DE2016774A1 - A catalytic composition comprising nickel, tungsten and a substrate of aluminosilicate material and an inorganic oxide, and a method of using this composition - Google Patents

Description

PATENT LAWYERS

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8 MDNOH EN 2,

STANDARD OIL COMPANY, Chicago, Illinois 60680, U.S.A.

A catalytic composition comprising nickel, tungsten and a support material of aluminosilicate material and an inorganic oxide, and a method of using this composition

The catalytic composition includes a hydrogenation component and a cocatalytic acidic cracking carrier. The hydrogenation component comprises a member of the following group, which consists of 1.) elemental nickel and elemental tungsten, 2.) their oxides, $.) their sulfides and 4.) mixtures thereof. Of the The carrier comprises a large pore crystalline aluminosilicate material and a refractory inorganic oxide. The preferred one Aluminosilicate material is ultra stable large pore crystalline aluminosilicate material; the preferred refractory inorganic Oxide is silicon dioxide / aluminum oxide. In general can nickel and tungsten in a total amount in the range of about 0.03 "to about 0.20 gram-moles of nickel oxide (NiO) and tungsten trioxide (WO,) be present per 100 g of the catalyst and in a tungsten to nickel ratio of about 0.6 to about 12.0.

Typical processes using this catalytic composition is a hydrodenitrogenation process and hydrocracking process.

The present invention relates to a catalytic composite and hydrocarbon conversion processes in which these

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catalytic composition is used. The catalytic composition comprises a hydrogenation component and a cocata-lytic acidic cracking component. The Kohlenwasserstoffkonversiorisverfahreii can reduce be for the treatment of mineral oils to substantially whose nitrogen and / or sulfur or a method may be a method for the treatment of mineral oils, which has the result that at least some of the Kohleriwasserstoffmoleküle of oil be chemically altered so that petroleum oils are formed with different properties. The first type of process can be referred to as a hydrodenitrogenation process and / or a hydrodesulfurization process. A hydrocracking process is a typical example of the latter type of process.

It is well known to those skilled in the art that catalysts which are useful in hydrodenitrogenation and / or hydrodesulfurization of petroleum hydrocarbon streams, generally one Hydrogenation component, W3> cobalt, molybdenum, nickel, tungsten or other sulfur active materials, their compounds or mixtures, and a non-acidic or weakly acidic carrier such as activated alumina. Typical hydrocracking catalysts include a hydrogenation component such as cobalt, molybdenum, nickel, platinum, palladium, their compounds or mixtures thereof, and an acidic cracking component. Such an acidic cracking component can be a mixture of refractory inorganic oxides, such as silicon dioxide / aluminum oxide or boron oxide-aluminum oxide or an acid-treated inorganic oxide such as fluorinated alumina or a mixture of a particular type of aluminosilicate material and a refractory inorganic Oxide, such as Y-type molecular sieves and silica / alumina.

A new catalytic composition has been found that has superior activity for certain types of hydrocarbon conversions as for hydrodenitrogenation and hydro he acken.

A new catalytic composition has been found. In general

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mine, this catalytic composition includes a hydrogenation component and a cocatalytic acidic cracking carrier, wherein the hydrogenation component is selected from the following group, which consists of 1.) elemental nickel and elemental Tungsten, 2.) their oxides, 3 ·) their sulfides and 4.) mixtures thereof and the carrier comprises a large-pore crystalline aluminosilicate material and a refractory inorganic oxide. The large pore crystalline aluminosilicate material can be ultra stable be large pore crystalline aluminosilicate material, which is characterized by a maximum unit cell dimension of 24.55 Å units, hydroxyl infrared bands around 3700 cm and around

-1
3625 cm and an alkali metal content of less than 1% by weight, preferably less than 0.5% by weight. The refractory inorganic oxide can be silicon dioxide / aluminum oxide. The large-pored crystalline aluminosilicate material is preferably suspended in a matrix of the heat-resistant inorganic oxide. The carrier is prepared by mixing the large-pore crystalline aluminosilicate material in a finely divided state with a sol or a gel of the refractory inorganic oxide and then drying the large-pore crystalline aluminosilicate material is present in an amount ranging from about 5% by weight to about 70% by weight based on the weight of the catalytic support, the preferred amount of the large pore crystalline aluminosilicate material being an amount exceeding 30% by weight based on the weight of the support.

