeding 500 F. Nothwithstanding these improvements, which admittedly result in increased catalyst stability and improved product quality additional improvements are afforded through the utilization of my invention. The present process improves the quality of the matreial being charged to both of the principal reaction zones, thereby further increasing catalyst stability, or catalyst life. Ultimate product quality is also further improved with respect to the quantity of mono-olefinic hydrocarbons. Heretofore, particularly in European and other foreign locales, the presence of mon-olefinic hydrocarbons, as a result of their excellent blending values, were tolerated in motor fuel gasolines. Since the burning of olefin-containing motor fuel has been found to be a contributing factor to pollution of the atmosphere, stringent requirements are imposed whereby the olefin concentration of motor fuel gasoline must be essentially nil-77 In accordance with the present process, through the use of particular catalytic composites, operating conditions and especially operating techniques, a coke-forming hydrocarbon distillate is selectively hydrogenated and desulfurized with minimum polymer formation and minimum degradation of desired constituents. Briefly, the present invention provides a series of guard chambers, in the form of catalytic reaction zones, employed in a swing-bed arrangement whereby one or more chambers are on-stream, while one or more chambers are simultaneously being regenerated. This feature of the present invention eliminates the necessity for re-running of the fresh feed charge stock and improves catalyst stability, and is based upon recognition of the fact that polymer formation, notwithstanding operating temperatures below about 500 F., takes place in the initial 5.0% to about 25.0% of the first stage catalyst bed. This feature leads to a continuous hydro-treating process through the elimination of frequent shut-down periods for regeneration of the catalytic composites.

In the absence of the guard chambers, but including the re-running and depentanizing of the fresh feed charge stock, a commercial unit must be shut down for hot hydrogen stripping, to remove olefin polymers, every ten to sixty days, the latter at the start of the run, the former near the end of the run when the catalytic composite must be burned Through the incorporation f the guard chambers, and elimination of re-running and depentanizing, shut downs for hydrogen stripping of the catalyst in the principal reaction zones are nonexistent a requirement for burning the catalyst in the guard chambers does not lead to a shut down, and the necessity for burning the principal catalytic beds is greatly reduced.

OBJ-ECTS AND EMBODIMENTS A principal object of the present invention is to effect the saturation of olefinic hydrocarbons and the desulfurization of sulfurous compounds present in coke-forming hydrocarbon distillates. A corollary objective is to provide catalyst stability in a process for stabilizing cokeforming pyrolysis gasoline.

A specific object of my invention is to provide a continuous hydrotreating process into which catalyst regeneration facilities are integrated with the result that catalyst stability is significantly improved and little degradation of aromatic hydrocarbons exists.

Therefore, in one embodiment, my invention provides a continuous process for hydrorefining a sulfurous, hydrocarbon charge stock containing mono-olefinic and diolefinic hydrocarbons, which process comprises the steps of: (a) reacting said charge stock and a regenerating hydrogen-rich gaseous phase in a first catalytic reaction zone at a maximum catalyst bed temperature in the range of about 200 F. to about 500 F.; (b) further reacting the resulting first reaction zone efiiuent, at substantially the same temperature in the range of about 200 F. to about 500 F., in a second catalytic reaction zone; (c) increasing the temperature of the resulting second reaction zone efiiuent to a level above about 500 F., and reacting the thus-heated second zone effluent in a third catalytic reaction at a maximum catalyst bed temperature below about l,000 F.; (d) separating the resulting third reaction zone efiiuent, at a temperature in the range of about 200 F. to about 500 F., in a first separation zone, to provide a first vaporous phase and a first liquid phase; (e) further separating said first vaporous phase, at a substantially lower temperature in a second separation zone, to provide a second liquid phase and a hydrogen-rich second vaporous phase; (f) heating said hydrogen-rich second vaporous phase to a temperature in the range of about 500 F. to about 700 F.; (g) passing the thus-heated second vaporous phase through a fourth catalytic reaction zone containing an olefin polymer-deactivated catalytic composite; (h) separating resulting fourth reaction zone efiiuent, in a third separation zone, to provide a polymer concentrate and a hydrogen-rich third vaporous phase; and, (i) recyclying said third vaporous phase to combine with said charge stock.

Other embodiments of my invention involve preferred processing techniques and operating conditions, as well as particular catalytic composites for utilization in the various catalytic reaction zones. With respect to the latter, it is preferred to utilize, in the initial major catalytic reaction zone, a catalytic composite capable of selectively hydrogenating conjugated di-olefinic hydrocarbons in a hydrocarbon charge stock, also containing sulfurous compounds, aromatic hydrocarbons and mono-olefinic hydrocarbons, to the corresponding saturates without significant saturation of the mono-olefins present. Thus, a preferred catalyst for use in the first reaction zone is a composite of a Group VIII noble metal component, an aluminacontaining non-acidic carrier material and an alkali metal component. With respect to the catalytic composite disposed within the second major reaction zone, intended for desulfurization and mono-olefin hydrogenation, accompanied by preservation of the aromatic hydrocarbons, the character of the catalytic composite, as hereinafter indicated, is dependent upon the maximum catalyst bed temperature within the reaction zone.

