Search Images Maps Play YouTube News Gmail Drive More »
Sign in
Screen reader users: click this link for accessible mode. Accessible mode has the same essential features but works better with your reader.

Patents

  1. Advanced Patent Search
Publication numberUS3699036 A
Publication typeGrant
Publication dateOct 17, 1972
Filing dateAug 21, 1970
Priority dateAug 21, 1970
Publication numberUS 3699036 A, US 3699036A, US-A-3699036, US3699036 A, US3699036A
InventorsHass Robert H, Reeg Cloyd P
Original AssigneeUnion Oil Co
Export CitationBiBTeX, EndNote, RefMan
External Links: USPTO, USPTO Assignment, Espacenet
Hydrocracking nitrogen containing feedstocks
US 3699036 A
Images(8)
Previous page
Next page
Description  (OCR text may contain errors)

Patented Oct. 17, 1972 3,699,036 HYDROCRACKING NITROGEN CONTAINING FEEDSTOCKS Robert H. Hass, Fullerton, and Cloyd P. Reeg, Orange, Calif, 'assignors to Union Oil Company of California, Los Angeles, Calif. No Drawing. Filed Aug. 21, 1970, Ser. No. 66,062 Int. Cl. Cg 13/02, 13/04, 13/10 US. Cl. 208-111 14 Claims ABSTRACT OF THE DISCLOSURE Hydrocarbons containing substantial amounts of organo-nitrogen compounds are converted in the presence of hydrogen over a catalyst comprising alumina, aluminosilicates, or combinations thereof with silica, zirconia, magnesia and the like with or without active metal components in the presence of controlled minor amounts of added water.

BACKGROUND Recent years have witnessed a phenomenal growth in the development and application of catalytic hydrocarbon conversion processes including hydrocracking, hydrofining, hydrogenation and combinations of these. By far the greatest part of these eiforts have been directed toward methods for hydrocracking gas oils to products boiling in the gasoline range. The best catalysts developed to date for this purpose are those comprising a highly active cracking base, e.g., of the crystalline zeolite type, combined with a highly active hydrogenation component such as palladium or platinum. These catalysts are'highly efiicient for converting middle distillate oils boiling in the 400-750 F. range to gasoline. Due to the extensive present and prospective use of middle distillate stocks as feed to these gasoline-producing hydrocracking units, and to other economic, geographic and seasonal factors, a need is being felt in the industry to provide additional middle distillate stocks to meet the demand for other products such as turbine and diesel fuels. An obviously desirable source of such additional middle distillate stocks comprises the heavy distillates boiling above about 700 F., which have heretofore been diverted largely to fuel oils because of the lack of an economical method for converting them to lower boiling products.

In the initial investigation of hydrocracking techniques for converting hydrocarbon feeds boiling above 700 F. it became apparent that severe hydrocracking to produce gasoline in a single pass conversion step was impractical, firstly because inordinate amounts of butanes and dry gas were produced, and secondly because the gasoline product was always of undesirably low octane quality. It thus became apparent that for purposes of gasoline production, especially on a once through basis, hydrocracking techniques were of primary value in connection with middle distillate feedstocks, and that the hydrocracking of higher boiling stocks would be economically desirable only if conversion to gasoline could be minimized and the production of middle distillates maximized. A primary objective in the hydrocracking of these heavy feeds was thus the realization of maximum middle distillate/ gasoline ratios in the resulting product.

It was then found that the catalysts most useful for converting middle distillates to gasoline were least useful for converting heavy feedstocks selectively to middle distillates in that large amounts of dry gases, butanes and gasoline were produced by those catalysts. A great many conventional hydrocracking catalysts were tested in an attempt to find one which would convert such feedstocks more selectively to middle distillate products. In all cases it was found that the desired selectivity could be maintained only by operating at very low overall conversion per pass, entailing prohibitively high recycle oil rates. However, upon developing and testing certain catalysts not conventionally regarded as hydrocracking catalysts, substantially more promising results were obtained. Specifically, it was found that by using a type of catalyst commonly employed for catalytic hydrofining composed of nickel and molybdenum oxides and/or sulfides supported on activated alumina, the selectivity of conversion to middle distillate products was excellent even at relatively high conversions, e.g., 40-60 volume-percent per pass. Moreover, for feedstocks containing organic nitrogen, it was found that these catalysts were not only more selective, but sometimes more active (on the basis of a lower temperature required for a given total conversion) than even the most active hydrocracking catalysts based on zeolite cracking bases.

Although those catalysts are particularly attractive for converting relatively high boiling nitrogenous feeds to middle distillates and high quality midbarrel fuels, they also have considerable utility in the conversion of similar feedstocks to high octane gasoline boiling range hydrocarbons. However, as is always the case with chemical conversion systems, there remains considerable room for improvement in numerous limiting process parameters such as catalyst activity, product distribution, catalyst life and the like. For example, it is known that catalyst life can be dramatically increased by lowering reaction temperature. However, such procedures decidedly limit the obtainable conversion'per pass requiring greater capital and operating expense due to the requirement of larger reactor sizes and recycle systems. Therefore, a reasonable balance must be drawn between catalyst deactivation rate, reaction temperatures and tolerable conversion levels that enable efiicient use of process equipment.

Yet another consideration which limits the effectiveness of these catalyst systems for hydrocarbon conversion, particularly hydrocracking and hydrogenation, is the deleterious effect exhibited on catalyst activity by organonitrogen compounds in the feed. Although some of these catalysts are particularly elfective for hydrofining, e. g., denitrogenation and desulfurization, we have observed that their hydrocracking and hydrogenation activity is markedly limited by the presence of even minor amounts of organonitrogen compounds when the concentration of those materials in the feed substantially exceeds 1 p.p.m. nitrogen. The most direct approach to the solution of this problem would appear to be the use of a hydrofining zone upstream of the hydrocracking system or the use of a larger hydrocracking catalyst bed. In these systems, and in other apparent alternative schemes, sufficient precontacting of the nitrogen containing feed with these or other hydrofining catalysts under hydrofining conditions can be effected to reduce the nitrogen content to tolerable levels, e.g., less than about 5 p.p.m. Although such procedures undoubtedly provide an available alternative, they are subject to certain obvious disadvantages. In order to reduce the nitrogen content of the feed to a level sufficient to prevent inhibition of hydrogenation and hydrocracking, provision must be made for rather severe hydrofining treatment upstream of the contacting zone. Pretreatments of this nature are known to involve consider expense.