Roughly speaking, -Nickel and Tungsten can be used in one total in the range of about 0.03 "to about 0.20 gram-moles of nickel oxide (HiO) and tungsten trioxide (WO,) may be present per 100 g of the catalyst and in a tungsten to nickel ratio in the range of about 0.6 to about 12.0. Preferably the nickel and tungsten can be present be in a total amount ranging from about 0.05 to about 0.18 Gram-moles NiO and WO, per 100 g of catalyst and for one tungsten to nickel ratio in the range of about 1.2 to about 8.0. Preferably Both nickel and tungsten are available in a total in the range of about 0.07 to about 0.13 gram-moles NiO and WO, per 100 grams of the catalyst and at a tungsten to nickel ratio in the range of about 2 to about 5 · The optimal tungsten to nickel ratio is 3.0.

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A typical hydrocarbon process of the present invention is eating a process for the hydrodenitrogenation of petroleum hydrocarbons » hydrogen coating material that is as much as 2000 ppm or more Contains nitrogen. This hydrodenitrogenation process includes contacting the feed material in the presence of hydrogen under hydrodenitrogenation conditions with that of the invention catolytic composition.

Another hydrocarbon conversion process according to the invention is a hydrocracking process, the hydrocracking process consists in making the petroleum-hydrocarbon feedstock in the presence of hydrogen and under hydrocracking conditions with the catalytic composition according to the invention contacted.

In the following, reference is made to the attached drawings.

FIG. 1 shows a highly simplified schematic flow diagram a preferred embodiment of a method according to the invention. In this embodiment the method is a method for hydrocracking petroleum hydrocarbons. Certain Additional devices such as valves, pumps, heat exchangers and the like are not shown since it is clear to the person skilled in the art that he must use them at certain selected points in the flow path of the process.

Figure 2 shows the relationship that was used to determine the relative Calculate the activity for the hydrodenitrogenation of the various catalysts described below.

Figure 3 shows the hydrocracking performances of two embodiments of catalytic compositions according to the invention and compares these performances with the performances more typical Hydrocracking catalysts.

FIG. 4 shows an activity contour diagram for the nickel-tungsten system the catalytic composition according to the invention.

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Two Important Processes for Fractional Petroleum Hydrocarbon Conversion are hydrodenitrogenation processes and hydration processes. A hydrodenitrogenation process is a petroleum refining process that wherein the nitrogen content of a petroleum hydrocarbon feed material is substantially reduced with increased Temperature and pressure / presence of a hydrogen containing Gas and a hydrodenitrogenation catalyst. When organic nitrogen is converted to ammonia, it becomes hydrogen consumed. Furthermore, at least a portion of the sulfur will be used in any particular hydrocarbon feedstock is present, converted to hydrogen sulfide.

In a typical hydrodenitrogenation process, naphthas, gas oils and even hydrocarbon feedstocks can be treated over the entire boiling range to cause their nitrogen content to be substantially reduced. The Beschickungcmaterialien can to about 538 G ⁰ in a range from about 177 (350 to 1000 ⁰ F) boiling or above and up to 2000 ppm nitrogen or more.

In the hydrodenitrogenation process according to the invention, the feed material to be treated is contacted in the hydrodenitrogenation reaction zone with the catalyst described below in the presence of a hydrogen-donating gas. An excess of hydrogen is maintained in the reaction zone. Preferably, a hydrogen to oil ratio of at least 890 cbm hydrogen per ecm of feed material (5000 standard cubic feet of hydrogen per barrel of feed SCPB) is used and the hydrogen to oil ratio can be up to 7120 cbm per cbm (40,000 SCFB) are sufficient. Preferably, a hydrogen to oil ratio of about 1780 to about 4450 cbm per cbm (about 10,000 to about 25,000 SCFB) is employed. The total pressure is usually between about 35 »2 atmospheres (500 psig) and about 352 atmospheres (5000 psig), and preferably between about 70.3 atmospheres (1000 psig) and about 140.6 atmospheres (2000 psig). The average catalyst bed temperature is in the range of about 204 ⁰ C (4CO ⁰ F) to about 427 ° C (800 ⁰ F), and preferably in a range of about 343 ° C (65O ° F) to about 399 ° C (750 ⁰ F). The throughput of liquid volume per hour (LHSV) is typically in the range

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from about 0.1 to about 20 volumes of hydrocarbon per hour per volume of catalyst and preferably in the range of about 0.5 to about 5 volumes of hydrocarbon per volume of catalyst.

