|Publication number||US3254017 A|
|Publication date||May 31, 1966|
|Filing date||Aug 23, 1963|
|Priority date||Aug 23, 1963|
|Publication number||US 3254017 A, US 3254017A, US-A-3254017, US3254017 A, US3254017A|
|Inventors||Arey Jr William Floyd, Burgess Mason Ralph, Charles Paule Robert|
|Original Assignee||Exxon Research Engineering Co|
|Export Citation||BiBTeX, EndNote, RefMan|
|Patent Citations (4), Referenced by (39), Classifications (13)|
|External Links: USPTO, USPTO Assignment, Espacenet|
y 1966 w. F. AREY. JR. ETAL 3,254,017
PROCESS FOR HYDROCRACKING HEAVY OILS IN TWO STAGES H2 '2 '3 Filed Aug. 25, 1963 n j vw 9 l4 FEED Y HYDROCRACKING I5 ZONE HYDROCRACKING ZONE 28 22 J DRY GAS 29 NAPHTHA 4 k FRACTIONATOR William Floyd Arey, Jr. Ralph Burgess Mason NVENTORS Robert Charles Poule BY 2 Z PATENT AGENT United States Patent 3,254,017 PROCESS FOR HYDROCRACKING HEAVY OILS IN TWO STAGES William Floyd Arey, Jr., Baton Rouge, Ralph Burgess Mason, Denham Springs, and Robert Charles Paule, Baton Rouge, La., assignors to Esso Research and Engineering Company, a corporation of Delaware Filed Aug. 23, 1963, Ser. No. 304,162 20 Claims. (Cl. 20859) This invention relates to the catalytic hydrocracking of heavy hydrocarbon oils. Particularly, it relates to a two-stage catalytic hydrocracking process for heavy hydrocarbon oils containing appreciable quantities of sulfur, nitrogen, and metal-containing compounds, which process comprises partial conversion of high boiling residual petroleum fractions, metals removal, denitrogenation, and desulfurization in the first stage utilizing a large pore catalyst; and selective hydrocracking to valuable lower boiling products in the second stage utilizing a fine pore catalyst, with the conversion of residual petroleum fractions in the second stage being strictly limited by control of operating conditions and the pore size of the catalyst. More particularly, it relates to said hydrocracking process wherein the fine pore catalyst utilized in the second stage comprises a crystalline alumino-silicate zeolite composited with a platinum group metal.
Hydrocracking has recently become a subject of considerable interest in the petroleum industry because of certain particularized advantages that it offers over conventional catalytic cracking operations. Chemically, hydrocracking may be thought of as a combination of hydrogenation and catalytic cracking, and is effected in the presence of a suitable bifuncti-onal catalyst capable of simultaneously cracking high boiling hydrocarbons to lower boiling fractions and hydrogenating olefinic and aromatic materials into saturated parafiins and naphthenes. An advantage of hydrocracking is its ability to selectively convert a wide variety of feed stocks, including refractive heavy aromatics, to lower boiling distillates with significantly less gas and coke yield and higher yields of good quality liquid products than are usually produced by catalytic cracking.
Generally, hydrocracking finds its highest utility in cracking hydrocarbons boiling in the heavy naphtha and gas oil range. It has however met with only limited acceptance in the upgrading of heavy hydrocarbon oils, particularly low grade heavy hydrocarbon oils such as petroleum residues, shale oil, sour crudes, bitumen deposits, etc., due to the contamination and deactivation of the hydrocracking catalyst by various sulfur, nitrogen and metal-containing organic compounds present in such oils, which materials, in addition to acting as catalyst poisons, tend to deposit coke during the hydrocracking operation.
Throughout the specification and the claims the following terms will be used as herein defined. The term gas oil is meant to include any hydrocarbon fraction derived from crude petroleum, shale oil and the like, or from a hydrocarbon conversion process such as cracking, coking, etc., which fraction boils substantially continuously within the range of from about 430 F. toabout 1000 F. This range will, of course, include the light, medium and heavy gas oils as those terms are commonly employed in the art.
The term residual fraction is meant to include those fractions which either have initial boiling points or remain undistilled above about 1000 F. Typical of such fractions are the heavy bottoms, such as tars, asphalts, etc.
The term naphtha is meant to include those fractions boiling below about 430 F. with the heavy naphtha fraction boiling between about 200 F. and 430 F. and the light naphtha fraction boiling between about 55 F. and 200 F. The term gasoline will be used interchangeably with the term C /43() F. and represents the fraction boiling between about 55 F. and 430 F.
Finally, the term dry gas refers to methane, ethane and propane produced in the hydrocracking process expressed as weight percent of initial charge.
Recently, a markedly superior hydrocracking catalyst has been used extensively for the hydrocracking of a variety of feed stocks. This catalyst comprises a platinum group metal, e.g. palladium, which is deposited on, composited with, or incorporated within a crystalline alumino-silicate zeolite. Prior to the discovery of this hydrocracking catalyst, various conventional catalysts, such as the noble metals, or oxides or sulfides of iron group metals, supported on amorphous material such as silica, silica-alumina, etc., had been found to be extremely sensitive to the presence of feed impurities, and especially to organic nitrogen compounds. Under such conditions these conventional catalysts exhibited low activity and additionally required either frequent regenerations or the use of nonoptimum reaction conditions which were not capable of producing maximum yield and quality of desired product. With the advent of the aforementioned platinum group metal on crystalline zeolite hydrocracking catalyst, the various difii-culties experienced with the prior art catalysts were found to be substantially lessened in the hydrocracking of conventional feeds, e.g. feeds having nitrogen contents up to about 500 p.p.m. such as virgin, coker and catalytic heating oils. Excellent yields of lower boiling hydrocarbons, e.g. gasoline, were obtained with minimum coke-forming tendency, the required frequency of regenerations was substantially reduced, and relatively mild reaction conditions were uti-' lizable without the extensive feed pretreatment (e.g. hydrotreating for the removal of nitrogen) usually required prior to contact with the conventional catalysts. It was additionally found that this new .type of catalyst exhibited a considerably higher degree of selectivity to desired products, e.g. gasoline, than did the conventional catalysts previously employed.