Although organonitrogen compounds are known to exhibit several deleterious eifects on the preferred midbarrel hydrocracking and hydrogenation catalysts such as those comprising Group VI and Group VIII metal components on a primarily alumina support, their influence on the activity, selectivity and elfective catalyst life of the zeolite based catalyst particularly those containing Group VIII noble metals such as platinum and palladium, is even more dramatic. For example it is known that with some of the zeolite catalysts primarily employed for gasoline production that the presence of substantially more than parts per million of organonitrogen compounds in the feed determined as elemental nitrogen almost completely inhibits the hydrocracking and hydrogenation activity of those compositions, at least in the upstream sections of the contacting zone, until sufiicient contacting has been effected to reduce the organonitrogen content to tolerable levels. As a general rule, the zeolite base catalysts are not particularly effective denitrogenation catalysts and many are permanently inhibited by exposure to substantial amounts of organonitrogen compounds. As a result, it is a matter of almost necessity in the use of such composition to provide for the almost complete 'den-itrogen'ation of a selected feedstock prior to hydrogenation or hydrocracking over these catalysts. It is usually necessary to reduce the organonitrogen content of such feeds to a point below about 5'p.p.m. and inmany cases, nitrogen levels substantially in excess of 1 p.p.m. cannot be tolerated. Obviously, the expense and process complexities involved in effecting these degrees of vdenitrogenation are considerable. However, to date they have been considered acceptable in view of the lack of available alternatives and the relative degree of comparative success realized by operating in that manner.

We have now discovered a procedure whereby several aspects of such hydrocarbon conversion systems can be markedly improved from the standpoint of performance and economy. This procedure enables the attainment of economically feasible conversion levels and product quality at less severe reactor conditions, e.g., lower reaction temperatures, without the necessity of severe prehydrofining. We have discovered that the addition of controlled amounts of water or water precursors to the hydrocarbon feed counteracts the effect of substantial nitrogen levels and enables the catalysts to effect acceptable conversion levels at reduced temperatures. Several of the advantages of these procedures are readily apparent. Firstly, longer catalyst life can be realized due to the lower reaction temperatures required. Secondly, operating and capital expense involved in upstream hydrofining zones can be reduced due to the tolerance of the described conversion systems to higher concentrations of nitrogen compounds. In addition, the activity of these systems for several forms of hydrocarbon conversion is substantially improved, thereby enabling higher conversions, lower operating temperatures, greater control of product distribution, and the like. Particularly dramatic efi'ects are evidenced in hydrocracking activity and hydrogenation, particularly the hydrogenation of aromatic constituents. Consequently, these systems can be operated under conditions which provide markedly superior results in the production of those gasolines and midbarrel fuels and even enable the production of such products from relatively highly aromatic refractory feedstocks under recycle conditions while minimizing the degree of refractory buildup in the recycle stream.

It is therefore one object of this invention to provide an improved hydrocarbon conversion process. It is another object of this invention to provide an improved process for producing middle distillate fuels and/ or gasoline. It is yet another object of this invention to provide a method for increasing the activity of hydrocracking systems. Yet another object of vthis invention is the provision of a method for increasing the tolerance of hydrocracking and hydrogenation systems to organic nitrogen compounds. Still another object of this invention is the provision of improved hydrogenation, hydrocracking and hydrofining methods. Another object of this invention is the provision of a method which enables the reduction of hydrocracking temperatures necessary to obtain prescribed conversion levels.

In accordance with one embodiment of this invention, hydrocarbons containing at least about 5 p.p.m. nitrogen as organonitrogen compounds are contacted in the presence of hydrogen under hydrocarbon conversion conditions with a catalyst comprising at least one of the metals, oxides and sulfides of Groups II, IV-B, VI-B, VII-B and VIII combined with a foraminous oxide support in the presence of a promoting amount of added water or water precursor.

In accordance with another embodiment of this invention, hydrocarbons containing at least about 5 p.p.m. nitrogen as organonitrogen compounds are contacted in the presence of hydrogen at elevated temperatures and pressures sufficient to promote the reactions of said hydrocarbons with the added hydrogen in the presence of a catalyst comprising at least one hydrogenation active metal, oxide or sulfide of Groups VI-B, VII-B and VIII combined with alumina, silica, magnesia, zirconia or clays comprising primarily alumina and/or silica and combinations thereof such as aluminosilicates in the presence of a controlled added amount of water equivalent to at least about 0.2 weight-percent and based on hydrocarbons.

During the investigation of this phenomena water was supplied to a hydrocracking zone at several different sets of operating conditions by several different procedures. In one instance water was admixed with a hydrocarbon phase by humidifying the hydrogen stream to the reactor. In another instance, the same results were effected by adding a Water precursor, e.g., n-butanol to the feed. Alcohols of this type are believed to be dehydroxylated to produce water under hydrocracking conditions. In every case, the addition of Water by these procedures improved the activity of the system. However, an even more surprising result was observedwhen the water injection was discontinued. Instead of the catalyst deactivating gradually to its original level as expected the hydrocracking performance exhibited a transient behavior in a reproducible cycle during which the activity improved still further reaching a maximum and then slowly declined to the activity exhibited by the system prior to water addition. This transient phenomenon was reversible and repeatable. However, despite all efforts, including periodic water injection, no means could be found whereby the maximum catalyst activity observed could be sustained on a continuous basis.

Although the exact mechanism by which this phenomenon is effected is not understood with certainty, it may be that the presence of watermitigates the inhibiting effect of organic nitrogen compounds by either preventing access of those poisons to the catalyst surface or by increasing the tolerance of the catalyst to those impurities. These suppositions are supported somewhat by observations made during operations on hydrocarbon feedstoeks having very low nitrogen concentrations, i.e., below about 1 p.p.m. organic nitrogen. In those systems, the catalystactivity was seen to decline upon water addition as would be expected since it is generally believed that water itself is a catalyst inhibitor. Nevertheless, in view of these observations, it seems reasonable to conclude that the phenomena herein described is at least in part a function of the combined effects of water and nitrogen.

The activity exhibited by these catalyst systems at a given set of process conditions is greater in the absence of either nitrogen or water. Both of these materials are generally considered catalyst inhibitors in otherwise clean systems. However, we have observed that considerable advantage can be realized by the addition of water to hydrocracking zones operating on hydrocarbon mixtures containing substantial amounts of organonitrogen compounds, i.e., 5 p.p.m. or greater.