For use in the present invention, the catalyst can Hydrodenitrogenierungsverfahren a few hours at a pressure ranging from about 0 atm (0 psig) to about 140.6 atm (2000 psig) in flowing hydrogen at a temperature in the range of about 177 ° C (35O ⁰ F) are pre-treated to about 399 ° C (75O ⁰ F). The hydrogen flow rate can range from about 1.24 cbm per hour and per kg of catalyst (20 standard cubic feet per hour per pound of catalyst SCFHP) to about 12.4 cbm per kg per hour (200 SCFHP). Typically the pressure and hydrogen flow rate will be the same as used in the subsequent feed treatment. Alternatively, auclj. a stream of hydrogen can be used to pretreat the catalyst containing small amounts of hydrogen sulfide, e.g. B. up to about 10 vol .-% contains.

Hydrocracking is a general term used for petroleum refining processes in which hydrocarbon feedstocks having relatively high molecular weights are converted to low molecular weight hydrocarbons at elevated temperature and pressure in the presence of a hydrocracking catalyst and in the presence of a hydrogen containing gas. In the conversion of organic nitrogen and sulfur to ammonia and / or hydrogen sulfide and in the Zei disturbance of the compounds with high molecular weight to compounds with low molecular weight and in the saturation of olefins and other unsaturated compounds, hydrogen is consumed. In Hydroerackungsverfahren hydrocarbon feedstocks are, such as gas oils, typically (F 1000 ⁰⁾ boiling in the range of about 177 ° C (35O ⁰ F) to about 538 ° C, catalytically treated cycle oils, between about 177 ° C (350 ⁰ F ) and 454 ° C (850 ° F) to low molecular weight products such as gasoline-boiling products and light distillates.

Typical hydrocarbon feedstocks contain nitrogen

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Substance compounds in such amounts that the amount of the existing Nitrogen is greater than 20 ppm. The nitrogen tends to decrease the activity of that used in the hydrocracking reaction To reduce the catalyst. Such a catalytic reduction Activity leads to ineffective process performance, poor product distributions and yields. The more the nitrogen content increases, the higher the reaction temperatures are required to maintain a given degree of conversion. In general, the nitrogen content of a hydrocarbon feed can be reduced by using this feedstock subjected to a batch preparation treatment. In one in such a case the nitrogen compounds are converted to ammonia. In addition, sulfur is converted to hydrogen sulfide.

Generally use hydrocracking processes at lower levels Temperature to maximize products in the gasoline boiling range two levels. In the first stage of the batch prep stage, the feedstock is treated with hydrogen to remove nitrogen and sulfur typically found in common raffinate ion feedstocks can be found. In the second The pre-treated hydrocarbon becomes the stage of the hydrocaking stage electricity converted to lower-boiling products.

There are also one-step hydro-baking processes. In a one-step process takes place in the first part of the catalyst bed or in the first reactor or in the first two or three reactors of a multi ire actuator system denitrogenation and desulfurization instead of. Therefore, denitrogenation, desulfurization and hydrocracking by the same catalyst can be carried out in a one-step process be effected. However, two different catalysts can also be used because the first catalyst for the Denitrogenation and desulfurization is used and the second catalyst for hydrocracking. However, the the denitrogenation and the desulfurization resulting ammonia and hydrogen sulfide passed over the second catalyst together with the hydrocarbons to be hydrocracked by the second catalyst. In a one-step process does not find a separation step between the first catalyst and

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the second catalyst instead of separating the ammonia and hydrogen sulfide from the hydrocarbons.

The catalytic composition according to the invention can advantageously are used for hydrocracking petroleum hydrocarbon fractions. It can be used in either one One stage hydrocracking process or in a two stage hydrocracking process .

