PROCESS FOR CONTINUOUS DEWAXING OF HYDROCARBON OILS

Abstract: A process for the shape-selective catalytic hydrodewaxing of a hydrocarbon oil, which comprises introducing the hydrocarbon oil, shape-selective dewaxing catalyst and hydrogen at an upper portion of a moving bed reactor, removing dewaxed hydrocarbon oil, partially spent catalyst, and unconsumed hydrogen at a lower portion of the moving bed reactor, separating the dewaxed hydrocarbon oil and the unconsumed hydrogen, reactivating the partially spent catalyst and returning the reactivated catalyst to an upper portion of the moving bed reactor at a rate such that the temperature within the reactor remains substantially constant.

Description: This invention relates to a process for dewaxing hydro carbon oils. Processes for dewaxing petroleum distillates have been known for a long time. Dewaxing is required when highly paraffinic oils are to be used in products which need to remain mobile at low temperatures, e.g., lubricating oils, heating oils, jet fuels. The higher molecular weight straight chain normal and slightly branched paraffins which are present in oils of this kind are waxes which are the cause of high pour points in the oils and, if adequately low pour points are to be obtained, these waxes must be wholly or partly removed. In the past, various solvent removal techniques were used, e.g., propane dewaxing, methyl ethyl ketone dewaxing, but the decrease in demand for petroleum waxes as such, together with the increased demand for gasoline and distillate fuels, has made it desirable to find processes which not only remove the waxy components, but which also convert these components into other materials of higher value. Catalytic dewaxing processes achieve this end by selectively cracking the longer chain n-paraffins to produce lower molecular weight products which may be removed by distillation. Processes of this kind are described, for example, in The Oil and Gas Journal , Jan. 6, 1975, pages 69-73, and U. S. Patent No. 3,668,113. In order to obtain the desired selectivity, the catalyst has usually been a shape-selective zeolite having a pore size which admits the straight chain n-paraffins either alone or with only slightly branched chain paraffins, but which excludes more highly branched materials, cycloaliphatics and aromatics. Zeolites such as ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35 and ZSM-38 have been proposed for this purpose in dewaxing processes, and their use in described in U. S. Patent Nos. 3,894,938; 4,176,050; 4,181,598; 4,222,855; 4,229,282 and 4,247,388. When dewaxing is effected, as is customary, in a fixed bed reactor, the zeolite catalyst utilized ages, loses activity and/or becomes deactivated with time. The rate of aging or deactivation depends, to a great exent, on the nature of the feedstock being dewaxed. Loss of activity is evidenced by the fact that it is necessary to raise the temperature within the reactor to achieve the degree of reaction desired. After a time, the temperature is raised to a point where undesirable side reactions, which detrimentally affect the nature of the product and/ or catalyst efficiency, increase to a value which makes it impractical to continue. Dewaxing is then discontinued and the catalyst is reactivated by treatment with hydrogen at an elevated temperature or regenerated by burning with oxygen in an oxygen-containing gas. An obvious disadvantage of such reactivation/ regeneration is that the dewaxing process cannot be carried out while the catalyst is undergoing reactivation/regeneration. This invention provides a continuous process for the shape-selective catalytic hydrodewaxing of a hydrocarbon oil which comprises introducing the hydrocarbon oil, shape-selective dewaxing catalyst and hydrogen at an upper portion of a moving bed reactor, removing dewaxed hydrocarbon oil, partially spent catalyst, and unconsumed hydrogen at a lower portion of the moving bed reactor, separating the dewaxed hydrocarbon oil and the unconsumed hydrogen, reactivating the partially spent catalyst and returning the reactivated catalyst to an upper portion of the moving bed reactor at a rate such that the temperature within the reactor remains substantially constant. In specific embodiments, reactivation of the spent catalyst is with hydrogen at an elevated temperature or by burning with oxygen in an oxygen-containing gas. In a further embodiment, both forms of reactivation may be utilized sequentially or in parallel, taking care that hydrogen-containing and oxygen-containing streams are not mixed and that the oxygen-regenerated catalyst is essentially free of oxygen gas before being returned to the dewaxing reactor. In the drawings, Fig. 1 is a flowchart illustrating overall operation of the process of the invention; and Fig. 2 is a flowchart illustrating an oxygen regeneration system for use in conjunction with the process illustrated in Fig. 1. Feedstock The present process may be used to dewax