Those skilled in the art readily recognize that the prehydrofining severity required to reduce the nitrogen content of hydrocarbon mixtures to levels below 5 p.p.m. is not easily attained and involves considerable expense. The method of this invention offers an alternative to such severe hydrofining procedures. It is possible to drastically reduce the severity and expense of prehydrofining treatments and compensate for the residual nitrogen in the hydrocarbon feed by conducting the conversion in the presence of controlled amounts of water as described. In fact, the advantages of this invention can be realized in the absence of any hydrofining whatever, provided of course that the nature of the feedstock is not such that it completely destroys the activity of the catalyst.

A wide variey of hydrocarbon feeds can be employed within the concept of this invention. Similarly, the conditions under which the feed hydrocarbons are reacted and the compositions of the catalysts employed in any specific case can be varied to promote selected reaction mechanisms such as hydrogenation, hydrofining, hydrocracking, and the like. The advantages of this invention can be realized in all of these situations when the feed employed contains a substantial amount of organonitrogen compounds. As a general rule, the hydrocarbon feed employed in these systems will boil above about 200 F. and will include light, medium and heavy naphthas, gas oils, coker distillates, vacuum gas oils, and the like. These systems may be designed for the simple hydrogenation of aromatic feedstock of essentially any molecular weight. However, particular advantage is realized in the hydrogenation of aromatic feedstocks particularly those boiling above about 400 F. and containing in excess of about and usually about to about 200 p.p.m. nitrogen as organonitrogen compounds which we have found to markedly inhibit the aromatics hydrogenation activity of the catalysts encompassed within the scope of this invention. However, the presently most preferred application of this phenomena is in the hydrocracking of raw or partially refined hydrocarbon feeds to produce gasolines and midbarrel fuels. We have observed that the addition of the described controlled water levels to hydrocracking systems dramatically improves the activity thereof and enables much greater control of product distribution than is otherwise possible. The feedstocks employed in such hydrocracking systems boil above about 400 F., usually within the range of about 400 to about 1200" F. Depending upon the degree of prerefining, the organonitrogen content of these feeds can vary substantially. Organonitrogen compounds are usually present in amounts equivalent to about 10- p.p.m. to about 1 weight-percent nitrogen, generally about 10 to about 5000 p.p.m. nitrogen. However, we have observed that the advantages of operating in the described manner are more apparent when the nitrogen content of the feed is within the range of about 10 to about 600 ppm. Consequently, it is presently preferred that feedstocks having nitrogen levels substantially in excess of 1000 p.p.m. be partially hydrofined prior to contacting in the hydrocracking systems operated in accordance with this invention.

Although the sulfur content of selected feedstock is not presently considered to be a determinative variable with regard to the results observed upon water addition, it is nevertheless a consideration relative to overall catalyst activity. The deactivating effect of excessive sulfur concentrations can be substantially mitigated by the use of sulfactive catalysts, which are generally well known in the art. Similarly, the presence of some sulfur compounds in the feed is often preferred in hydrocracking systems in that these materials tend to sulfide the catalyst maintaining it in the generally more active sulfided state. As a general rule, the raw hydrocarbon feeds may contain organic sulfur concentrations ranging up to about 3 weight-percent sulfur based on the total feed. It is often preferable to substantially reduce this sulfur content to a level below about 100 ppm. by pre'hydrofining. Such desulfurization can be conveneiently effected by contacting at conventional hydrofining conditions in the presence of catalysts similar to those herein described. The hydrogen sulfide resulting from the desulfurization of the organic sulfur compounds can be separated prior to introduction of the feed to the hydrocracking zone or can be passed directly over the hydrocracking or hydrogenation catalyst without substantial detriment.

A wide variety of catalysts can be employed within the concept of this invention. Essentially any catalyst comprising an active form of a metal or metal compound of Groups H, VI, VII and VIII of the Periodic Chart supported on a foraminous refractory oxide can be employed in these systems. Since catalysts of this nature are generally well known in the art they need not be described in detail herein. However, brief reference will be made to several of the presently most popular compositions within this class.

The foraminous refractory oxides above referred to can be selected from alumina, silica, magnesia, zirconia and alumina and silica containing clays such as bentonite and attapulgus clay and combinations thereof such as silica, alumina, silica zirconia and the like. One presently preferred combination of these materials is silica stabilized alumina containing about 2 to about 30 weightpercent silica. This material in itself is conventional in the art and can be prepared by several procedures such as cogelation or sequential gelation or physical admixture of partially or completely precipitated gels of alumina and silica. It has been observed that the presence of minor amounts of silica in alumina matrices greatly improves the physical properties of the resultant combination. The natural and synthetic crystalline and amorphous aluminosilicate Zeolites are also widely employed for the preparation of hydrocarbon conversion catalysts and are applicable to the systems herein described.

Exemplary of these compositions are the aluminum containing ion exchangeable clays, particularly acid extracted clays such as bentonite, montmorillonite, beidellite, halloysite, endelite, kalolinite and the like. The acid extraction of these materials is generally well known in the art and is designed to effect the removal of metals found in the aluminosilicates in their naturally occurring state. Such acid extraction can be effected by contacting the clay with a mildly acidic solution or a strong mineral acid such as nitric, sulfuric, hydrochloric, or phosphoric acids and the like, having a pH of about 4 or above.

A second generally recognized class of amorphous aluminosilicates having the above-described characteristics are the partially degraded crystalline aluminosilicates which have been subjected to acidic and/or thermal en vironments sufficient to at least partially destroy the characteristics crystalline structure of those materials. As a general rule, there is little economic interest in obtaining starting materials, i.e., aluminosilicates, by this latter procedure in that the crystalline aluminosilicates are usually far more expensive than alternative forms of aluminosilicates suitable for use herein. Nevertheless, such materials are suitable for application within the concept of this invention.

Another class of aluminosilicates having the desired characteristics are the silica-alumina cogels. These materials are usually prepared by either coprecipitating silica and alumina from an aqueous solution at soluble salts of silicon and aluminum or by grafting one of these constituents, i.e., either the silica or alumina, onto a previously precipitated alumina or silica gel by acidifying a solution of the water soluble salt of the second component. For example, alumina can be precipitated in the presence of a hydrous silica gel by acidifying a solution of sodium aluminate with a mineral acid such as sulfuric, nitric, and the like. The relative amounts of silica and alumina in the combinations can vary considerably although it is presently preferred that the silica concentration be equivalent to about 10 to about 40 weight-percent of the total silica-alumina combination on a dry weight basis. The cogels resulting from such coprecipitation or sequential precipitation of separate constituents can then be dried and ion exchanged to replace the undesirable cations with ammonium, hydrogen and/or Group VIII metal-containing cations as previously described.