In another embodiment, the first and fourth catalytic reaction zones function in swing-bed fashion and are interchanged when the deactivated catalytic composite in the fourth reaction zone is substantially free from olefin polymers, with the result that the charge stock and regenerating hydrogen gaseous phase are then introduced into the fourth reaction zone, and the heated hydrogen-rich second vaporous phase is introduced into the first catalytic reaction zone to remove olefin polymers from the catalyst disposed therein.

SUMMARY OF THE INVENTION As hereinbefore set forth, the principal purpose of the present invention is to provide a selective and highly stable continuous process for hydrogenating coke-forming hydrocarbon distillates. As utilized herein, the term hydrogenating is intended to be synonymous with hydrotreating and hydrorefining. In essence, this purpose is accomplished through the use of a two-stage fixed-bed catalytic reaction system, having integrated therein a swingbed regeneration technique.

Previously, there existed two separate, desirable routes for the treatment of coke-forming distillates. One such route was directed toward recovering a hydrocarbon product suitable for use in certain gasoline blending. With this as the desired object, the hydrorefining process could be effected in a. single-stage, or reaction zone, in general utilizing the catalytic composite hereinafter specifically described as the third reaction zone catalyst. The attainable selectivity in this instance resides primarily in the hydrogenation of highly reactive double bonds, while leaving native mono-olefins in the charge stock unreacted; with respect to styrene, the hydrogenation is inhibited to produce alkyl benzenes without ring saturation. Until the stringent restrictions were placed upon the olefin content of motor fuels, the single stage process was suitable for use in some locales. Where, however, the desired end result was aromatic hydrocarbon retention, intended for subsequent recovery via extraction, the two-stage route was required. Currently, regardless of the ultimate use of the aromatic-containing product, the two-stage route becomes necessary. It has previously been noted that the sulfurous compounds and the mono-olefins remain substantially unchanged in the single, or the first-stage reaction zone. This restrictive hydrogenation selectivity is accomplished with reduced polymer formation to gums Other advantages of restricting the hydrogenation to the conjugated di-olefins and the olefinic bond in the styrenes include: lower hydrogen consumption, a lower heat of reaction and, obviously, increased aromatic retention including the alkyl benzenes resulting from styrene hydrogenation. The mono-olefins must be saturated in the second stage, not only to `facilitate aromatic extraction by way of currently known and utilized methods, but also to comply with the impositions placed upon the olefin content of motor fuels. Thus, the desired, necessary second-stage hydrogenation technique involves saturation of the monoolefins, as well as sulfur removal, the latter required for an acceptable aromatic product.

As hereinafter set forth, in the description of the accompanying drawings, the present continuous process requires more than two catalytic reaction zones, and at least four catalytic reaction zones. Two initial reaction zones serve as swing-bed guard chambers, each of which, when considered in conjunction with the first major, or principal reaction zone, contains about 5.0% to about 25.0% by weight of the total catalytic composite disposed in both zones. In the present specification and appended claims, the first and fourth catalytic reaction zones are the guard chambers, and generally contain the same catalytic composite as is preferably disposed within the second catalytic reaction zone. By way of clarification, the second and third reaction zones of the present invention are those to which the prior art refers to as the first and second stages, respectively. In most applications of the present process, the catalytic composite will be the same in the first, second and fourth catalytic reaction zones.

The principal function of the catalytic composite utilized in the second reaction zone, and preferably in the first and fourth reaction zone guard chambers, involves selective hydrogenation of conjugated di-olefinic hydrocarbons. One particular catalytic composite possesses unusual stability notwithstanding the presence of sulfurous compounds in the fresh hydrocarbon charge stock. The catalytic composite is hereinafter described in greater detail. Briefly, however, the catalyst is a composite of an alumina-containing, non-acidic carrier material, a Group VIII noble metal component and an alkali metal component, the latter being preferably lithium and/ or potassium. In order to avoid hydrocracking activity, it is particularly preferred that this catalytic composite be substantially free from any acid-acting propensities. With respect to the third catalytic reaction zone, the catalytic composite functions for the principal purpose of effecting Ithe destructive Cit conversion of sulfurous and nitrogeneous compounds, as well as mono-olefin saturation, and is a composite of an alumina-containing carrier material, a Group VIII metal component; it may or may not also contain a halogen component. As hereinafter set forth, the character of the Group VIII metal utilized in the third catalytic reaction zone, is considered in conjunction with the maximum catalyst bed temperature employed therein.