a variety of hydrocarbon oil feedstocks ranging from relatively light distillate fractions up to high boiling stocks, such as whole crude petroleum, reduced crudes, vacuum tower residua, fluid catalyst cracking (FCC) tower bottoms, gas oils, vacuum gas oils, deasphalted residua and other heavy oils. The feedstock will normally be a C feedstock, since lighter oils will usually be free of significant quantities of waxy components. The process is particularly applicable to waxy distillate stocks, such as gas oils, kerosenes, jet fuels, lubricating oil stocks, heating oils and other distillate fractions, whose pour point and viscosity need to be maintained within certain specification limits. The feedstock will normally contain paraffins, olefins, naphthenes, aromatics and heterocyclic compounds and with a substantial proportion of higher molecular weight n-paraffins and slightly branched paraffins which contribute to the waxy nature of the feedstock. During the dewaxing process, the n-paraffins and slightly branched chain paraffins undergo limited cracking or hydrocracking to form liquid range materials, thereby both reducing the pour point and lowering the viscosity of the feedstock being dewaxed. Some isomerization of straight chain and slightly branched chain paraffins to more highly branched aliphatics may also take place, and contribute to lowering the pour point. Catalyst The preferred catalysts for hydrodewaxing are those shape-selective zeolites having a Constraint Index within the approximate range of 1 to 12. These zeolites retain a degree of crystallinity for long periods, in spite of the presence of steam at high temperature, which induces irreversible collapse of the framework of other zeolites, e.g., of the X and A type. Furthermore, carbonaceous deposits, when formed, may be removed by burning at higher than usual temperatures to restore activity. In many environments, the zeolites of this calss exhibit very low coke-forming capability, conducive to very long times on stream between burning regenerations. An important characteristic of the crystal structure of this class of zeolites is that it provides constrained access to, and egress from, the intracrystalline free space by virtue of having a pore dimension greater than about 5 Angstroms, and pore windows of about a size such as would be provided by 10-membered rings of oxygen atoms. It is to be understood, of course, that these rings are those formed by the regular disposition of the tetrahedra making up the anionic framework of the crystalline aluminosilicate, the oxygen atoms themselves being bonded to the silicon or aluminum atoms at the centers of the tetrahedra. Briefly, the preferred type zeolites useful in this invention possess, in combination, a silica-to-alumina mole ratio of at least about 12, and a structure providing constrained access to the crystalline free space. The zeolite will have a silica/alumina ratio greater than 12. The silica-to-alumina ratio referred to may be determined by conventional analysis. This ratio is meant to represent, as closely as possible, the ratio in the rigid anionic framework of the zeolite crystal and to exclude aluminum in the binder or in cationic or other form within the channels. Such zeolites, after activation, acquire an intracrystalline sorption capacity for normal hexane which is greater than that for water, i.e., they exhibit "hydrophobic" properties. It is believed that this hydrophobic character is advantageous in the present invention. The type of zeolites described freely sorb normal hexane and have a pore dimension greater than about 5 Angstroms. In addition, the structure must provide constrained access to large modules. It is sometimes possible to judge from a known crystal structure whether such constrained access exists. For example, if the only pore windows in a crystal are formed by 8-membered rings of oxygen atoms, then access by molecules of larger cross-section than normal hexane is excluded, and the zeolite is not of the desired type. Windows of 10-membered rings are preferred, although, in some instances excessive puckering or pore blockage may render these zeolites ineffective. Twelve-membered rings do not generally appear to offer sufficient constraint to produce the advantageous conversions, although puckered structures exist, such as TMA offretite, which is a known effective zeolite. Also, structures can be conceived, due to pore blockage or other cause, that may be operative. Rather than attempt to judge from crystal structures whether or not a zeolite possesses the necessary constrained access, a simple determination of the "Constraint Index" may be made by passing continuously a mixture of an equal weight of normal hexane and 3-methylpentane over a sample of zeolite at atmospheric pressure, according to the following procedure. A sample of the zeolite, in the form of pellets or extrudate, is crushed to a particle size about that of coarse sand and mounted in a glass