The presently preferred aluminosilicates are the crystalline species having Slo /A1 ratios of at least about 2. This class includes both synthetic and naturally occurring zeolites. Illustrative of the synthetic zeolites are zeloite X, U.S. 2,882,244; zeolite Y, U.S. 3,130,007; zeolite A, U.S. 2,882,243; zeolite L, Belg. 575,117; zeolite D., Can. 611,981; zeolite R U.S. 3,030,181; zeolite S, U.S. 3,054,657; zeolite T, U.S. 2,950,952; zeolite Z, Can. 614,995; zeolite E, Can. 636,931; zeolite F, U.S. 2,995,358; zeolite O, U.S. 3,140,252 zeolite B, U.S. 3,008,- 803; zeolite Q, U.S. 2,991,151; zeolite M, U.S. 2,995,- 423; zeolite H, U.S. 3,010,789; zeolite J, U.S. 3,011,869; zeolite W, U.S. 3,012,853; zeolite KG, U.S. 3,056,654. Illustrative of the naturally occurring crystalline aluminosilicates which can be suitably treated by the methods herein described are levynite, dachiardite, erionite, faujasite, analcite, paulingite, noselite, ferrierite, haulandite, scolecite, stilbite, clinoptilolite, harmotome, phillipsite, brewsterite, flakite, datolite, chabazite, gmelinite, cancrinite, leucite, lazurite, scolacite, mesolite, ptilolite, mordenite, nepheline, natrolite, and sodalite. The natural and synthetic faujasite-type crystalline aluminosilicate zeolites, e.g., zeolites X and Y, arepreseutly particularly preferred.

When these catalysts are prepared in whole or in. part from zeolitic aluminosilicates, it is often desirable to substantially reduce the alkali metal content of the aluminasilicate during some stage of the catalyst preparation and convert the zeolite into an acidic form. Consequently, the alkali metals which are present both in natural and synthetic aluminosilicates in their original state are preferably removed during some stage of the catalyst preparation by ion exchange. Any one of numerous procedures can be employed for this purpose. Probably the most convenient procedures include ion exchange with aqueous solutions of ammonium or hydrogen or soluble salts of metal cations of Groups II, VII and VIII of the Periodic Chart. As a general rule, the alkali metal content of the aluminosilicates should be reduced to less than 5 weightpercent preferably less than about 2 weight-percent determined as a corresponding oxide. This objective can be realized by continuing the exchange with the cations above mentioned until the desired reduction in alkali metal content has been achieved. When exchange with the noted metal cations is employed, the resultant concentration of the respective metals in the final composition will usually be within the range of about 0.2 to about 5 weight-percent determined as the elemental metal. These metals can also be added to the aluminosilicate or other foraminous oxides by impregnating the support with an aqueous solution of a water soluble thermally deco-mp0sable salt of the desired metal. The Group VI-B metals can also be added by this procedure. Illustrative of suitable water soluble thermally decomposable compounds are the nitmtes, sulfates, carbonates and halides of calcium, magnesium, molybdenum, tungsten, manganese, rhenium, nickel, cobalt, platinum, palladium and the like.

Catalysts presently particularly preferred for the production of midbarrel fuels comprise composites of nickel and/ or cobalt metals, oxides and sulfides and one or more of molybdenum and tungsten metals, oxides and sulfides combined with an aluminum support.

The aluminas presently preferred are high surface area amorphous materials having relatively low cracking activities corresponding to Cat-A indices below 25. However, higher activities can be tolerated when higher relative gasoline production is desired. It is also preferred that the alumina support be stabilized by combination with a minor amount of amorphous silica gel. Silica can be present in concentrations of up to about 40 weightpercent based on the combined weight of silica and alumina. Presently preferred compositions have silica concentrations within a range of about 3 to about 25 weight-percent.

The presently preferred midbarrel catalyst compositions contain about 0.5 to about 10 Weight-percent nickel and/or cobalt metals, oxides or sulfides determined as the corresponding oxides, and about 4 to about 30 weightpercent of molybdenum and/or tungsten metals, oxides or sulfides based on the corresponding oxides. Combinations of nickel and molybdenum sulfides are particularly preferred. Such compositions are normally prepared by impregnating the selected alumina carrier with an aqueous solution or solutions of soluble salts of the respective metals followed by draining, drying and calcining in air at temperatures of 800 to about 1200 F. The calcined catalysts are preferably presulfided prior to contact with a feedstoc k as by reaction with a gaseous mixture of hydrogen and hydrogen sulfide, carbon disulfide, elemental sulfur and the like.

The preferred carrier for these midbarrel fuels catalysts is activated alumina gel containing a minor proportion of coprecipitated silica gel. The silica content should not exceed about 40 percent by weight. Higher silica contents tend to decrease cracking selectivity to midbarrel fuels. The preferred silica content is between about 3 and 25 percent by Weight. The preferred supports comprise at least about 60 weight-percent alumina, the remainder being silica, magnesia, aluminosilicates, and the like. Prior to impregnation, the carrier is preferably formed into pellets of about 4 to inch diameter (by extrusion or die-compression). Alternatively, the catalyst may be employed in a powder form.

Particularly preferred catalyst compositions are prepared by impregnation with the stabilized impregnating solutions discussed in U.S. Pats. 3,232,887 and 3,287,280. By these procedures the selected support is contacted with a solution of the desired metal salts containing stabilizing amounts of phosphate ions introduced by the addition of an acid of phosphorus, e.g., phosphoric acid, phosphorous acid and the like. Phosphorus concentrations in the final catalyst are at least about 0.2 weightpercent, usually within a range of about 0.5 to about 8 weight-percent, preferably about 2 to about 5 weightpercent based on the free metal.

Hydrocracking conditions generally considered to be most desirable include reaction temperatures Within a range of 500 to about 900 F., preferably 550 to 850 F., reactor pressures of at least about 500 p.s.i.g., preferably 1000 to about 5000 p.s.i.g., and liquid hourly space velocities (LHSV) of about 0.1 to about 10, preferably 0.3 to about 5. The free hydrogen content in the reaction zone should be equivalent to at least about 5 00, generally about 500 to about 20,000 and preferably about 1000 to about 15,000 standard cubic feet per barrel of reactor charge.

Hydrogenation and hydrofining, e.g. denitrogenation and desulfiurization can be effected at these same conditions and in fact, are etfected under hydrocracking conditions when there are constituents in the feed subject to hydrogenation and/hydrofining. However, it is sometimes desirable to effect denitrogenation, desulfurization and/ or hydrogenation in the absence of substantial degrees of hydrocracking. In those instances, it is desirable to contact the feed under somewhat milder conditions of reduced temperature and pressure, e.g., pressures-of about to about 1200 p.s.i.g. Catalyst compositions also determine the relative degrees of hydrocracking and hydrogenation and can be selected to favor one over the other as desired.