Through utilization of a particular sequence of processing steps, the essence of which involves an integrated regeneration technique when hot hydrogen, employed as the regenerating gas, is introduced into the processin-g section, catalyst stability is improved to a degree which affords a continuous process and makes it unnecessary to rerun or depentanize the feed. Regeneration of the catalyst disposed in the major processing reactors can thus be postponed until such time as the unit is scheduled for its normal maintenance shut down. This sequence of processing steps, hereinafter set forth in greater detail, regulates the overall hydrorefining process in such a manner that the charge to the major processing section is not at conditions which induce the olefin polymerization reactions. The present process provides an improvement in the quality of charge stock to both the second and third catalytic reaction zones.

PROCESS CONDITIONS AND OPERATIONS The process of the present invention is effected in a sequence of reaction zones, each of which is maintained at operating conditions consistent with the chemical characteristics of the reactants passing therethrough. In order to afford a clear understanding of the integrated process, brief reference to the accompanying drawing is believed warranted. In the drawing,

reactors

12 and 18 are those which have previously been referred to as guard chambers. For the purpose of the following discussion, reactor 12 will be considered as undergoing regeneration with a hot hydrogen-rich gaseous phase, whereas reactor 18 is operating within the process. Throughout the specification, reactor 18 is referred to as the first reaction zone and reactor 12, the fourth reaction zone. The principal catalytic reaction zones are reactors I7 and 27, and are herein referred to as the second and third reaction zones, respectively. With respect to the distribution of the catalytic composite in the first, second and fourth reaction zones, each of the first and fourth zones will contain from about 5.0% to about 25.0% of the total catalyst disposed in each zone in conjunction with the quantity of catalyst disposed in the second catalytic reaction zone. That is to say, the distribution of catalyst with respect only to the first and second reaction zones is such that the quantity of catalyst in the first reaction zone is Within the aforesaid range.

Processing some charge stocks, having an extreme degree of unsaturation and high concentrations of contaminating influences, may result in too great a rise in temperature due to the exothermicity of the reactions. In such instances, a desirable technique constitutes providing for multi-point introduction of either the liquid feed, or internally recycled diluents, or recycled hydrogen-rich gaseous phase at various intermediate loci of the reaction zones. This tends to prevent excessive saturation from occurring in one particular portion of the catalyst, and also cools the charge stream as it passes through the reaction zone.

A hydrogen-rich recycle gaseous phase, the source of which is hereinafter set forth, is heated to a temperature in the range of about 500 F. to about 700 This hot hydrogen-rich stream is utilized to strip the catalytic composite which has become deactivated as a result of the formation of olefin polymers. The regenerating effluent is introduced into a suitable separation zone, such as knock-out pot, from which the stripped polymers and gums are withdrawn. Following its use as a heat-exchange medium, to reduce its temperature, the

hot regenerating hydrogen stream is combined with the hydrocarbon distillate charge stock, and liquid diluent where used, the mixture being introduced into the rst operating catalytic reaction guard chamber. The hydrocarbon distillate charge stock, for example a light naphtha by-product from a commercial thermal cracking unit designed and operated for the production of ethylene, having a gravity of about 39.9 API, a bromine number of about 42.0, a diene value of about 32.0 and containing about 100 p.p.m. of sulfur, 100 p.p.m. of nitrogen, and about 71.5 vol. percent aromatic hydrocarbons, is admixed with an internally recycled liquid diluent, the source of which is hereinafter set forth, to provide a combined liquid feed ratio in the range of about 1.1 to about 6.0. The hydrogen concentration is within the range of from about 500 to about 10,000 scf/bbl., and preferably in a narrower range of about 1,000 to about 6,000 sci/bbl. The maximum catalyst bed temperature is maintained in the range of about 200 F. to about 500 F., and preferably at a level of about 250 F. This guard chamber functions under a pressure of from about 100 p.s.i.g. to about 1,000 p.s.i.g., and preferably at a level in the range of from about 500 p.s.i.g. to about 900 p.s.i.g.

Without substantial change in either temperature, or pressure, the eliiluent from the guard chamber is introduced into a second catalytic reaction zone, being the rst of the two major reaction zones. This second reaction zone preferably contains a catalyst of the same character as that in the guard chamber, being a composite of alumina-containing non-acidic carrier material, a Group VIII noble metal component and an alkali metal component. As hereinbefore set forth, considering only the operating guard chamber and this second catalytic reaction zone, the former contains from about 5.0% to about 25.0% of the catalyst disposed in both zones.