However, although the development of this new type of catalyst has provided a marked improvement in the hydrocracking of conventional'feed stocks, the hydrocracking of heavy petroleum feeds has still proved not entirely satisfactory. It is well known that such feeds as total crudes, and atmospheric and vacuum residua contain soluble ash or metal-containing contaminants. Residual petroleum fractions are'known to contain such metals as vanadium, nickel, iron, copper, sodium and calcium, which are believed to occur as organo-metallic compounds. (The exact nature and composition of these compounds is not fully known, but the presence of vanadium porphyrins in several crude oils and residua has been demonstrated.) It is also known that these metalcontaining contaminants tend to accumulate in the petroleum (and shale oil) fractions boiling above about 1000 F. (1000 F.+). The concentration of these materials is thus greater in atmospheric and vacuum residua than in the corresponding crude oils. It is further known .that conversion of residua to lighter boiling fractions by such means as coking, visbreaking and the like will remove a portion of the metals from the oil but that the high boiling products from such processes will still contain appreciable quantities of metal-bearing compounds. However, in the hydrocracking of such feeds, deposition of metals on the crystalline zeolite catalyst occurs and results in a reduction of the activity and selectivity of the catalyst for the desired conversion. In addition, such metal-bearing compounds frequently increase the deposition of coke on the catalyst with a resulting rapid loss of catalyst activity. While the coke can be periodically removed by regeneration, e.g. burning with air, the presence of these impurities on the catalyst usually causes a pronounced loss of effective catalyst surface area when the catalyst is exposed to the high temperature, oxidizing conditions employed in regeneration. These contaminants, therefore, are not only detrimental to catalytic performance, but also interfere with the regeneration of the catalyst.
It has now been found, that with certain of the crystalline zeolite catalysts containing platinum group metal, the 'hydrocra'cking reaction can be controlled with regard to the type of hydrocarbon molecule which will be hydrocracked. For example, by control of the hydrocracking conditions and the pore size of the crystalline zeolite catalyst, good conversion of hydrocarbon fractions boiling below about 1000 F. can be obtained with little hydrocracking of the high boiling residual fractions of the oil feed. It will be observed, therefore, that due to the accumulation of metal-bearing compounds in the 1000 F.+ residual fractions, ash and coke deposition on the crystallinecatalyst can be substantially avoided by limiting the amount of conversion of those fractions.
The present invention is concerned with an improved process for hydrocracking heavy hydrocarbon oil feeds, which process involves a two-stage conversion employing a different type of catalyst in each stage. The first stage of the present process utilizes a large pore amorphous gel type catalyst which is capable of converting the residual fractions in the feed to gas oil. Simultaneously with this conversion the feed is substantially demetallized, desulfurized and partially denitrogenated. It
has been found that the large pore catalysts contemplated for the first stage of the present process readily remove the soluble metal contaminants contained in heavy oil feeds by deposition of ash within the catalyst pores. In this manner the feed is substantially freed of contaminating materials and at the same time partially hydrocracked to a more readily convertible form prior to contact with the second stage catalyst. The first stage catalyst will naturally become progressively deactivated and may be periodically reactivated by conventional means 'such as hydrotreating, stripping, or burning of the coke deposits followed by hydrotreating.
The second stage of the present process utilizes a catalyst comprising a crystalline alumino-silicate zeolite having a platinum group metal incorporated therein, and having uniform pore diameters of about 5 to about 20, preferably about 7 to about 13, angstrom units. This type of catalyst, although normally highly effectivein the hydrocracking of conventional feed stocks, is of only limited effectiveness, at equivalent reaction conditions, with the heavy oil feeds herein contemplated. As previously mentioned, such feeds tend to deposit ash and coke on the catalyst. Furthermore, at the hydrocracking conditions normally employed with conventional feed stocks, most of the large hydrocarbon molecules contained in the residual fractions are unable to gain 'entry into the framework of the crystalline zeolite structure and little hydrocracking of such fractions occurs. Conversion of such fractions is possible at more severe operating conditions, e.g. at higher temperatures, possibly due to thermal cnacking effects or expansion of the pores ,of the catalyst to allow entry of the large hydrocarbon molecules. However, this is undesirable since ash and coke deposition usually increases, thereby severely contaminating the catalyst. The second stage of the present process is, therefore, operated to take advantage of the relative inability of certain of the crystalline zeolite catalysts (e.g. those having pore diameters of about 5 to about 20 A.) to effectively 'hydrocrack residual fractions,'in order to avoid ash deposition in the second stage. This is accomplished by utilizing a crystalline zeolite catalyst of limited pore size which is substantially incapable of imbibing the large hydrocarbon. molecules contained in residual fractions, and by limiting the amount of conversion of residual fractions in the second stage by control of the operating conditions.