tube. Prior to testing, the zeolite is treated with a stream of air at 1000°F (538°C) for at least 15 minutes. The zeolite is then flushed with helium and the temperature adjusted between 550° to 950°F (288°-510°C) to give an overall conversion of between 10 and 60%. The mixture of hydrocarbons is passed at 1 liquid hourly space velocity (i.e., volume of liquid hydrocarbon per volume of zeolite per hour) over the zeolite with a helium dilution to give a helium to total hydrocarbon mole ratio of 4:1. After 20 minutes on stream, a sample of the effluent is taken and analyzed, most conveniently by gas chromatography, to determine the fraction remaining unchanged for each of the two hydrocarbons. The "Constraint Index" is calculated as follows: The Constraint Index approximates the ratio of the cracking rate constants for the two hydrocarbons. Zeolites suitable for the present invention are those having a Constraint Index in the approximate range of 1 to 12. Constraint Index (CI) values for some typical zeolites are: It is to be realized that the above Constraint Index values typically characterize the specified zeolites, but that such are the cummulative result of several variables used in determination and calculation thereof. Thus, for a given zeolite, depending on the temperatures employed within the aforenoted range of 550° to 950°F (288°-510°C), with accompanying conversion of between 10 and 60%, the Constraint Index may vary within the indicated approximate range of 1 to 12. Likewise, other variables, such as the crystal size of the zeolite, the presence of possible occluded contaminants and binders intimately combined with the zeolite may affect the Constraint Index. It will accordingly be understood by those skilled in the art that the Constraint Index, as utilized herein, while affording a highly useful means for characterizing the zeolites of interest, is approximate, taking into consideration the manner of its determination. While the above experimental procedure will enable one to achieve the desired overall conversion of 10 to 60% for most catalyst samples and represents preferred conditions, it may occasionally be necessary to use somewhat more severe conditions for samples of very low activity, such as those having a very high silica-to-alumina ratio. In those instances, a temperature of up to about 1000°F (538°C) and a liquid hourly space velocity of less than 1, such as 0.1 or less, can be employed in order to achieve a minimum total conversion of about 10%. The class of zeolites defined herein is exemplified by ZSM-5, ZSM-11, ZSM-12, ZSM-35, ZSM-38 and other similar materials. U. S. Patent No. 3,702,886 describes and claims ZSM-5. ZSM-11 is described in U.S. Patent No. 3,709,979. ZSM-12 is described in U. S. Patent No. 3,832,449. ZSM-35 is described in U. S. Patent No. 4,016,245. ZSM-38 is more particularly described in U. S. Patent No. 4,046,859. The specific zeolites described, when prepared in the presence of organic cations, are catalytically inactive, possibly because, the intracrystalline free space is occupied by organic cations from the forming solution. They may be activated by heating in an inert atmosphere at 1000°F (538°C) for 1 hour, for example, followed by base-exchange with ammonium salts followed by calcination at 1000°F (538°C) in air. The presence of organic cations in the forming solution may not be absolutely essential to the formation of this type zeolite; however, the presence of these cations does appear to favor the formation of this special type of zeolite. More generally, it is desirable to activate this type catalyst by base-exchange with ammonium salts followed by calcination in air or about 1000°F (538°C) for from about 15 minutes to about 24 hours. Natural zeolites may sometimes be converted to this type zeolite catalyst by various activation procedures and other treatments, such as base-exchange, steaming, alumina extraction and calcination, in combinations. Natural minerals which may be so treated include ferrierite, brewsterite, stilbite, dachiardite, epistilbite, heulandite, and clinoptilolite. The preferred crystalline aluminosilicates are ZSM-5, ZSM-11, ZSM-12, ZSM-35 and ZSM-38, with ZSM-5 being particularly preferred. In a preferred aspect of this invention, the zeolites hereof are selected as those having a crystal framework density, in the dry hydrogen form, of not substantially below about 1.6 grams per cubic centimeter. It has been found that zeolites which satisfy all three of these criteria are most desired. Therefore, the preferred zeolites of this invention are those having a Constraint Index as defined above of about 1 to about 12, a silica-to-alumina ratio of at least about 12 and a dried crystal density of not less than about 1.6 grams per cubic centimeter. The dry density for known structures may be calculated from the number of silicon plus aluminum atoms per 1000 cubic Angstroms, as given, e.g., on paper 19 of the article on Zeolite Structure by