The product efuent from the second catalytic reaction zone is introduced into a suitable direct-tired heater wherein the temperature is increased to a level above about 500 F. and preferably in the range of about 550 F. to about l000 F. When the process is functioning eiciently, the diene value of the normally liquid hydrocarbonaceous material entering the third catalytic reaction zone is in the range of about 0.1 to about 0.5, and very often nil. The conversion of nitrogenous and sulfurous compounds, and the saturation of mono-oletins, contained within the second reaction zone eluent, is effected in the third catalytic reaction zone. The second catalytic reaction zone is maintained under an imposed pressure of from about 100 to about 1,000 psig., and preferably at a level of from about 500 to about 900 p.s.i.g. Process operations are facilitated when the focal point for pressure control is the cold separator following the third catalytic reaction zone and, therefore', it will be maintained at a pressure slightly less than that imposed upon the rst and second catalytic reaction zones. The third catalytic reaction zone and, therefore, it will be maincatalyst sufficient to provide an LHSV of from about 0.5 to about 10.0, based upon fresh feed. The hydrogen concentration will be in the range of from about 500 to about 10,000 s.c.f./bbl., based upon fresh feed, and preferably from about 1,000 to about 8,000 s.c.f./bbl.

Following its use as a heat-exchange medium, to decrease its temperature to a level in the range of about 200 F. to about 500 F., and preferably from about 250 F. to about 400 F., the product eluent from the third catalytic reaction zone is introduced into a hot separator. A hot liquid phase is withdrawn from the hot separator and recycled to combine with the fresh feed charge stock to the tirst catalytc reaction zone as hereinbefore described. Where desired or necessary, a portion of the hot separator liquid may be employed to quench the reaction in the third catalytic reaction zone by being introduced thereto at some intermediate locus. The remainder of the hot separator liquid is then introduced into a suitable separation system which removes hydrogen sulfide and recovers the aromatic-rich naphtha boiling range fraction. The hydrogen-rich vaporous phase from the hot separator is cooled and condensed to a temperature of from 60 F. to about 140 F., and passed into a suitable cold high pressure separator. One suitable operating technique involves injecting water into the hot vaporous phase from the hot separator and equipping the cold separator with a water dip-leg from which sour Water containing ammonia and hydrogen sulfide is removed. A hydrogen-rich recycle gaseous phase is withdrawn from the cold separator, and, by way of compressive means, is recycled through a heater and a bed of olefin polymer deactivated catalyst as hereinabove set forth. Make-up hydrogen, to supplant that consumed in the overall process may be introduced from any suitable external source into any suitable location within the processing system; preferably, the make-up hydrogen is introduced by way of the efuent line from the second catalytic zone.

The gaseous phase from the cold separator may be treated in any manner for the removal of hydrogen sulfide and/ or light paraflinic hydrocarbons, in order to increase the concentration of hydrogen therein. The principally liquid phase from the hot separator may be admixed with the liquid phase from the cold separator the mixture being further separated in a hydrogen sulfide-stripping zone, the bottom fraction from which, as previously stated, represents the normally liquid product of the process. With respect to the naphtha boiling range portion of the product effluent, the sulfur concentration is about 0.1 p.p.m. by weight, the aromatic concentration is about 71.1% by volume, the bromine number is about 0.1 and the diene value is essentially nil.

In some situations, it may be desirable to recycle a portion of the hot separator liquid to the inlet of the heater employed to increase the temperature of the hydrogen-rich gaseous phase required to strip polymerization material from the catalyst in the guard chamber. This recycled diluent also facilitates the removal of polymer from the catalyst. In another embodiment, a portion may be recycled directly to the guard chamber undergoing stripping in order to cool the same prior to swinging it into the system. For the most part, however, the recycled diluent is employed for the dual purpose of inhibiting the temperature increase within the reaction system and to provide a cleaner combined feed stock to the second and third reaction zones. Since the product eluent emanating from the third catalytic reaction is (l) substantially completely saturated with respect to olenc hydrocarbons, (2) virtually completely desulfurized, it constitutes the source of the diluent employed in the present continuous hydrorening process. The techniques of separating the effluent in the hot separator, at a substantially reduced temperature in the range of about 200 F. to about 500 F., and preferably from about 250 F. to about 400 F., provides a hot recycle stream consisting essentialy of the higher-boiling liquid components of the third reaction zone etllucnt. Thus, the recycle diluent comprises principally Cc-Cg aromatic hydrocarbons which are substantially free from olelins, dissolved hydrogen sulfide, ammonia and lighter parafnic hydrocarbons. The benelits accruing as a result of utilizing the hot recycle diluent, in addition to considerations involved with heater duties and heat-exchangers, will be evident to those having expertise in the art of petroleum refining technology. However, it must be noted that there is also effected a decrease in the quantity of mon-olens in the total liquid feed to the third catalytic reaction zone. Thus, excessively great temperature increases are avoided leading to increased catalyst stability and a more economical operation from the standpoint of the size and character of quench facilities which might otherwise be provided.

DESCRIPTION OF CATALYTIC COMPOSITES The catalytic composites employed in the present process comprise metallic components selected from the metals,

and compounds thereof of Group VI-B, and I-A and VIII of the Periodic Table. Thus, in accordance with the Periodic Table of the Elements, E. H. Sargent & Company, 1964, suitable metallic components include lithium sodium, potassium, rubidium, cesium, molybdenum, tungsten, cobalt, nickel, ruthenium, rhodium, palladium, osmium iridium and platinum.