The equivalent water concentration in the reaction zone should be snfiicient to promote the activity of the catalyst as evidenced by higher conversion levels at otherwise identical conditions. We have found that the most dramatic improvement in catalyst activity is effected at equivalent water concentrations in excess of about 0.1 weight-percent, generally about 0.2 to about 5 weight-percent and preferably about 0.2 to about 2 weight-percent. Depending upon the characteristics of the particular systems such as reaction conditions and the selected catalyst and the organo nitrogen content of the selected feedstock, the preferred water level in the reaction zone may -vary considerably from about 0.1 to about weight-percent. A constant improvement in activity can be realized by maintaining a predetermined controlled water concentration in the reaction zone. We have observed that the rate of the increase in catalyst activity is related to the maximum equivalent water concentration in the reaction zone, certain reaction parameters such as liquid hourly space velocity and mixing characteristics and of course the duration of water injection into the reaction zone. Therefore, it is possible when employing the cyclic water injection procedure previously described, to extend the cycle length during which these beneficial effects are realized by operating at the higher water concentrations and/or lower liquid hourly space velocities.

An essential aspect of this procedure is that the prescribed controlled water concentration be provided in the reaction zone. This can be accomplished by the addition of either free water or a water precursor to the hydrocarbon feed, to one or more of the streams entering the reactor, or directly to the reactor. Exemplary of readily available water precursors are the organic alcohols. Presently preferred due to their availability and cost are those having from about 2 to about 12 carbon atoms and one to about five hydroxyl groups per molecule.

Although it is presently preferred that the water concentrations be maintained only in the hydrocracking zone, it is within the scope of this invention to provide for the addition of water to process systems upstream of the hydrocracking zone. The most apparent alternative in this regard is the addition of water or water precursor to the hydrocarbon feed entering a hydrofining stage immediately preceding the hydrocracking zone. This mode of operation is particularly convenient in integral hydrofining-hydrocracking systems employing catalysts similar to those herein described in both hydrofining and hydrocracking zones. Nevertheless, it is not necessary to introduce water into the hydrofining zone to take advantage of the results herein described. For example, in a single stage hydrofininghydrocracking system, the water or water precursor can be injected alone or in combination with recycle hydrogen or hydrocracker recycle oil at an intermediate point in the reaction zone. However, procedures employing this approach will generally involve the use of two or more reactors in one or more stages. The distinction between stages is conventionally characterized by the separation, or partial separation, of intermediate products such as ammonia and hydrogen sulfide from the first stage product prior to introduction of the hydrocarbon phase to the second stage. As a consequence, these systems can involve the use of a single reaction zone in which conditions are controlled such that hydrofining occurs in the first portion of the catalyst bed while hydrocracking is effected in the terminal portion of the catalyst bed. In the alternative, a single stage process can be employed using two or more reactors in which the upstream reactor are controlled so as to effect primarily hydrofining while the downstream reactors are operated to promote hydrocracking. In these more complex plural reactor systems the advantages of this invention can be realized by injection of water or suitable water precursors to one or more of the hydrocracking stages whether or not some hydrofining occurs in each respective stage.

Particular advantage can be obtained when employing multizone hydrocracking systems wherein the organic nitrogen content of the hydrocarbon phase is gradually reduced upon sequential passage through the several hydrocracking zones. As the nitrogen content of the hydrocarbon phase is reduced to a level below 5 p.p.m., the beneficial effects realized by water addition are mitigated. In fact, at nitrogen levels substantially below 5 p.p.m., particularly 1 p.p.m. or less, the addition of water is detrimental rather than beneficial due to its inherent catalyst inhibiting characteristics. Consequently, it is presently preferred to maintain the described water levels in the hydrocracking zones in which the nitrogen level in the feed is substantially above 5 p.p.m. while avoiding water addition and/or substantial carryover to the terminal hydrocracking zones or stages or otherwise reducing the water content of the hydrocarbon phase therein when the organic nitrogen content falls substantially below 5 p.p.m. nitrogen. By this procedure, the advantages of this invention can be most efliciently realized while preventing any deleterious effects attributable to the presence of water in the absence of substantial nitrogen concentrations.

As pointed out in the examples hereinafter detailed, the activity of the catalyst immediately subsequent to addition of water or water precursor gradually approaches an increased equilibrium level and remains at that level in the absence of significant catalyst deactivation as long as the water level is maintained at the same concentration. However, when Water addition is discontinued after the catalyst has reached this equilibrium activity, the hydrocarbon conversion activity of the system increases still further reaching a maximum and then gradually declines to the original activity of the system prior to water addition. In view of these observations, it is possible to control the catalyst activity in several ways. For example, the concentration of water in the hydrocracking zone can be maintained at a predetermined constant level throughout the run duration. However, in that mode of operation, the activity of the catalyst will reach an equilibrium intermediate level and continue at that level subject only to gradual catalyst deactivation usually observed in such systerns. An alternative procedure that provides substantially higher average catalyst activities involves periodic water addition. By this procedure, a predetermined water concentration is maintained until the catalyst approaches equilibrium. Water addition is then discontinued to increase the catalyst activity even further. Effective utilization of this procedure must take into account the amount of water added, the time required for the catalyst to approach its equilibrium activity and the cycle time of the catalyst activity increase and decrease subsequent to the discontinuance of water addition. The duration of the several cycles involved in this transient behavior can be readily ascertained by simply adding water to the conversion zone in the prescribed amounts and observing either the increased conversion realized at otherwise identical operating conditions or by reducing reaction temperature sufficiently to maintain a constant conversion level until the catalyst activity has equilibrated at its higher intermediate value. The time required for the completion of this response will vary somewhat with water concentration and with the prevailing liquid hourly space velocity. Nevertheless, the cycle time for the first step of the procedure can be readily ascertained for any given set of operating conditions by the procedure described.

After the catalyst has equilibrated at the intermediate activity level, water addition can be discontinued and the increase in catalyst activity can be observed in a manner similar to that above described until the activity preceding Water addition is reestablished. The data obtained by this procedure for any system will enable the most effective utilization of the observed phenomenon by directing the control of intermittent water addition depending upon the cycle time of the selected system. For example, it may be desirable to allow the catalyst activity to approach its original value subsequent to the discontinuance of water addition before reintroducing water to the hydrocracking zone. However, it is not necessary to defer the reintroduction of Water to that point. On the contrary, additional advantage can be obtained by reintroducing water to the system when the declining catalyst activity approaches the intermediate equilibrium value established during continuous water addition. By this procedure, the activity of the catalyst can be maintained at its equilibrium level or above by intermittently adding water or a selected Water precursor.