By means of the present process, the metal contaminants are substantially removed in the first stage by deposition on the relatively inexpensive amorphous gel catalyst, and the residual fractions in the feed are partially hydrocracked to gas oil in the first stage thereby forming smaller molecules which may subsequently gain entry into the pores of the second stage zeolite catalyst and be converted therein to the valuable naphtha fractions. Also, ash contamination of the first stage catalyst does not substantially affect the yield and quality of the ultimate product naphtha, since only a small amount of naphtha is producedin this stage.
Although the first stage catalyst of the present process substantially demetallizes the heavy oil feed, there will still be a significant amount of soluble metal compounds contained in the residual fractions of the first stage efiluent. therefore, is that the second stage conversion (to gas oil) of unconverted residual petroleum fractions present in the first stage effiuent be strictly limited to less than about 30 vol. percent, preferably less than about 10 vol. percent, (based on first stage efiluent), in order to prevent substantial ash and metal deposition on the second stage crystalline zeolite catalyst. Control of the degree of conversion of the residual fractions in the second stage is accomplished by strict control of the second stage operating conditions as will be hereinbefore described.
In summary, the process of the present invention enables the highly effective crystalline zeolite catalysts to be successfully utilized in the conversion of heavy oil feeds by combining the abilities of two separate types of catalysts to accomplish the desired result; i.e., the ability of the amorphous gel type catalyst to remove metals and to partially convert residual petroleum fractions to gas oil, and the ability of the crystalline zeolite type catalyst to selectively hydrocrack gas oil to gasoline. By means of the partial conversion and metals removal by the large pore first stage catalyst, the crystalline zeolite catalyst is successfully and effectively utilized for final hydrocracking to valuable lower boiling products, e.g., gasoline, with limited conversion of residual fractions to avoid ash disposition and consequent deactivation. Furthermore, as an incidental advantage, most of the ash formed during the process is deposited on the relatively inexpensive amorphous catalyst in the first stage, rather than the relatively expensive zeolite catalyst in the second stage.
The heavy hydrocarbon oil feeds contemplated for use in the process of the present invention will include crude oils, petroleum residues, shale oil, coal tar, heavy coker distillates, bitumen, etc. More specifically, the feeds comprise high boiling carbonaceous oils containing soluble ash and/or organo-metallic compounds. These oils will usually contain relatively high proportions of organic sulfur and nitrogen compounds as well as metal compounds. The nitrogen and metals content of the heavy oils have been found to be convenient identifying characteristics; the metal impurities being associated with the residual fractions in the feed and the nitrogen impurities being generally the most difiicult to remove. Concurrently with the destruction of the nitrogen impurities, the other reactive species, e.g., asphaltenes, sulfur compounds, etc., are usuallydecomposed to substantially non-contaminating materials. Accordingly, the heavy hydrocarbon oil feeds herein contemplated will contain more than about 5 ppm. ash or metal contaminants, and usually about 5 to 300 p.p.m. metals; and more than about ppm. of nitrogen up to about 20,000 ppm. nitrogen. Most preferably, the feeds contemplated will contain at least a proportion, e.g., about 2 to 100 vol. percent of a residual petroleum fraction boiling above about 1000 F. wherein said fraction contains at least about 500 ppm. nitrogen, and from about 5 to about 1000 ppm. metals.
A critical requirement of the present process,
These heavy oils may also be characterized by their specific gravities, sulfur contents, Conradson carbon residues, etc. However, because of the wide variance of these properties in the heavy oils contemplated, it will be preferred to utilize the aforementioned description. By way of illustration, typical heavy oils may have specific gravities of about 1 to 40 API, sulfur contents of about 0.1 to 5 wt. percent Conradson carbon residues of about 0.5 to 30 wt. percent.
The first hydrocracking stage of the present process is elfected in the presence of hydrogen and a large pore catalyst comprising a suitable hydrogenation metal sup.- ported on an amorphous gel. Such catalysts are well known in the art as hydrofining or hydrogenation catalysts. In the present process, however, the operating conditions in the first stage are more severe than those used in conventional hydrofining operations, since a substantial conversion of residual fractions to gas oil is desired. Suitable catalysts of this type include the metals of GroupsVI and VIII of the Periodic Table, and their oxides or sulfides, either alone or in admixture with each other, composited with an amorphous metal oxide support wherein the metal oxide is selected from the group consisting of silica, oxides of metals in Groups II-A, III-A, and IV-B of the Periodic Table, and mixtures thereof. Examples of the Group VI metals are tungsten and chromium, with the preferred Group VI metal being molybdenum in the form of a molybdate; examples of the Group VIII metal components are cobalt and nickel; and examples of the metal oxides in the amorphous support are alumina, silica, zirconia, magnesia, titania, ceria, thoria, etc. In particular, such catalysts are typified by nickel sulfide-tungsten sulfide, molybdenum sulfide or oxide, combinations of metal sulfides or oxides such as ferric oxide, molybdenum oxide or sulfide and cobalt oxide, all of which are supported on the above amorphous metal oxide supports. A particularly preferred catalyst is cobalt molybdate on alumina. The preferred composite catalysts will contain from about 1 to 10, preferably 2 to 4 wt. percent of a Group VIII metal oxide, preferably cobalt oxide; and from about 5 to 30, preferably to wt. percent of a Group VI metal oxide, preferably molybdenum; supported on the amorphous metal oxide support, preferably alumina or silica-alumina.