While neither the precise composition, nor the method of manufacturing in the catalyst is essential to my invention, certain considerations are preferred. For example, since the charge stocks to the present process are generally naphtha boiling range fractions, and the desired normally liquid product eiiluent is a naphtha boiling range fraction, it is preferred that neither catalytic composite exhibit an excessive degree of hydrocracking activity, under the operating conditions utilized herein, to the extent that the naphtha boiling range material is converted into lower-boiling, normally gaseous hydrocarbon products. Furthermore, since a principal object of the present invention resides in the retention of aromatic hydrocarbons, excessive hydrogenation activity is to be avoided. Although an acidic function may be incorporated within the catalytic composite utilized in the third catalytic reaction zone, in order to facilitate the destructive conversion of nitrogeneous and sulfurous compounds, the catalytic composite disposed within the first, second and fourth reaction zones is, for the most part, non-acidic. Thus, with respect to the latter, the catalytically active metallic components are preferably combined with a non-siliceous substantially halogen-free carrier material such as alumina. Substantially halogen-free composite is one wherein a halogen component is not intentionally added and, where a halogen compound (chloropalladic acid) is utilized in the catalyst manufacturing procedure, steps are intentionally taken to remove halogen from the resulting composite.

The catalytic composites disposed within the first, second and fourth catalytic reaction zones are preferably of the same composition and character. This catalyst might be said to serve a dual-function; that is, it must be nonsensitive to the presence of sulfurous compounds at the operating conditions employed, while at the same time capable of effecting the hydrogenation of conjugated dioleinic hydrocarbons to the corresponding saturates, while simultaneously possessing a degree of selectivity such that the mono-oleiins and aromatic hydrocarbons are not substantially saturated. A catalyst comprising an alumina-containing inorganic oxide, combined with a Group VIII noble metal component and an alkali metal component, is very etlicient in carrying out the desired operations. In contrast to the catalytic compositie utilized in the third catalytic reaction zone, hereinafter described, it is preferred that this catalyst be substantially free from acid-acting components, especially a halogen component As hereinbefore stated, a halogen compound is often employed during one or more steps of the overall catalyst manufacturing technique. For example, alumina is commonly prepared by a method which involves digesting substantially pure aluminum metal in hydrochloric acid, and the Group VIII noble metal is often impregnated throughout the finished alumina through the use of, for example, chloropalladic or chloroplatinic acid. It is Well 'known to be extremely diflicult to remove combined halogen from the iinished catalyst to a level lower than about 0.1% by weight. The presence of this halogen, which imparts undesired acidity to the catalytic composite, is countered and inhibited through the use of the alkali metal component.

The carrier material may be prepared in any suitable manner, and may be naturally-occurring. Following its preparation, the carrier material may be formed into any desired shape including spheres, pills, cakes, extrudates, powders, granules, etc. Neither the form, nor the method of manufacturing the carrier materials is considered to be an essential feature of my invention. One component of the catalyst disposed in the first, second and fourth reaction zones, is an alkali metal component employed for the purpose of attenuating the inherent acidity possessed by residual halogen, or by the carrier material itself. ,Suitable alkali metals are selected from the group of lithium, sodium, potassium, rubidium, cesium and mixtures thereof, particularly preferred metals being lithium and/or potassium, Regardless of the particular state in which the alkali metal component exists, the quantities thereof, from about 0.05% to about 1.5% by weight, are calculated as if this component existed in the elemental state. Generally, it is preferred to incorporate the alkali metal cornponent during the separation of a carrier material; therefore, the carrier material is often refered to as, for example, lithiated alumina.

At relatively low temperatures, the Group VIII noble metals possess the propensity for effecting the virtually complete hydrogenation of the reactive di-oleiins and styrenes, the latter selectively to alkyl benzenes. The noble metal components, selected from the group of ruthenium, osmium, rhodium, iridium, palladium, platinum and mixture thereof, are utilized in an amount of from about 0.01% to about 2.0% by weight, calculated as if existing in the elemental state. Group VIII noble metal components may be incorporated within the catalytic composite in any suitable manner including co-precipitation with the carrier, ion-exchange or impregnation of the carrier material with a suitable water-soluble compound of the metal. Following the incorporation of the metallic components, for example, by way of impregnation, the carrier material is dried at a temperature of about 200 F. to about 600 F., and subsequently calcined, in an atmosphere of air at an elevated temperature of about 700 F. to about 1200 F.

With respect to the catalytic composite disposed within the third reaction zone, for the primary purpose of destructively removing sulfurous compounds through the conversion thereof to hydrogen suliide and hydrocarbons, it should be noted that this reaction zone functions at a higher temperature level in the range of about 500 F. to about 1,000 F.