In the Hydrocrackungsverfahren · the introduced hydrocarbon feedstock in the range between about 177 ° C (35O ⁰ F) and about 538 ° C (1000 ⁰ F) boiling. If one so performs the process to maximize gasoline production, the feed material preferably has a final boiling point that is not higher than about 171 to 399 ° C (700 - 75O ⁰ F). Typically, a light catalytically treated is circulated oil or light gas oil original or mixtures thereof, the (650 ⁰ F) / used in the range of about 177 ° C (35O ⁰ F) to about 34-3 ⁰ C as a feed material. The feed can be pretreated to remove compounds of sulfur and nitrogen. The feed can have a substantial sulfur content ranging from 0.1 to 3 weight percent and nitrogen can be present in an amount up to 500 ppm or more. The temperature flow rate and other process variables can be adjusted to compensate for the effects of nitrogen on hydrocracking catalyst activity.

The hydrocarbon feed preferably contains a substantial one Amount of cyclic hydrocarbons, d. H. more aromatic and / or naphthenic hydrocarbons, since such has been found to be Hydrocarbons are particularly suitable for Horstel-. processing of highly aromatic hydrocracked gasoline products. Preferably the feed contains at least about 35 to 40% aromatics and / or naphthenes.

In the hydroacking process, the feed material is mixed with mixed with a hydrogen donating gas and heated to the hydrocracking temperature preheated and then into one or more hydrocracking reactors convicted. Preferably the feed

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substantially, and completely evaporated, "before it is introduced into the reactor system. For example, it is preferred that the feed be completely evaporated before it has passed through more than about 20 % of the catalyst bed in the reactor mixed vapor liquid phase and the temperature, pressure, return, etc. can then be adjusted to the particular feed material to achieve the desired degree of evaporation.

The feed is in the hydrocracking reaction zone with the catalyst described below in the presence of a Hydrogen-releasing gas contacted. In the hydro-drying process hydrogen is consumed and an excess of hydrogen is maintained in the reaction zone. Preferably a hydrogen to oil ratio of at least 890 cbm per cbm (5000 SOFB) is used and the hydrogen to oil ratio is used can range up to 3560 cbm per cbm (20,000 SCFB). Preferably, a ratio of about 1424 to about 2670 cbm per cbm (8000 up to 15,000 SCFB). From the point of extension of the When viewed from catalyst activity, it is a high partial pressure of hydrogen desirable.

The hydrocracking reaction zone is operated at elevated temperatures and pressures. The total hydraulic pack pressure is usually between about 49.2 and 281 atm (700 and 4000 psig), and preferably between about 70.3 and 127 atm (1000 and 1800 psig). The partial pressure of hydrogen will generally range from about 47.5 atmospheres (675 psig) to about 279 atmospheres (3970 psig). The hydrocracking reaction is inherently exothermic and thus there is a temperature rise across the catalyst bed. The average Hydrocrackungskatalysatorbett temperature is between about 343 and 454 ⁰ C (650 and 850 ⁰ F), and preferably a temperature between about 360 and 427 ° C (680 and 800 ⁰ F) is maintained. The liquid volume per hour throughput (LHSV) is typically in the range of about 0.1 to about 10 volumes of hydrocarbon per hour per volume of catalyst, preferably in the range of about 0.5 to about 5 volumes of hydrocarbon per hour. Hour and per volume of catalyst. Optimally, the LHSV is in the range from about 1 to about 2.

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The catalyfcische composition according to the invention can also be used for The disproportionation of petroleum hydrocarbon can be used, with alkyl groups of aromatic hydrocarbon atoms be transferred from one molecule to another. This latter method consists in taking the petroleum hydrocarbon stream in a hydrocarbon conversion zone with a catalytic composition according to the invention in the presence of a Hydrogen donating gas under suitable hydrocarbon disproportionation conditions contacted.

For the disproportionation typical feedstocks petroleum hydrocarbon streams containing monocyclic and dicyclic aromatic hydrocarbons boiling below about 34-3 ⁰ C (65O ⁰ P). Such hydrocarbon granules can be a petroleum hydrocarbon fraction which contain aromatics or which are aromatic hydrocarbons. For example, the feedstock may contain toluene, o-xfiol, m-xylene and p-xylene, trimethylbenzenes and tetramethylbenzenes. Typically, in the disproportionation process, the feedstock is mixed with a hydrogen donating gas and preheated to a suitable disproportionation temperature and then introduced into the disproportionation reaction zone, which may include one or more reaction vessels. Preferably, the feed is substantially completely evaporated before it is introduced into the reaction zone.