The pore size of the above-described first stage catalyst is to be regarded as critical. Suitable catalysts should have average pore diameters of at least about Angstrom units, preferably about 50 to about 300, mose preferably about 80 to about 150 angstrom units, in order that sufficient conversion of residual fractions and metal removal be attained.
The hydrocracking operating conditions employed with the above-described large pore amorphous catalyst in the first stage of the present process include atemperature of from about 600 F. to about 850 F., preferably 750 to 825 F.; a pressure of from about 500 to about 3000 p.s.i.g., preferably 1500 to 2500 p.s.-i.g.; a liquid hourly space velocity of from about 0.1 to about 10.0, preferably 0.5 to 2.0, volumes of feed per volume of catalyst per hour; and a hydrogen rate of from about 3000 to about 20,000, preferably 5000 to 15,000 s.c.f. per barrel of feed. Such conditions will usually be sufficient to ob tain the desired degree of conversion of residual petroleum fractions to gas oil. Typical conversion levels, for heavy oil feeds will generally be within the range of about 20 to about 80, expressed as volume percent conversion of residual fractions to gas oil.
The first-stage hydrocracked efiluent is then passed with or without separation of ammonia and hydrogen sulfide gas to the second hydrocracking stage containing the fine pore crystalline zeolite catalyst. If preferred, the firststage effluent may be scrubbed clean of such gases, or sent directly to a fractionator with the fractionator bottoms being recycled to the second-stage hydrocracker. Alternatively, and frequently preferred from an economic standpoint, the first-stage eflluent may be sent in toto. to
the second-stage reactor, Without the added expense of gas separation, which expense should be weighed against the catalyst life attainable.
The second hydrocracking stage of the present process is effected in the presence of hydrogen and a crystalline alumino-silicate zeolite composited with a platinum group metal. The hydrocracking operating conditions em.- .ployed in the second stage include a temperature of from about 500 to about 800 F., preferably 725 to 775 F.; a pressure of from about 500 to about 3000 -p.s.i.g., preferably 1500 to 2500 p.s;i.g.; a liquid hourly space velocity of from about 0.1 to about 10.0, preferably 0.5 to 2.0 volumes of feed per volume of catalyst per hour; and a hydrogen rate of :fnom about 3000 to about-20,000, preferably 8000 to 15,000 s. c.f. per barrel of feed. Such conditions will generally be sufiicient to obtain the desired degree of conversion of gas oil to naphtha, e.g. 20 to 80 volume percent. As hereinbefore mentioned, a critical requirement is that the conversion of residual petroleum fractions be 'limited in the second stage to avoid ash deposition on the zeolite catalyst. The conversion of residual fractions to lower boiling matenial should be less. than about volume percent (based on first-stage effiuent), e.g. 0 to less than about 30 vol. percent, and preferably less than about 10 volume percent,
erg. 0 to less than about 10 volume percent. This conversion limitation will, in turn, place a maximum limit on the temperature in the second stage, which may have to be reduced should the above limiits be exceeded. Thus, while the first and second stage temperatures and pressunes, as set forth above, may be approximately equal, it will usually be preferable to operate the second stage at a somewhat lower temperature than the first stage, since the conversion 01f residual fractions in the second stage will usually be less than that of the first stage. Furthermore, the crystalline zeolite catalyst used in the second stage will usually have a higher activity for a longer period of time than the first-stage catalyst. Because of the more rapid deactivation of the first-stage catalyst, it may be necessary to raise the temperature in the first stage in order to compensate for loss of activity. Such adjustments of operating conditions will be readily determinable by those skilled in the art without departing from the spirit of the present invention.
The hydrocracking catalyst utilized in the second stage of the present process comprises a crystalline metallo alumino-sili cate zeolite, well known in the art as a molecular sieve, having a platinum group metal (e.g. palladium) deposited thereon -or composited therewith. These crystalline zeolites are characterized by their highly ordered crystalline structure and uniformly dimensioned pores, and have an alumino-silicate anionic cage structure wherein alumina and silica tetra'hedra are intimately connected to each other so as to provide a large number of active sites, with the uniform pore openings facilitating entry of certain molecular structures. For use in the present invention these crystalline alumino-silicate zeolites should have effective pore diameters of about 5 to about 20, preferably about 6 to about 15, and more preferably about 7 to about 13 angstrom units. The size of the pore openings is regarded as critical since smaller openings, e.g. 4 angstrom units will be too small to allow entry of certain feed molecules in the high boiling fractions: and larger pores, i.e. greater than about 20 angstrom units will allow unimpeded entry of the metal-containing components in the feed such as nickel and vanadium porphyrins. Additionally, the crystalline nature of the catalyst is important, since the crystalline structure determines the unifonmity of the pore openings.