When the catalytically active metallic component is a Group VIII noble metal component, the maximum catalyst bed temperature will be in the range of about 800 F. to about 1,000 F. When an ion-group metal component is employed, either lone, or in combination with a Group VLB metal component, the lower temperature range of about 500 F. to about 800 F. will be employed. Although the third reaction zone catalytic composite may be similar to that utilized in the lirst major conversion zone, it is also distinctly different therefrom. For example, an alkali metal lcomponent is `generally not combined therewith, and a halogen component is often combined therewith. The halogen may be either fluorine, chlorine, iodine, bromine, or mixtures thereof, with fluorine and chlorine being preferred. When utilized, the halogen component will be composited in such a manner as results in a iinal composite containing about 0.1% to about 1.5% by weight, and preferably from about 0.4% to about 0.9% by weight, calculated on an elemental basis. With respect to the alumina-containing carrier material, the alumina may be advantageously employed in and of itself, or in combination with minor quantities of silica or other refractory inorganic oxides. When combined with, for example, silica, it is preferred that the alumina/silica weight ratio be within the range of from about 63/ 37 to about /10. When the third reaction zone catalytic composite contains a Group VIII noble metal component, it is also selected from the group of ruthenium, rhodium, palladium, osmium, iridium and platinum. Of these, a palladium and/or platinum metallic component is especially preferred. This component may exist within the final catalyst composite as a compound, including the oxide, sulfide, halide, etc., or in an elemental state. The Group VIII noble metal component generally comprises about 0.01% to about 1.0% by weight of the final catalytic composite, calculated on the basis of the element.

When an ion-group metal component, particularly nickel and/or cobalt, is utilized, it will generally be present in an amount of about 1.0% to about 10.0% by weight, again calculated as the elemental metal. When utilized, the Group VI-B metal component, particularly molybdenum and/or tungsten will be incorporated in an amount of from about 4.0% to about 20.0% by weight.

With respect to interchanging or swinging the two reaction zones serving as guard chambers, referred to as the first and fourth reaction zones (in the drawing,

reactors

18 and 12, respectively), a preferred technique is to effect the interchange at such times as the catalyst undergoing regeneration becomes substantially free from olefin polymerization products. This point will be noted by monitoring the material entering the knock-out pot for separation and removal of the polymer products from the hot hydrogen gaseous phase. When this stream indicates that the olefin polymer products have been stripped from the catalyst, the guard chamber may `be interchanged with the first reaction zone within the process. Simultaneously, the hot hydrogen becomes diverted through the guard chamber just removed from the integrated process system. When this catalyst is substantially free from olefin polymer products, the interchange is again effected. Another scheme involves monitoring the di-olefin content of the first reaction zone efiluent. As the di-olefin content increases, there is an indication that the catalytic composite is losing activity as the result of the formation of olen polymers.

DESCRIPTION OF DRAWINGS My invention, as directed toward the multiple-stage hydroening of coke-forming hydrocarbon distillates, may be more clearly understood upon reference to the accompanying drawing which illustrates one embodiment thereof. In the drawing, various ow valves, control valves, instrumentation and start-up lines, coolers, pumps and/or compressors, heat-exchangers, etc., have either been eliminated, or reduced in number; only those connecting lines necessary for a complete understanding of the continuous process are shown. The use of other miscellaneous appurtenances is well within the purview of one having skill in the art of petroleum processing techniques.

The drawing will be described in conjunction with a commercially-scaled unit designed to effect the multiplestage catalytic hydrotreating of a thermally-cracked pyrolysis gasoline. The charge stock has a gravity of about 45.0 API, an initial boiling point of about 132 F. and an end boiling point of about 350 F. The charge stock contains about 700 p.p.m. by weight of sulfur and about 50.0% by volume of aromatic hydrocarbons, and indicates a bromine number of about 60.0 and a diene value of about 50.0. The fresh feed, in amount of about 2,614 bbL/day, enters the process by Way of line 1, and is admixed with a liquid recycle diluent from line 2, the latter in an amount of 3,721 bbl/day. The mixture continues through line 2 and, via line 3, containing valve 5, and line 36, into reactor 18. In the present illustration, it will be presumed that reactor 18 is in operating status within the system, whereas the catalyst in reactor 12 is undergoing regeneration. Therefore, with respect to the various valves in the manifolding attendant these two reaction zones, the open valves are designated as 5, 9, 22, 14, and 7, with the closed valves being 4, 3, 20, 11 and 17.