The feed material is contacted in the disproportionation reaction zone with the catalyst described below in the presence of a hydrogen donating gas. Preferably, a hydrogen to oil ratio of at least 178 cbm per cbm (100 SCFB) is employed and the hydrogen to oil ratio can range up to 8900 cbm per cbm (50,000 SCFB). Preferably the hydrogen to oil ratio is in the range of about 890 to about 534-0 cbm per cbm (5000 to 30,000 SCFB). Other operating conditions include an elevated temperature in the range of about 371 ° C (700 ⁰ F) and about 593 ° C (1100 ⁰ F), preferably between about 454 ⁰ C (850 ⁰ F) and about 538 ⁰ G (1000 ° F ); an elevated pressure in the range of about 7 * 0 atmospheres (100 psig) and about 70.3 atmospheres (1000 psig), preferably between about 14.1 atmospheres (200 psig) and

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about 35.2 atmospheres (500 psig); and an hourly weight throughput (WHSV weight hourly space velocity) in the range of about 0.1 and about 20 weight units of hydrocarbon per hour per unit weight Catalyst, preferably between about 1 and about 10 weight units of hydrocarbon per hour per weight unit of the catalyst. The exothermic demethanization reaction that occurs in the disproportionation reaction zone can be controlled are obtained by treating the catalyst with sulfur compounds such as hydrogen sulfide and carbon disulfide either before or at the start of the disproportionation reaction.

In addition to the fact that the catalytic composition of the present invention represents the catalyst for hydrodenitrogenation for hydrocracking for disproportionation, is that according to the invention catalytic composition a suitable catalyst for the isomerization of alkyl aromatics, the transalkylation r: H see aromatic hydrocarbons, the hydrodealkylation of hydrocarbon-substituted aromatic compounds and for reforming petroleum hydrocarbon naphtha streams.

The catalytic composition according to the invention comprises a hydrogenation component on a cocatalytic acidic cracking support. The hydrogenation component comprises one ingredient selected from the group consisting of 1.) elemental nickel and elemental wolf races, 2.) their oxides, 3.) their sulfides and 4.) Mixtures of these. The cocatalytic acidic cracking carrier comprises a large pore crystalline aluminosilicate material and a hit-resistant inorganic oxide.

The hydrogenation component can be deposited on the acidic cracking carrier or can be incorporated into the acidic cracking carrier by impregnation with heat-decomposable salts of the desired metals. Each of these metals can be applied to the carrier separately or together with the other metals. Alternatively, the hydrogenation component metals can be coprecipitated with a hydrogel of the refractory inorganic oxide In this latter process, finely divided processes are used crystalline aluminosilicate material mixed well with the hydrogel

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and then each metal is added separately to the hydrogenation component added to the mixture in the form of a heat-decomposable salt of this metal. The mass is then dried in the following and calcined to decompose the salts and remove the unwanted anions.

The composition according to the invention has a hydrogenation component which is selected from the following group

1.) elemental nickel and elemental tungsten, 2.) their oxides, 3 ·) their sulfides and 4-.) Mixtures thereof. In general nickel and tungsten can be present in a total amount ranging from about 0.05 to about 0.20 gram-moles of NiO and WO * per 100 grams Catalyst and in a tungsten to nickel ratio in the range of about 0.6 to about 12.0. Preferably this can be nickel and tungsten be present in a total amount ranging from about 0.05 to about 0.18 g-molesNiO 2 and WO, per 100 g of catalyst and one Tungsten to nickel ratio in the range from about 1.2 to about 8.0. Preferably the nickel and tungsten can be present in one Total amount ranging from about 0.07 to about 0.1 i> g-mol NiO and WOx per 100 g catalyst and with a tungsten to nickel ratio ranging from about 2 to about 5 · The optimal tungsten too Nickel ratio is 3jO.

The cracking carrier of the catalytic composition according to the invention comprises a large pore crystalline aluminosilicate material and a refractory inorganic oxide. Albeit various large pore crystalline aluminosilicate materials in the The catalytic composition of the present invention can be used as an ultrastable large pore crystalline aluminosilicate material the preferred aluminosilicate material. The cracking carrier material can be from about 5 wt .-% to about 70 wt .-% contain large pore crystalline aluminosilicate material. Preferably, the cracking vehicle can be from about 30% to about Contain weight percent aluminosilicate material. It is preferred that the large pore crystalline aluminosilicate material fully dispersed and in a porous matrix of refractory inorganic oxide is suspended. A preferred refractory inorganic oxide is silicon dioxide / aluminum oxide. Both silica / alumina cracking catalysts with low aluminum oxide

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_ 20T6774

A high aluminum oxide content can be present in the carrier material the catalytic composition of the present invention can be used. In general, silica contain alumina cracking catalysts with low alumina, which are preferred, about 10 to about 15 weight percent alumina. Silica / alumina cracking catalysts high alumina contain from about 20 to about 40 weight percent alumina.