Naturally-occurring large pore crystalline aluminosilicate zeolites may be exemplified by the mineral faujasite, which may be beneficially employed. Synthetically produced alumino-silicate zeolites having the required pore diameters such as synthetic faujasite and mordenite 7 are also available and willbe preferred in the present invention. In general, all crystalline alumino-silicate zeolites, in natural or synthetic form contain a substantial portion, e.g. above about 10 wt. percent, of an alkali metal oxide, normally sodium oxide. 7
More specifically, the major component of the catalyst used in the second stage of the present process is a crystalline alumino-sileicate zeolite having an effective pore diameter within the aforementioned ranges, wherein a substantial portion of the alkali metal, e.g. sodium, has been replaced with a cation (either a metal cation or a hydrogen-containing cation, e.g. NH so as to reduce the alkali metal oxide (e.g. Na O) content to less than 10 wt. percent, and preferably to about 1 to wt. percent (based on zeolite). The anhydrous form of the baseexchanged crystalline alumino-silicate zeolite prior to compositing with platinum group metal may be generally expressed in terms of moles by the formula:
wherein Me llS selected from the group consisting of hydrogen and metal cations (so that the alkali metal oxide content is less than wt. percent of the zeolite), n is its valence and X is a number from 2.5 to 14, preferably 3 to 10 and most preferably 4 to 6. having these silica to alumina ratios have been found to .be highly active, selective and stable.
The above-described base exchanged alumino-silicate zeolites serve as the supports for the platinum group metal. By platinum group metal it -is meant to include members of Group VIII of the Periodic Table; e.g. platinum, palladium, rhodium, cobalt, nickel, etc. The final catalyst should contain at least about 0.1 wt. percent platinum group metal, preferably palladium, based on the weight of the final catalyst. Preferred ranges will be about 0.5 to 10 wt. percent, more preferably about 1 to 5 wt. percent, and most preferably about 1.5 to 3 wt. percent.
The processes for synthetically producing the crystalline alumino-silicate zeolite component of the hydrocracking catalyst herein contemplated are well known in the art. They typically involve crystallization from reaction mixtures containing: A1 0 as sodium aluminate, alumina sol and the like; SiO as sodium silicate and/or silica gel and/or silica sol; and metal oxide as alkaline hydroxide, preferably sodium hydroxide, either free or in combination with the above components. Careful control is kept over the metal oxide (e.g. Na O) concentration of the mixture, the proportions of silica to alumina and metal oxide (soda) to silica, the crystallization period, etc., to obtain the desired product. A conventional procedure for producing crystalline alumino-silicate zeolite having a silica to alumina mole ratio of about 4 to 6 is as follows:
Colloidal silica is mixed with a solution of sodium hydroxide and sodium aluminate at ambient temperatures to produce a reaction mixture having the following molar ratios of reactants:
Reactants: Mole ratio Na O/SiO 0.28 to 0.45 SiO /Al O 8 to Ego/N320 t0 The reaction mixture may then be allowed to digest at ambient temperatures for up to 40 hours or more, in
Crystalline zeolites earth metals, and preferably metals in Groups II, III, VIII, and rare earth metals. Where a hydrogen-containing cation is used to replace the sodium, the hydrogen or decationized form of the zeolite is produced. A convenient method of preparing the hydrogen or decationized form is to subject the zeolite to base-exchange with an ammonium cation solution followed by controlled heating at elevated temperature, e.g. 600 to 1000 F., to drive ofi ammonia and water.
The hydrogen form or the ammonium form of the zeolite is composited or impregnated with platinum group metal by means of a wet impregnation or a base exchange reaction, by treatment with a platinum or a palladium salt or ammonium complex, e.g. ammonium chloroplatinate, palladium chloride, etc. palladium catalyst may be prepared by simply slurrying the desired quantity of the ammonium form of the zeolite in water, subsequently adding an ammoniacal palladium solution having the desired quantity of palladium, and mixing the resulting slurry for a short period of time at ambient.temperature. The catalyst is then preferably subjected to calcination at elevated temperatures, e.g. about 500 to 1000 F., in order to expel ammonia gas to leave a material consisting of palladium on the hydrogen form zeolite.
The above-described catalyst is most preferably used in the hydrogen or ammonium form wherein the sodium content of the sieve has been reduced with either hydrogen I ion or ammonium ion. However, under certain circumseparated and recovered, and the gas oil and residual fractions being recycled to the first-stage hydrocracking reactor after mixing with fresh feed.
The process of the invention may be further understood by reference to the accompanying drawing which is a diagrammatic view of a preferred embodiment of the invention.
Referring to the drawing, a heavy oil feed is introduced through line 11 and joins with pressurized hydrogen gas introduced through line 12. The combined stream flows through heater 13 wherein it is heated to operating temperature, and is then introduced into hydrocracking zone 15, via line 14, where it contacts the large pore amorphous catalyst of the type hereinbefore described, e.g. cobalt molybdate on alumina or silica-alumina. The efiluent from hydrocracking zone 15, containing partially converted residual petroleum fractions and gas oil, is withdrawn through line 16. Additional pressurized hydrogen may be introduced through line 19 if desired. The combined stream is again heated (or cooled if necessary) to operating temperature in heater (or cooler) 20, and introduced into hydrocracking zone 22 via line 21, where it contacts the crystalline zeolite catalyst of the type hereinbefore described. The product stream from hydrocracking zone 22, containing unconverted residual fractions, gas oil, naphtha and dry gas is sent via line 23 to fractionator 27, which serves to fractionate dry gas and gasoline from the portion of the product boiling above the gasoline range. Dry gas is removed through line 28 and 430 F.- naphtha product is removed through line 29. The portion of the product boiling above about 430 1 is recycled via line 30 and joins incoming feed flowlng in line 11. If desired, distillates such as heating oil and gas oils may be taken from the tower through line 33 rather than being recycled to the conversion process.
For purposes of simplicity the hydrogen recycle system has not been shown. Recycle of hydrogen may be accomplished in any conventional manner. For example,
hydrogen in the overhead product from the fractionator.