W. M. Meier. This paper is included in "Proceedings of the Conference on Molecular Sieves, London, April 1967", published by the Society of Chemical Industry, London, 1968. When the crystal structure is unknown, the crystal framework density may be determined by classical pykometer techniques. For example, it may be determined by immersing the dry hydrogen form of the zeolite in an organic solvent which is not sorbed by the crystal. It is possible that the unusual sustained activity and stability of this class of zeolites is associated with its high crystal anionic framework density of not less than about 1.6 grams per cubic centimeter. This high density, of course, must be associated with a relatively small amount of free space within the crystal, which might be expected to result in more stable structures. This free space, however, is important as the locus of catalytic activity. Crystal framework densities of some typical zeolites are: When synthesized in the alkali metal form, the zeolite is conveniently converted to the hydrogen form, generally by intermediate formation of the ammonium form as a result of ammonium ion-exchange and calcination of the ammonium form to yield the hydrogen form. In addition to the hydrogen form, other forms of the zeolite, wherein the original alkali metal has been reduced to less than about 1.5 wt %, may be used. Thus, the original alkali metal of the zeolite may be replaced by ion-exchange with other suitable ions of Groups IB to VIII of the Periodic Table, including, by way of example, nickel, copper, zinc, palladium, calcium or rare earth metals. In practicing the desired conversion process, it may be desirable to incorporate the above-described crystalline aluminosilicate zeolite in another material resistant to the temperature and other conditions employed in the process. Such matrix materials include synthetic or naturally occurring substances, as well as inorganic materials such as clays, silica and/or metal oxides. The latter may be either naturally occurring or in the form of gelatinous precipitates or gels, including mixtures of silica and metal oxides. Naturally occurring clays which can be composited with the zeolite include those of the montmorillonite and kaolin families, which families include the sub-bentonites and the kaolins commonly known as dixie, McNamee-Georgia and Florida clays, or others in which the main mineral constituent is halloysite, kaolinite, dickite, nacrite or anauxite. Such clays can be used in the raw state, as originally mined, or initially subjected to calcination, acid treatment or chemical modification. In addition to the foregoing materials, the zeolites employed herein may be composited with a porous matrix material, such as alumina, silica-alumina, silica-magnesia, silica-zirconia, silica-thoria, silica-berylia, silica-titania, as well as ternary compositions, such as silica-alumina-thoria, silica-alumina-zirconia, silica-alumina-magnesia and silica-magnesia, zirconia. The matrix may be in the form of a cogel. The relative proportions of zeolite component and inorganic oxide gel matrix may vary widely, with the zeolite content ranging from between about 1 to about 99 wt %, and more usually in the range of about 5 to about 80 wt % of the composite. The catalysts employed in this invention are constituted by a zeolite, as described above, in intimate combination with a hydrogenating component, such as tungsten, vanadium, molybdenum, rhenium, nickel, cobalt, chromium, manganese, or a noble metal, such as platinum or palladium. Such component can be exchanged into the composition, impregnated therein or physically intimately admixed therewith. Such component can be impregnated in or onto zeolite, such as, for example, by, in the case of platinum, treating the zeolite with a platinum metal-containing ion. Thus, suitable platinum compounds include chloroplatinic acid, platinuous chloride and various compounds containing the platinum amine complex. The compounds of the useful platinum or other metals can be divided into compounds in which the metal is present in the cation of the compound and compounds in which it is present in the anion of the compound. Both types of compounds which contain the metal in the ionic state can be used. A solution in which platinum metals are in the form of a cation or cationic complex, e.g., Pt(NH₃)₄Cl₂, is particularly useful. Process In the embodiments illustrated in Figs. 1 and 2, hydro­ dewaxing is accomplished by contacting a waxy lube base stock with hydrogen and catalyst in a moving-bed reaction zone containing a mixture of reactivated and/or regenerated catalyst. (As used herein, reactivation generally refers to treatment with hydrogen and regeneration to burning with oxygen. However, in practice, those terms are often used interchangeably.) Catalyst is continuously or intermittently removed from the bottom of the hydrodewaxing reactor and passed to a reactivation section, where catalyst is either contacted