A hydrogen-rich recycle gaseous phase in line 33 is introduced into heater 35, wherein the temperature is increased to a level of about 690 F. The heated hydrogen passes through line 36 and open valve 7 in line 6 into reactor 12. Hydrogen and stripped olefin polymers from the catalyst disposed in reactor 12 are withdrawn by way of line 13, containing open valve 14, and introduced thereby into knock-out pot 15. Polymer products are withdrawn from the system by way of line 38 containing open valve 39. The hot hydrogen is withdrawn by way of line 37 and, after cooling to a temperature of about 265 F., passes through line 10 and open valve 9 into line 36 wherein it is admixed with the fresh hydrocarbon charge stock and liquid diluent. Guard chamber 18 functions at a maximum catalyst bed temperature of about 270 F., a liquid hourly spaced velocity of about 17.0, based upon fresh feed, and a pressure of about 860 p.s.i.g. The catalytic composite disposed in reactor 18, as well as in reactors 12 and 17, is a composite of alumina, 0.5% by weight of lithium and about 0.4% by weight of palladium. The effluent from reactor 18 is withdrawn by way of line 19, passes through open valve 22 and line 21, and, by way of line 16, is introduced into reactor 17 at a pressure of about 830 p.s.i.g.

Reactor 17 functions at a maximum `catalyst bed temperature of about 360 F., representing an increasing temperature gradient of about F., and at a liquid hourly space velocity of about 1.8. With respect to the unit being illustrated, reactor 17 contains approximately 305 cubic feet of catalyst while

reactors

18 and 12 each contain amout 34 cubic feet of catalyst. Thus, with respect to reactors 17 and 18, containing a total of 339 cubic feet, reactor 18 contains about 10.0% of the total catalyst in the two reactors. The effluent from reactor 17 is withdrawn by way of line 23, and introduced thereby into heater 24, in admixture with make-up hydrogen from line 26, wherein the temperature is increased to a level of about 525 F. The heated mixture passes through line 25, and is introduced into reactor 27 at a pressure of about 790 p.s.i.g. Reactor 27 contains a catalytic cornposite of alumina, 5.0% by weight of nickel and 10.0% by weight molybdenum and in amount such that the liquid hourly space velocity therethrough, based upon fresh feed, is about 2.0.

The vincreasing temperature gradient `is maintained at a level of about 75 F., resulting in a maximum catalyst bed temperature of 600 F. The reactor efiiuent is withdrawn by way of line 28, and, following its use as a heatexchange medium to decrease its temperature to about 300 F., is introduced into hot separator 29 at a pressure of about 770 p.s.i.g.

A principally liquid phase is withdrawn from hot separator 29 by way of line 34, and about 3,721 bbl/day is diverted through line 2 to combine with the fresh feed charge stock, in line 1. Where desired, a portion of the hot separator liquid may be recycled through line 34 to be combined with the regenerating hydrogen stream in line 33. A portion of the hot separator liquid may be recycled to the third catalytic zone.

The principally gaseous phase from hot separator 29 is withdrawn by way of line 30, and, following its use as a heat-exchange medium and further cooling, is introduced into cold separator 31 at a temperature of about 100 F. and a pressure of about 750 p.s.i.g. The hydrogen-rich recycle stream, for use in regenerating the catalyst in reactor 12, is withdrawn by way of line 33, and introduced thereby into heater 35. The normally liquid hydrocarbons stream is withdrawn from cold separator 31 through line 32, and is admixed therein with the excess liquid phase from hot separator 29 in line 42, the mixture constituting the normally liquid product of the process.

Prior to being sent to an aromatic recovery system, the product in line 32 is introduced into a hydrogen sulfide stripping column, the top temperature of which is about 250 F., the top pressure Ibeing about 190 p.s.i.g. and the bottoms temperature being about 430 F. At these conditions, the bottom stream is recovered substantially free from butanes and lighter hydrocarbons, and can be transported directly to the aromatic recovery system. With respect to as-produced aromatic concentrate, the sulfur concentration is effectively nil being only about 0.1 p.p.m. by weight. The diene value is nil, the bromine with re-running and depentanizing, the catalytic composite in reactors 17 and 27 can be expected to attain a catalyst life of 140 and 150 `bbL/lb., respectively; through the use of the present invention, the expected ultimate catalyst life is increased to about 280 and 300 bbL/lb., respectively, and no re-running or depentanizing is required. The erected cost of the reaction system, including heaters, where re-running and depentanizing mustl be provided, is about $500,000.00. The erected cost of a unit, incorporating the present invention, including heaters, is about $420,000.00, or about $30,000.00 per 1,000 bbl. of charge stock.

number indicative of the quantity of mono-olefins, is less than about 0.3 and about 98.0% recovery of the aromatics, by weight, is attained.