For example, a suitable eliiuent in line 23 to remove ammonia and hydrogen sulfide gas, and a gas-liquid high pressure separator may be inserted in line 23, with liquid product, water and recycle gas being separated therein. The recycle gas, consisting predominantly of hydrogen, may be further purified by conventional gas treating means and then recycled, under pressure, to the two hydrocracking zones. Optionally, a similar arrangement may be employed between the first and second hydrocracking zones, e.g. in line 16.
The process of the present invention is subject to various minor modifications. While the process has been described as a two-stage process, it is not intended to be limited to a two-reactor system. Each stage hereinbefore described can be readily composed of more than one reactor. For example, one modification herein contemplated comprises a three-reactor system, two reactors being included within the second-stage hereinbefore described. In this modification, the effluents from both the second and third reactors are passed to the fractionator; the gas oil and residual fractions are separately withdrawn from the fractionator; the gas oil fraction is recycled to the third hydrocracking reactor preferably containing the same crystalline zeolite catalyst as the second reactor, and instead of recycling both gas oil and residual fractions to the first reactor containing the large pore catalyst, e.g. cobalt molybdate on alumina, only the residual fractions are so recycled. Since the third reactor will be operating on low nitrogen, distillate feed, milder conditions can be employed therein.
Additionally, while it will usually be preferred to operate with separate reactors for the different types of catalysts, it is also within the contemplation of the present invention to utilize an intimate mixture of the catalysts employed.
Furthermore, while the process of the invention has been previously described with reference to fixed bed catalytic hydrocracking, it will be appreciated that various other conventional processing techniques may also be used. Thus, ebulating bed, moving bed, slurry or sump phase types of operation are to be included within the scope of the present invention.
The invention may be further understood by reference to the following examples which are not intended to be limiting.
EXAMPLE 1 This example demonstrates the ability of the large pore amorphous catalyst utilized in the first hydrocracking stage of the present process to convert a heavy oil feed to gas oil and to remove metal contaminants.
A West Texas atmospheric residuum was hydrocracked in a once-through fixed bed operation using a cobalt molybdate on alumina catalyst having an average pore diameter of about 80 A., and containing 3 wt. percent C and 12 wt. percent M00 The West Texas atmospheric residuum had a specific gravity of 19.7 API, a sulfur content of 2 wt. percent, a total nitrogen content of 2300 p.p.m., a metal content of 50 p.p.m., a Conradson carbon of 5.2% and a distillation distribution as follows:
Vol. Percent Nitrogen,
Initial to 430 F. 430 to 650 F The hydrocracking conditions and performance in converting the feed to gas oils and in removing contaminants are shown below:
10 Table l HYDROCRACKING OF WEST TEXAS ATMOSPHERIC RESIDUUM WITH COBALT MOLYBDATEALUMINA CATALYST OF ABOUT A. PORE DIAMETER Temperature, F. 760
The above data demonstrate that appreciable ash removal and conversion of residual fractions to 1000 F. material can be obtained witha desirably low naphtha make, in the first stage of the present process.
EXAMPLE II This example demonstrates the relationship between the conversion of 1000 F residual fractions and metals removal obtained with a catalyst of the type used in the second stage of the present process; namely, a crystalline alumino-silicate zeolite having a pore diameter of about 13 A. and composited with 2 wt. percent palladium.
The feed utilized was the West Texas atmospheric residuum of Example I.
The catalyst was prepared as follows:
A slurry mixture of 1655 grams of commercial sodium aluminate containing 65 wt. percent NaAlO and 5300 grams of sodium hydroxide (97% NaOH), contained in 37 pounds of water, was added with rapid stirring to 78.5 grams of a commercially available colloidal silica sol containing 30 wt. percent silica (Ludox solution, supplied by E. I. duPont de Nemours & Co., Inc.). Mixing was conducted at ambient temperature of 75 F. The total relative molar composition of the resultant reaction mixture was as follows: 27 SiO 11 Na O: 1 A1 0 366 H O. Stirring was continued to form a homogeneous mixture and the composite slurry was then kept at ambient temperature, in an open vessel with stirring for a digestion period of about 2 hours, after which time it was'heated to 212 F. The vessel was sealed and heated at 212 F. for 6 days, which was the point of maximum crystal'linity as determined by periodic sampling and analysis. The vessel was then cooled and opened, and the crystalline slurry was filtered, washed with water, and oven dried at 275 F, A sample of the crystals was calcined for 4 hours at 850 F. and analyzed to show the following composition: 64.8 wt. percent SiO 14.0 wt. percent Na O, and 20.9 wt.. percent Al A which corresponds to anapproximate molar composition of about 1.1 Na O:Al O :5.3 SiO The zeolite had an average pore diameter of about 13 A.
Five pounds of the above product crystals were then ion-exchanged at room temperature with 7.5 gallons of a 19 wt. percent aqueous solution of ammonium chloride solution. The composite solution was stirred intermittently over about a 2-hour period at room temperature and the solids were then filtered. This ion-exchange procedure was repeated with fresh solution five times at a temperature of about F. After the final treatment the filter cake was water washed to substantially remove excess chloride ion.