with a hot hydrogen stream to volatilize and strip entrained and adsorbed hydrocarbons or, alternatively, contacted with an oxygen-containing gas to burn off coke and other accumulations or hydrocarbons. Reactivated/regenerated catalyst is returned to the dewaxing zone. A small segment of the total catalyst inventory can be continuously or periodically removed and replaced by fresh catalyst to maintain overall catalyst activity. Figs. 1 and 2 describe certain configurations for practicing the process of the present invention. The particular embodiments described are chosen to illustrate the invention, and the description is not intended to exclude other modifications and configurations within the spirit of the invention. In order to simplify the description, furnaces, heat exchangers, pumps and other conventional heat integration equipment are not shown. Referring to Fig. 1, liquid feed is fed to the top of reactor 4 through line 1 along with recycle and makeup hydrogen, typically at a pressure in the range 2170 to 8377 kPa (300 to 1200 psi), from lines 2 and 3, respectively. The mixed phase stream passes downward through the catalyst bed and exits the reactor through lines 5 and 6 into the high temperature separator 7. The liquid product is removed through line 8 and the vapor stream, containing primarily hydrogen, light hydrocarbons, ammonia and hydrogen sulfide, is removed through line 9. After cooling (not shown), the stream is further split in the low temperature separator 10 into a liquefied light hydrocarbon stream exiting through line 11, and a hydrogen-rich stream in line 12. The hydrogen stream may be further purified in an amine adsorption system (not shown), and additionally may have water removed by molecular sieve adsorption (not shown) before pressuring back to process conditions in the recycle gas compressor 13. Pressurized hydrogen in line 14 is used as recycle hydrogen in line 2, as preliminary catalyst deoiling gas in lines 15 and 16, for pressurizing lock hopper 35 via line 39 and for transporting catalyst via line 17. Fresh catalyst is intermittently fed to the top of reactor 4 from storage vessel 18 by way of lock hopper 19, where the catalyst charge is brought up to reactor pressure. Here, fresh catalyst mixes with recirculated catalyst transferred from the reactivation/regeneration section through line 20. The mixture of reactivated-regenerated and fresh catalyst from the storage vessel is passed to the reactor at a rate such that the temperature within the reactor remains substantially constant at a predetermined value, generally in the range 288° to 426°C (550 to 800°F), chosen, in part, depending on the nature of the feedstock being dewaxed. The temperature within the reactor is adjusted by adjusting the feed rate of catalyst being supplied to the reactor. Operation with a continuously reactivated and/or regenerated catalyst at a substantially constant temperature results in the continuous production of an essentially uniform dewaxed product having a substantially constant pour point. Catalyst gradually travels down through the hydrodewaxing reactor and is removed through line 21 to a catalyst stripping hopper 22, where some deoiling is effected by stripping with the before mentioned split stream of recycle hydrogen from line 16. Catalyst is then transported through line 23 by way of a star valve 24, and by pneumatic transport through line 25 to lock hopper 40, where it is fed by star value 41 to reactor 26 for reactivation. As is well known, reactivation with hydrogen involves treating or contacting the catalyst to be reactivated with hydrogen at elevated temperatures, most generally in the range of 427° to 538°C (800° to 1000°F). Such treatment results in removal of a portion of the carbonaceous deposits formed on the catalyst during dewaxing, and in striping various volatile materials adsorbed onto the catalyst. As shown in Fig. 1, reactivation is effected continuously, or in a batch operation, in reactor 26 with hydrogen from compressor 27 fed through lines 28 and 29. Hydrogen exits the reactor through line 30 to knock-out pot 31 for removal, via line 32, of any liquids stripped from the catalyst. A separate stream of hydrogen 33 may also be used as transport gas in line 25. Reactivated catalyst leaving reaction 26 via line 34 enters lock hopper 35, where it reduces until it is transported back to reactor 4 through line 36, star valve 37 and lines 38 and 20. Pressure in lock hopper 35 is equalized by hydrogen from line 39. In addition to reactivation with hydrogen, the catalyst is regenerated with oxygen, as required. As is well known, regeneration with oxygen involves burning off carbonaceous and other deposits with oxygen in an oxygen-containing gas at elevated temperatures. Since burning with oxygen is exothermic, regeneration generally initiated at a lower temperature, such as 316°C (600°F), with 1 to 10% oxygen (the oxygen in air is diluted with an inert gas, such as nitrogen), and