Referring once again to the guard chambers,

reactors

12 and 18, during the initial stage of regenerating catalyst in the former, valve 41 in line 40 will be open in order 5 to recover residual light liquid material which can be introduced into reactor 27 by way of line 25. When the gums and other oleins polymer appears in the eiuent fromknock-out pot 15, valve 41 is closed, valve 39 is open, and the polymerization products are withdrawn from the process. When the catalytic composite disposed in reactor 12 is substantially free from olen polymer products and reactor 12 has been cooled to the process temperature, the following valves are opened: valve 4, line 3, valve 11, line 36, valve 8 in line 10, valve 20 in line 19, and valve 17 in line 16; and, the following valves are closed in order to regenerate the catalyst in reactor 18: valve 7 in line 6, valve 9 in line 10, valve 5 in line 3, valve 14 in line 13, and valve 22 in line 21. The procedure is, of course, reversed after the catalytic composite disposed within reactor 18 is regenerated to ythe extent of being substantially free from olen polymer products.

In the absence of the swing-bed guard chambers but 'I'he foregoing speciiication, and example integrated into the description of the drawing, clearly illustrates the method of effecting the present invention and the benelits to be afforded through the utilization thereof.

I claim as my invention: 1. A continuous process for hydrorening a sulfurous,

(a) reacting said charge stock and a hydrogen-rich gaseous phase as hereinafter identified in a first catalytic reaction zone with a catalyst composite which is non-sensitive to sulfur and which is present therein in an amount of about 5.0% to about 25.0% by weight of the total catalyst in the rst and second reaction zones, in said first catalytic reaction zone at least partial hydrogenation of di-olelinic hydrocarbons is eiected and olein polymer formation occurs, at a maximum catalyst bed temperature in the range of about 200 F. to about 500 F.;

(b) further reacting the resulting first zone euent, at substantially the same temperature in the range of about 200 F. to about 500 F. in a second catalytic reaction zone containing said catalyst composite wherein at least partial hydrogenation of di-oleiinic compounds is substantially completed;

(c) increasing the temperature of the resulting second reaction zone effluent to a level above about 500 F., and reacting the thus-heated second zone eiuent in a third catalytic reaction zone containing a hydrodesulfurization catalytic composite wherein monoolen hydrogenation and desulfurization are eiected at a maximum catalyst bed temperature below about 1,000 F.;

(d) separating the resulting third reaction zone eiuent, 75

at a temperature in the range of about 200 F. to about 500 F., in a first separation zone, to provide a rst vaporous phase and a first liquid phase;

(e) further separating said first vaporous phase, at a substantially lower temperature, in a second separation zone, to provide a second liquid phase and a hydrogen-rich second vaporous phase;

(f) heating said hydrogen-rich second vaporous phase to a temperature in the range of 500 F. to about 700 F.;

(g) passing the thus-heated second vaporous phase through a regenerating guard chamber as a fourth catalytic reaction zone containing an olefin polymerdeactivated catalytic composite;

(h) separating the resulting fourth reaction zone efuent, in a third separation zone, to provide a polymer concentrate and a hydrogen-rich third vaporous phase; and,

(i) recycling said third vaporous phase to combine with said charge stock as said hydrogen-rich gaseous phase in step (a).

2. The process of claim 1 further characterized in that said rst and fourth catalytic reaction zones are interchanged when the deactivated catalytic composite in said irst reaction zone is substantially free from olefin polymers, so that said charge stock and regenerating hydrogen gaseous phase are introduced into said fourth reaction zone, and said heated second vaporous phase is introduced into said first catalytic reaction zone to remove olefin polymers from the catal'y'st disposed therein.

3. 'I'he process of claim 2 further characterized in that said first and fourth reaction zones are interchanged when the catalyst in said first reaction zone is substantially free from olelin polymers.

4. The process of claim 1 further characterized in that at least a portion of said rst liquid phase is recycled to combine with said charge stock.

5. The process of claim 1 further characterized in that at least a portion of said tirst liquid phase is recycled to combine with said second vaporous phase.

6. The process of claim 1 further characterized in that at least a portion of said rst liquid phase is recycled to said third catalytic reaction zone.

7. The process of claim 1 further characterized in that the catalyst disposed in said iirst, second and fourth reaction zones is a composite of a Group Vllll noble metal component, an alumina-containing non-acidic carrier material and an alkali metal component.

8. The process of claim 1 further characterized in that the catalyst bed temperature in said third reaction zone is in the range of 500 F. to about 800 F. and the catalyst disposed therein contains an iron-group metal component.

l9. The process of claim 1 further characterized in that the maximum catalyst bed temperature in said third reaction zone is in the range of 800 F. to about l,000 F. and the catalyst disposed therein contains a Group VIII noble metal component.

References Cited UNITED STATES PATENTS 3,457,163 7/ 1969 Parker 208-211 2,833,698 5/ 1958 Patton et al. 208-211 3,239,449 3/1966 Graven et al. 208-210 2,833,697 5/1958 Oettinger 208-211 3,167,498 1/ 1965 Leverkusen et al. 20S-210 3,492,220 1/ 1970 Lempert et al. 20S-210 DELBERT E. GANTZ, Primary Examiner G. I. CRASANAKIS, Assistant Examiner U.S. C1. XJR. 208--97