The resulting ammonium form of the zeolite was separated by filtration and treated with a solution containing the ammonium complex of palladium chloride in an amount suflicient to produce a 2 wt. percent palladium catalyst. Specifically, an ammoniacal palladium solution having a palladium content of 0.0504 gram of palladium per cc. was prepared by dissolving palladium chloride in aqueous ammonium hydroxide. 1702 cc. of this ammoniacal palladium solution was added to a water slurry of the ammonium form of zeolite which contained 8193 grams of zeolite and had'a solids content of 52.2 wt. percent. This was equivalent to about 0.4
cc. of palladium .solution per gram of solids in the zeolite SIDUUM WITH Pd ON CRYSTALLINE ZEOLITE CATA- LYST OF ABOUT 13 A. PORE DIAMETER Run Number Conversion, Vol.
430 F.+ to 430 F. 17 40 47 58 58 1,000 F.+ to 1,000 F. 4 7 18 21 32 Metals Removal, Wt. Percent-.- nil nil 10 21 41 As indicated, at low conversion of 1000 F.+ residual fractions, the amount of metals removal is correspondingly low, i.e. the amount of ash deposited on and retained by the catalyst is relatively small. The desirability of limiting the conversion of residual fractions in order to avoid contamination of the catalyst is thus demonstrated.
EXAMPLE III The West Texas atmospheric residuum feed of the previous examples is hydrocrackedv according to the process of the present invention.
Specifically, 1000 barrels/day of feed together with 1345 barrels/day of 430 F.+ recycle material are fed to hydrocracker containing the cobalt molybdate on alumina catalyst having an average pore diameter of about 80 A. and comprising 3 wt. percent C00 and 12 wt. percent M00 The operating conditions in hydrocracker 15 include a space velocity of l v./v./hr., a hydrogen gas rate of 10,000 s.c.f./b., a temperature of 760 F. and a pressure of l500 p.s.i.g. The conversion of l000 F.+ residual fractions to 1000 F. fractions in hydrocracker 15 is about 36 vol. percent and the hydrocracked efiluent stream contains less than 10 p.p.m. ash.
The eflluent is fed at a rate of 2425 barrels/day to hydrocracker 22 containingthe 2 wt. percent palladium on crystalline zeolite catalyst having a pore diameter of about 13 A. as prepared in Example II. The operating conditions in hydrocracker 22 include a temperature of 725 F., a pressure of 1500 p.s.i.g., a hydrogen gas rate of 10,000 s.c.f./b. and a space velocity of 0.5 v./v./hr. At these conditions, the conversion of 1000 F.+ residual, material is limited to about 10 vol. percent.
The efiluent from hydrocracker 22 is fractionated and yields about 400,000 s.c.f./d. of C -C gas, about 1250 barrels/day of C to 430 F. naphtha, about 695 'barrels/ day of 430 F./ 1000 F. gas oil and about 650 barrels/day of 1000 F.+ residual bottoms. The 1345 barrels/day of 430 F.+ material is recycled to hydrocracker 15.
If desired, a portion of the 430 F./1000 F. distillate can be separated for use as is, e.g. heating oil, or as recycle to a third hydrocracker containing crystalline zeolite catalyst.
What is claimed is:
1. A process for hydrocracking a heavy hydrocarbon oil feed contaminated with nitrogen and metal impurities and containing residual petroleum fractions, which process comprises contacting said feed, in a first reaction zone,
12 with a catalyst comprising a hydrogenation metal supported on an amorphous gel having an average pore diameter of at least about 20 angstrom units, in the presence of added hydrogen, at hydrocracking conditions sufficient to at least partially convert said residual fractions to gas oil and to substantially reduce the nitrogen and metals content of said feed; and contacting the normally liquid efiluent of said first reaction zone, in a second reaction zone, with a catalyst comprising a crystalline alumino-silicate zeolite composited with a platinum group metal, said zeolite having pore openings of about 5 to about 20 angstrom units and containing less than about 10 wt. percent alkali metal oxide, in the presence of added hydrogen, at hydrocracking conditions suflicient to at least partially convert gas oil to naphtha while limiting the conversion of residual fractions to lower boiling material to less than about 30 volume percent, based on the first reaction zone efiluent.
2. The process of claim 1, wherein said feed contains more than about 5 p.p.m. metals, more than about p.p.m. nitrogen and about 2 to 100 volume percent of a residual petroleum fraction, said fraction containing at least about 500 p.p.m. nitrogen and about 5 to about 1000 p.p.m. metals.
3. The process of claim 1, wherein said alkali metal oxide is Na O and said zeolite has been base exchanged with a hydrogen containing cation. v
4. The process of claim 1, wherein said zeolite is composited with at least about 0.1 wt. percent platinum group metal.
5. The process of claim 1, wherein said zeolite is composited with about 0.5 to 10 wt. percent platinum group metal.
6. The process of claim 5, wherein said platinum group metal is palladium.
' 7. The process of claim 1, wherein the catalyst in said first reaction zone comprises a Group VI-metal and a Group VIII metal supported on an amorphous metal oxide base selected from the group consisting of silica, oxides of metals ln Groups II-A, III-A, and IV-B, and mixtures thereof.
8. The process of claim 1, wherein the hydrocracking conditions in said first reaction zone include a temperature of from about 600 to about 850 F., a pressure of from about 500 to about 3000 p.s.i.g., a liquid hourly space velocity of from about 0.1 to about 10.0 volumes of feed per volume of catalyst per hour, and a hydrogen rate of from about 3000 to about 20,000 standard cubic feet per barrel of feed; and wherein the hydrocracking conditions in said second reaction zone include a temperature of from about 500 to about 800 F., a pressure of from about 500 to about 3000. p.s.i.g., a liquid hourly space velocity of from about 0.1 to about 10.0 volumes of feed per volume of catalyst per hour, and a hydrogen rate of from about 3000 to about 20,000 standard cubic feet per barrel of feed.