As previously observed, the cycle time required to traverse the complete cycle of this transient catalyst activity response depends on the characteristics of the catalyst employed, the hydrocarbon feedstock, the nitrogen content in the hydrocracking zone, the water level employed and the liquid hourly space velocity. However, as a general rule, when operating within the conditions above referred to, the time required to complete a cycle of this transient response is within a range of about 30 minutes to about 100 hours. This range is of course narrowed considerably for narrower ranges of operating conditions. For example, if the preferred liquid hourly space velocity is between about 0.3 to about 5, and the maximum water level is within a range of about 0.2 to about 2 weight-percent, the observed cycle time will usually be within the range of about 1 to about 60 hours. Consequently, it is presently preferred to continue Water addition to the hydrocracking zone until the conversion level is stabilized which usually necessitates continuance of water addition for at least about 10 minutes, preferably from about 1 to about 100 hours. The supply of added water to the reactor is then preferably discontinued for a period sufficient to enable the activity of the catalyst to pass through its maximum and decline to approximately its equilibrium value which involves periods of at least about 2 hours, usually about 1 to about 100 hours.

The-effectiveness of this procedure is demonstrated by the following example which is intended only to be illustrative of the concept of this invention and not limiting thereof.

EXAMPLE 1 The hydrocarbon feed employed in this example had the following characteristics.

TABLE 1 Boiling range, by D-1160, F 555-1016 Composition, by Univ. Hi Mass, weight-percent:

This feed was partially hydrofined by contacting with a catalyst containing 3 weight-percent NiO, 18 weightpercent M 3 percent P, 3 percent Si-O and 73 percent alumina. Operating conditions in the hydrofining zone included a reactor temperature of 740 F., a liquid hourly space velocity of 1.0, a total reactor pressure of 2500 p.s.i.g. and a hydrogen feed rate of 8000 s.c.f./bbl. feed. The product was then flashed to remove ammonia and hydrogen sulfide and passed to the hydrocracking zone. The described hydrofining procedures were sufficient to reduce the organic nitrogen content of the reactor charge to the hydrocracking zone to 50 p.p.m. and the organic sulfur content to 40 p.p.m.

Hydrocracking conditions included a total reactor pressure of 2500 p.s.i.g., a liquid hourly space velocity of 0.5 LHSV and a dry hydrogen feed rate of 8000 s.c.f./bbl. at an overall 40 percent conversion per pass. The second zone catalyst was identical to that employed in the above described hydrofining zone. Water was provided in the reaction zone by the addition of n-butanol to the hydrocarbon feed in amounts equivalent to 1.0 weight-percent water based on hydrocarbon.

Prior to water addition, the reaction temperature was maintained at 740 for a total of 11 days. This temperature was suflicient to provide a constant conversion per pass of 40 percent to products boiling below the initial boiling point of the feed. Water injection was then commenced and continued for a period of 168 hours. Immediately after water addition the catalyst activity increased as evidenced by a drop in temperature requirement needed to maintain the desired conversion per pass. This gradual increase in catalyst activity continued for 24 hours and then equilibrated at a level of 40 percent conversion per pass at 736 F. The introduction of water was discontinued after 168 hours. The catalyst activity again began to increase and continued to increase for a period of 12 hours at which time the maximum catalyst activity was obtained. The activity of the catalyst at this point was sufficient to obtain the same conversion per pass observed prior to water introduction at a reaction temperature of 710 F., 34 F. below the original reaction temperature. However, after this point, the activity began to decline and continued its decline gradually for about 24 hours until the system equilibrated at a conversion level of 40 percent conversion per pass at 744 F. These conditions were identical to those observed before water addition. No significant variation in product quality or distribution was observed at any time throughout the cycle.

From these observations, it is readily apparent that substantial advantage can be achieved by the described procedures. These results were apparently the result of the influence of water in mitigating the inhibiting effect of nitrogen on both hydrogenation and hydrocracking activity. The aromatics hydrogenation activity of the illustrated system increased along with the hydrocracking activity as evidenced by the proportionately higher aromatics conversion rate which is 'known to necessitate hydrogenation.

EXAMPLE 2 The catalyst employed in this example comprised a refractory oxide base consisting of 20 weight-percent stabilized zeolite Y, 65 weight-percent Harshaw alumina and 15 weight-percent peptized catapal alumina. The active metal components were nickel and tungsten present in amounts equivalent to 4.1 weight-percent NiO and 22.8 weight-percent W0 based on the total dry weight of the active metal components and refractory oxide. The resulting composition had an apparent bulk density of 0.957 gram/cc. and a surface area of 248 meters per gram. The catalyst was sulfide by contacting with a 10% mixture of hydrogen sulfide in hydrogen suflicient to convert substantially all of the metal oxides to the corresponding sulfides.

The, feed employed in this operation had the following characteristics.

TABLE 2 Boiling point range, by D-1160, F. 5 -986 Composition, by Univ. High Pass, wt.-percent:

Total saturates 45.1 Total aromatics 27.7 Sulfur, X-ray, wt.-percent 2.91 Nitrogen, total, wt.-percent 0.082 AP I gravity, D-287 22.3

Conversion was effected in la recycle system in which the above-described feed was passed into admixture with a recycle stream and the combination was then introduced into a hydrocracking zone in which the described catalyst was retained in a fixed bed. The product from the reaction zone was admixed with water (during those periods in which water injection was employed) and the combination was flashed to produce a vapor phase comprising primarily hydrogen, water vapor, and hydrocarbons having 3 carbon atoms per molecule or less. This vapor phase was employed as hydrogen recycle to the hydrocracking zone with hydrogen makeup being added as required to maintain the desired hydrogen levels in the reactor. The water concentration in the hydrogen recycle stream Was determined by the saturation level of water in the vapor phase and was equivalent approximately 0.1 weight-percent water based on total hydrocarbon feed to the hydrocracking zone.

The liquid hydrocarbon phase from the separation stage was treated by caustic scrubbing and fractionation to recover an overhead product having an end boiling point of 545 F. ASTM and a recycle product boiling above 545 which was continuously recycled to the hydrocracking zone.