the temperature and concentration of oxygen are gradually increased until all of the carbonaceous deposits have been oxidized. This generally occurs at a temperature of 483° to 538° (900° to 1000°F), with an oxygen concentration of 1 to 20% (air contains about 20 vol % of oxygen). Care should be taken not to exceed the temperature at which the crystalline structure of the zeolite becomes disrupted or the metal catalyst becomes deactivated. Also, care should be taken to ensure that oxygen-containing and hydrogen-containing gas streams do not mix and oxygen-regenerated catalyst is stripped of adsorbed oxygen prior to return to the dewaxing process. Regeneration with oxygen can be done either continuously or an be alternated with hydrogen reactivation. In preferred embodiments, a separate oxygen regeneration system is installed in parallel or in series with a hydrogen reactivation system, illustrated in Fig. 1. The oxygen regeneration system illustrated in Fig. 2 includes a double valve isolation system 61, 61′, 62 and 62′ upstream and downstream, respectively, to separate the oil/hydrogen atmosphere of the reactor system from the oxygen-containing atmosphere of the regeneration system. In operation, the catalyst to be regenerated by treatment with oxygen is passed through initial isolation valve system 61 and 61′ into lock hopper 63 and purged free of hydrogen (purge system not shown). The purged catalyst is then fed by valve 70 to regenerator 64, where it is contacted with recycle gas stream 65 containing combustion gas mixed with makeup air 66. Combustion gases exit the regenerator via line 71 and are passed to separator-cooler 67, where water separates and is removed. The catalyst passed downflow via valve 72 to catalyst cooler/lock hopper 68, where oxygen-containing gases are purged (purge system not shown). The purged catalyst passes through isolation valves 62 and 62′ to transfer line 69 for return to catalyst storage vessel 18 (shown in Fig. 1). In an alternate embodiment, the reactivation vessel 26 (shown in Fig. 1) is used for both hydrogen reactivation and oxygen regeneration (the system is run in alternating modes for treatment with hydrogen and oxygen). As before, care should be taken to ensure that the oxygen-containing and hydrogen-containing gas streams do not mix and oxygen-regenerated catalyst is stripped of adsorbed oxygen prior to return to the dewaxing process. The dewaxing process of the present invention involves more than the conversion of a batch process to a continuous one. In addition to providing a convenient means for dewaxing at a substantially constant temperature, the present process permits operation at lower temperature (in conventional processes, the temperature is progressively raised to compensate for loss in catalyst activity until it becomes necessary to stop dewaxing and reactivate and/or regenerate the catalyst). In addition to permitting operation at a relatively lower temperature, by also providing a catalyst of relatively uniform activity, there is obtained, starting with a particular feedstock, a more uniform dewaxed product.

Claims: 1. A process for the shape-selective catalytic hydro­ dewaxing of a hydrocarbon oil, which comprises: introducing said hydrocarbon oil, shape-selective dewaxing catalyst and hydrogen at an upper portion of a moving bed reactor; removing dewaxed hydrocarbon oil, partially spent catalyst, and unconsumed hydrogen at a lower portion of the moving bed reactor; separating the dewaxed hydrocarbon oil and the unconsumed hydrogen; reactivating the partially spent catalyst; and returning the reactivated catalyst to an upper portion of the moving bed reactor at a rate such that the temperature within the reactor remains substantially constant. 2. A process according to Claim 1, wherein the hydro carbon oil is a lubricating oil stock distillate. 3. A process according to Claim 1 or 2, wherein the partially spent catalyst is reactivated with hydrogen. 4. A process according to any one of Claims 1-3, wherein the par tially spent catalyst is reactivated by burning with oxygen in an oxygen-containing gas. 5. A process according to any one of Claims 1-4, wherein the spent catalyst is reactivated first with hydrogen and then by burning with oxygen in an oxygen-containing gas. 6. A process according to any one of Claims 1-5, wherein the dewaxed hydrocarbon oil has a substantially constant pour point. 7. A process according to any one of Claims 1-6, wherein the unconsumed hydrogen is recovered and recycled to the process. 8. A process according to any one of Claims 1-7, wherein the shape-selective dewaxing catalyst has a Constraint Index in the range of about 1 to 12. 9. A process according to Claim 8, wherein said catalyst has a silica-to-alumina ratio of at least about 12.

Inventors: KENNEDY, CLINTON ROBERT, TABAK, SAMUEL ALLEN
Applicants: MOBIL OIL CORPORATION
Application year: 1987
Main classification: C10G45
IPC codes: B01J38, C10G45, B01J29, C10G73