9. The process of claim 8, wherein the conversion of residual fractions to lower boiling material is limited .to less than about 10 volume percent.
10. A process for hydrocracking a heavy hydrocarbon oil feed to obtain lower boiling product, said feed being contaminated with about 5 to 300 p.p.m. metals and more than about 100 to about 20,000 p.p.m. nitrogen, and containing about 2 to 100 volume percent of a residual petroleum fraction containing at least about 500 p.p.m. nitrogen and about 5 to about 1000 p.p.m. metals, which process comprises:
(1) contacting said feed, in a first reaction zone, with a catalyst comprising a Group VI metal and a Group VIII metal supported on an amorphous metal oxide base having an average pore diameter of at least about 20 angstrom units, said metal oxide being selected from the group consisting of silica, oxides of metals in Groups IIA, III-A, and IV-B, and mixtures thereof, in the presence of added hydrogen, at
hydrocracking conditions sufiicient to at least partially convert said residual fraction to gas oil and to substantially reduce the nitrogen and metals content of said feed, said hydrocracking conditions including a temperature of from about 600 to about 850 F., a pressure of from about 500 to about 3000 p.s.i.g., a liquid hourly space velocity of from about 0.1 to about 10.0 volumes of feed per volume of catalyst per hour, and a hydrogen rate of fro-m about 3000- to about 20,000 standard cubic feet per barrel of feed,
(2) subsequently contacting the normally liquid effluent of said first reaction zone, in a second reaction zone, with a catalyst comprising a crystalline alumino-silicate zeolite composited with at least about 0.1 wt. percent platinum group metal, said zeolite having pore openings of about 5 to about 20 angstrom units and containing less than about wt. percent Na O, in the presence of added hydrogen, at hydrocracking conditions sufiicient to at least partially convert gas oil to naphtha while limiting the conversion of residual fractions to lower boiling material to less than about 30 volume percent based on the first reaction zone effluent, said hydrocracking conditions including a temperature of from about 500 to about 800 F., a pressure of from about 500 to about 3000 p.s.i.g., a liquid hourly space velocity of from about 0.1 to about 10.0 volumes of feed per volume of catalyst per hour, and a hydrogen rate of from about 3000 to about 20,000 standard cubic feet per barrel of feed, and
(3) recovering naphtha from the efiluent of said second reaction zone.
11. The process of claim 10, wherein the catalyst in' said first reaction zone comprises cobalt molybdate supported on alumina.
12. The process of claim 10, wherein said catalyst in said first reaction zone has an average pore diameter of about 50 to about 300 angstrom units.
13. The process of claim 10, wherein said catalyst in said first reaction zone has an average pore diameter of about 80 to about 150 angstrom units.
14. The process of claim 10, wherein said zeolite is composited with about 0.5 to 10 wt. percent platinum group metal and has a silica to alumina mol ratio of 2.5 to 14.
15. The process of claim 14, wherein said platinum group metal is palladium.
16. The process of claim 10, wherein said zeolite has pore openings of about 6 to about 15 angstrom units and is composited with about 1 to 5 wt. percent palladium.
17. The process of claim 10, wherein said zeolite is composited with about 1.5 to 3 wt. percent palladium, and has a silica to alumina mol ratio of about 3 to 10 and pore openings of about 7 to about 13 angstrom units.
'18. The process of claim 10, wherein the conversion of residual fractions to lower boiling material is limited to less than about 10 volume percent.
19. The process of claim 10, wherein the hydrocracking conditions in said first reaction zone include a temperature of about 750 to 825 F., a pressure of about 1500 to 2500 p.s.i.g., a liquid hourly space velocity of about 0.5 to 2.0 v./v./hr., and a hydrogen rate of about 5000 to 15,000 s.c.f. per barrel of feed, with the conversion of residual fractions to gas oil in said first zone being maintained within the range of about 20 to about volume percent; and wherein the hydrocrackingconditions in said second reaction zone include a temperature of about 725 to 775 F., a pressure of about 1500 to 2500 p.s.i.g., a liquid hourly space velocity of about 0.5 to 2.0 v./v./hr., and a hydrogen rate of about 8000 to 15,000 s.c.f. per barrel of feed, with the conversion of gas oil to naphtha in said second zone being maintained within the range of about 20 to 80 volume percent, and the conversion of residual fractions to lower boiling material in said second zone being limited to less than about 10 volume percent.
20. The process of claim 10 which further comprises recycle of residual fractions and gas oil contained in the effluent of said second reaction zone to said first reaction zone.
References Cited by the Examiner UNITED STATES PATENTS 2,839,450 6/1958 Oettinger 208 2,983,670 5/1961 Seubold '208111 3,132,087 5/ 1964 Kelley et a1. 20859 3,172,838 3/1965 Mason et a1. 20861 DELBERT E. GANTZ, Primary Examiner.
ALPHONSO D. SULLIVAN, PAUL M. COUGHLAN,
Examiners. A. RIMENS, Assistant Examiner.
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|U.S. Classification||208/59, 208/111.3, 208/111.35, 208/111.15, 208/110|
|International Classification||C10G65/00, C10G47/00, C10G47/18, C10G65/10|
|Cooperative Classification||C10G65/10, C10G47/18|
|European Classification||C10G47/18, C10G65/10|