The hydrocracking stage was continuously operated at a pressure of 1990 p.s.i.g., a liquid hourly space velocity of 1.5 and a hydrogen concentration equivalent to 7000 standard cubic feet of hydrogen per barrel of total feed. Reactor temperature was varied as required to maintain a 50% conversion level per pass to products boiling below 545 F. By this means the effect of water in the reaction zone was determined as a function of the temperature required to maintain the established conversion level of 50% per pass.

The run was initiated and continued for a period of 9 days (216 hours) with water injection at a rate equivalent to 0.1 weight-percent based on total reactor charge. Water injection via saturation of the hydrogen recycle stream was then discontinued by omitting water introduction into the reactor product as previously described. Within 3 hours following the discontinuance of water injection the equivalent catalyst activity had increased to a point at which 50% conversion per pass could be maintained at a temperature of 714 After reaching this maximum activity the catalyst being to deactivate and continued deactivation gradually for a period of about 33 hours after which time a temperature of 733 F. was required to maintain the same conversion level. Water injection was commenced after ll days into the run (252 hours) after which time the catalyst activity began to gradually increase and reached an activity after an additional 48 hours sufficient to maintain 50% conversion per pass at a temperature of 722 F. The total system lined out at that temperature (722 F.) indicating a gradual catalyst deactivation during the 4 day test period equivalent to a daily temperature increase requirement (TIR) of 0.714 F. Obviously the gradual catalyst deactivation indicated by the noted temperature increase requirement tended to normalize the effect of water addition and deletion. Consequently, it becomes apparent that the differences in reaction temperature required to maintain 50% conversion would have been even greater than those determined in this operation were it not for the gradual catalyst deactivation observed. Nevertheless, the unusual cycle of catalyst activity resulting from the removal of water from a hydrocracking system was apparent from these observations. The catalyst activity first increased by an amount equivalent to a reaction temperature difference of 5 F. and then gradually decreased by an amount equivalent to about 19 F. Activity then gradually increased to the pre-run activity after reintroduction of 0.1 weight-percent water (taking TIR into account). These results are summarized in Table 3.

TABLE 3 Water level,

percent Run length, hrs.

B Required to maintain 50% conversion per pass to 545 F. minus prodnet; in view of catalyst deactivation equivalent to adaily TIR at 0.7 F.

If desired, reactor temperature can be varied while cycling water addition and deletion to maintain a constant conversion level. The markedly reduced reaction temperatures required to accomplish this result are known to greatly reduce catalyst deactivation rate and consequently extend the run length obtainable before regeneration is necessitated. Conversely, reaction conditions can be maintained constant while cycling water addition as described to vary the overall conversion level. Depending upon the downstream systems necessary to post-treat the hydrocracker eflluent one or the other of these procedures may be preferred. In most situations it is desirable to maintain a constant conversion level in view of the complexities involving in modifying the operation of downstream fractionation and recycle systems. Consequently, in most situations it may be simpler to vary reaction temperature as a function of water addition and conversion level in order to maintain a constant conversion level at all times. In any event, considerable improvement in hydrocarbon conversion can be achieved by either of these procedures or other alternatives readily apparent to one skilled in the art in view of the aforegoing disclosure in the appended claims.

We claim:

1. The method of hydrocracking hydrocarbons containing at least about 10 p.p.m. organically bound nitrogen including the steps of contacting said hydrocarbons at conditions of temperature, pressure and hydrogen partial pressure sufficient to effect substantial hydrocracking of said hydrocarbons with a hydrogenation active catalyst comprising a catalytically active amount of at least one active component selected from the metals, oxides and sulfides of Groups II, IV-B, VI-B, VII-B and VIII and a foraminous refractory oxide in the presence of a controlled amount of water and/or water precursor corresponding to at least about 0.1 weight-percent water based on said hydrocarbon.

2. The method of claim 1 wherein said hydrocarbon contains about 10 to about 1000 p.p.m. nitrogen as organonitrogen compounds and said foraminous refractory oxide is selected from alumina, silica, zirconia, magnesia and combinations thereof.

3. The method of claim 1 wherein said hydrocarbons are hydrocracked in the presence of said catalyst at a temperature of about 400 to about 900 F. in the presence of at least about 500 standard cubic feet of hydrogen per barrel of said hydrocarbon.

4. The method of claim 1 wherein said refractory oxide comprises at least one of alumina, silica-alumina cogels, silica-magnesia and zeolitic aluminosilicates, and said hydrocarbon boils primarily above about 200 F.

5. The method of claim 1 wherein said hydrocarbon boils above about 200 F. and contains about 10 to about 1000 p.p.m. nitrogen as organonitrogen compounds, said foraminous oxide is selected from alumina, silica-alumina cogels, silica-magnesia and zeolitic aluminosilicates, said active component comprises at least one of the metals, oxides and sulfides of calcium, magnesium, molybdenum, tungsten, manganese, rhenium, nickel, cobalt, platinum and palladium and said hydrocarbon is contacted with said catalyst at a temperature of about 400 to about 900 F. with at least about 500 standard cubic feet of hydrogen per barrel of said hydrocarbon in the presence of about 0.1 to about 5 weight-percent water.

6. The method of claim 5 wherein said hydrocarbon is contacted with said catalyst at a total pressure of at least about 1000 =p.s.i.'g. and a liquid hourly space velocity of 0.2 to about 10.

7. The method of claim 1 wherein said foraminous oxide comprises at least one of alumina, silica-alumina and zeolitic aluminosilicates, said active component comprises at least one of the metals, oxides and sulfides of molybdenum, tungsten, nickel, cobalt, platinum and palladium, said hydrocarbon boils primarily above about 200 F. and contains about 10 to about 600 p.p.m. nitrogen as organonitrogen compounds and is hydrocracked in the presence of said catalyst at a temperature of about 400 to about 900 F. and a liquid hourly space velocity of about 0.2 to about 10 with at least about 500 standard cubic feet of hydrogen per barrel of said hydrocarbon in the presence of about 0.1 to about 2 weight-percent water.

8. The method of claim 1 wherein said foraminous oxide is selected from alumina, silica stabilized alumina and zeolitic aluminosilicates and said catalyst further comprises at least about 0.2 weight-percent phosphorus.

9. The method of claim 1 wherein said catalyst comprises at least one of alumina, silica stabilized alumina and aluminosilicates containing a promoting amount of at least one of nickel and cobalt metals, oxides and sulfides and a promoting amount of at least one of molybdenum and tungsten metals, oxides and sulfides 15 and at least about 0.2 weight-percent phosphorus and said hydrocarbon is contacted with said catalyst at a temperature of about 500 to about 900 F. in the presence of. at least about 500 standard cubic feet of hydrogen per barrel of said hydrocarbon sufiicient to hydrocrack at least a substantial proportion of said hydrocarbon.

10. The method of hydrocrac'king hydrocarbons boiling between about 200 and about 1200 F. containing about to about 1000 p.p.m. nitrogen as organonitrogen compounds which comprises contacting said hydrocarbons with a hydrogenation active catalyst comprising a promoting amount of at least one of nickel, cobalt, molybdenum and tungsten metals, oxides and sulfides distended on a refractory oxide support comprising at least one of alumina, silica, magnesia and aluminosilicates in the presence of a controlled amount of Water and/or water precursor corresponding to about 0.2 to about 2 weight-percent water at a temperature of about 400 to about 900 F., a pressure of at least about 1000 p.s.i.g. and a liquid hourly space velocity of about 0.2 to about 10 in the presence of hydrogen added at a rate of at least about 1000 standard cubic feet per barrel of said hydrocarbon.

11. The method of claim 10 wherein. said catalyst comprises about 0.5 to about 10 weight-percent of at least one of nickel and cobalt metals, oxides and sulfides determined as the corresponding oxide and at about 4 to about 30 weight-percent of at least one of molybdenum and tungsten metals, oxides and sulfides determined as the corresponding oxide distended on said refractory oxide support comprising at least about 60 percent alumina.

12. The method of claim 11 wherein said catalyst further comprises at least about 0.2 weight-percent phosphorus, said hydrocarbon boils primarily above about 400 and is contacted with said catalyst at a temperature of at least 500 F. and under conditions 16 of pressure and hydrogen concentration suflicient to bydrocrack a substantial proportion of said hydrocarbon.

13. The method of claim 10 wherein said hydrocarbon is contacted with said catalyst over a protracted run length and said water is provided to the system for a period of at least about 10 minutes in the form of at least one of free Water and Water precursors convertible to water under the conditions at which said hydrocarbon is contacted with said catalyst.

14. The method of claim 10 in which said water is provided to said system for a period of at least about 1 to about hours in the form of at least one of free water and alcohols convertible to water and the conditions under which said hydrocarbon is contacted with said catalyst.

References Cited UNITED STATES PATENTS 3,546,100 12/1970 Yan 208111 3,023,159 2/ 1962 Ciapetta et al. 208- 3,037,930 6/1962 Mason 208111 XR 3,058,906 10/1962 Stine et al. 208-111 3,157,590 11/1964 Scott, Jr., et al. 208-111 3,197,397- 7/1965 Wight 208111 3,238,120 3/ 1966 Sale 208-111 3,242,067 it 3/ 1966 Arey, -Jr. et al. 208-111 3,278,417 10/ 1966 Van Driesen 208--109 XR 3,531,396 9/ 1970 Messing et al. 208-111 3,173,853 3/1965 '-P'eralta 208.112 XR 3,501,396 3/1970 Gatsis 208112 XR TOBIAS E. LEVOW, Primary Examiner P. F. SHAVER, Assistant Examiner US. Cl. X.R.

Referenced by
Citing PatentFiling datePublication dateApplicantTitle
US3869522 *Jun 18, 1973Mar 4, 1975Shell Oil CoAromatics hydrogenation process
US4025428 *May 7, 1975May 24, 1977Ciba-Geigy CorporationProcess for the purification of industrial effluents
US4097364 *Mar 24, 1976Jun 27, 1978Chevron Research CompanyHydrocracking in the presence of water and a low hydrogen partial pressure
US4097376 *Feb 24, 1977Jun 27, 1978Ciba-Geigy CorporationProcess for the purification of industrial effluents
US4104160 *Feb 25, 1977Aug 1, 1978Ciba-Geigy CorporationCellulosic absorbent pretreated with polyalkyleneimine and metal salt
US4149960 *Mar 23, 1978Apr 17, 1979Mobil Oil CorporationDewaxing by contacting with an aluminum silicate zeolite catalyst
US4401556 *Apr 12, 1982Aug 30, 1983Union Carbide CorporationUsing a hydrogenation component-modified zeolite; catalyst selectivity; hydrocarbon fuels
US4585751 *Jun 24, 1985Apr 29, 1986Phillips Petroleum CompanyHydrotreating catalysts
US4649127 *Jan 9, 1985Mar 10, 1987Mobil Oil CorporationProcess for hydrogenative reactivation of spent catalysts
US4670134 *May 2, 1986Jun 2, 1987Phillips Petroleum CompanyCatalytic hydrofining of oil
US4790927 *Jan 9, 1984Dec 13, 1988Union Oil Company Of CaliforniaProcess for simultaneous hydrotreating and hydrodewaxing of hydrocarbons
US4877762 *Aug 22, 1988Oct 31, 1989Union Oil Company Of CaliforniaCatalyst for simultaneous hydrotreating and hydrodewaxing of hydrocarbons
US5629451 *Jun 30, 1994May 13, 1997Chemical Research & Licensing CompanyReduction of nitrile contaminants by selective hydrogenation
US5741930 *Apr 28, 1997Apr 21, 1998Chemical Research & Licensing CompanyReduction of nitrite contaminants by selective hydrogenation
US5779883 *Mar 17, 1997Jul 14, 1998Catalytic Distillation TechnologiesCatalytic hydrogenation of petroleum streams
US5837130 *Oct 22, 1996Nov 17, 1998Catalytic Distillation TechnologiesCatalytic hydrotreatment of crude petroleum for vacuum gas oils, catalytic hydrodesulfurization of vacuum gas oils
US5961815 *Sep 23, 1996Oct 5, 1999Catalytic Distillation TechnologiesHydroconversion process
US6416659Aug 17, 2000Jul 9, 2002Catalytic Distillation TechnologiesHydrodesulfurization of diesel stream in a distillation column reactor; liquid diesel oil is fed above the catalyst bed and the hydrogen is fed below
US7323100Jul 13, 2005Jan 29, 2008Conocophillips Companyamorphous catalyst support suitable for making a hydrocracking catalyst; amorphous mixed inorganic oxide, such as silica-alumina, silica-titania, silica-magnesia, or alumina-titania
Classifications
U.S. Classification208/111.15, 423/564, 208/108, 208/111.35, 585/752, 208/112, 585/268, 208/109, 208/110, 208/114, 208/111.3, 585/269, 585/489, 585/270, 585/841, 208/111.25, 585/266, 585/275
International ClassificationC10G47/00
Cooperative ClassificationC10G47/00
European ClassificationC10G47/00