The properties of many large pore crystalline aluminosilicate materials and methods of making them are known. Their structure includes a network of relatively small aluminosilicate cavities, which are connected to one another by a large number of pores. These pores are smaller than the cavities and have a substantially uniform diameter at its narrowest cross-section. Basically, the network of cavities is a solid one three-dimensional and ionic network of silicon dioxide and aluminum oxide tetrahedra. These tetrahedra are common by Oxygen atoms cross-linked with one another. In the crystal structure of the In the aluminosilicate material, cations are included to provide electrovalence to balance the tetrahedron. Examples of such cations are metal ions, hydrogen ions and hydrogen ion precursors, like ammonium ions. Using a process known as cation exchange, one cation can be exchanged for another be replaced. The crystalline aluminosilicate materials used in the catalytic compositions of the present invention are large-pore materials. The term "large pore material" means that the material has pores that are sufficient are large to allow the penetration of benzene molecules and larger molecules and the escape of reaction products. When used in petroleum hydrocarbon conversion processes, it is preferred to use a large pore aluminosilicate material with a pore size of at least 9 to 10 Å units. The large pore crystalline aluminosilicate materials used in the catalytic compositions of the present invention have such a pore size.

The preferred large pore crystalline aluminosilicate material is stable to elevated temperatures and stable to repeated cycles Humidification-drying cycles. This stability is demonstrated

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by the surface after calcination at 940.5 ⁰ C (1725 ° F). After calcination at a temperature of 940.5 ° C (1725 ⁰ F) for a period of 2 hours, a surface is maintained, which is greater than "150 square meters per gram. In addition, this stability is demonstrated by the surface after a steam treatment with an atmosphere containing 25% steam, and at a temperature of 829.5 ° C (1525 ⁰ F) for 16 hours. the surface after this steam treatment is greater than 200 sq m / g.

The ultra stable large pore crystalline aluminosilicate material having particularly good stability towards wetting, which is defined to maintain by the ability of a particular Aluminosilicatmaterials, its surface or its nitrogen adsorption capacity after contact with water or water vapor _t A sodium form of the ultra-stable large pore crystalline aluminosilicate material s (about 2.15 Wt% sodium) showed a loss of nitrogen adsorption capacity that was less than 2% per humidification when stability to humidification was examined by subjecting the material to a number of consecutive cycles, each cycle consisting of humidification and a drying existed. The ultra-stable large pore crystalline aluminosilicate material used in the catalytic composition of the present invention has unit cell cubic dimension and hydroxyl infrared bands that distinguish it from other aluminosilicate materials.

The cubic unit cell dimension of the ultra-stable, large-pored crystalline aluminosilicate material ranges from about 24.20 ^ units to about 24.55 Å units.

The hydroxyl infrared bands obtained with the ultra-stable large pore crystalline aluminosilicate material are a band

—1 —1

de around 3750 cm, one band around 3700 cm and one band around 3625

—1 —1

cm. The band around 3750 cm can be found in many decationized aluminosilicate materials be found in the hydrogen form, j «3-

—1 —1

but the band around 3700 cm and the band around 3625 cm are characteristic for the ultra stable large pore crystalline aluminosilicate material used in the catalytic composition of the present invention.

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The ultra-stable large pore crystalline aluminosilicate material is also characterized by an alkali metal content of less than 1 % .

An example of the ultra stable large pore crystalline aluminosilicate material is Z-1WS, which is described in U.S. Patent 3,293,192.

The cocatalytic acidic cracking carrier of the invention catalytic composition can be well known by a variety Processes can be manufactured and formed into pellets, spherules and extrudates of the desired size. For example the large pore crystalline aluminosilicate material can be pulverized into finely divided material and this latter material is then finely mixed with the silica-alumina cracking catalyst. The finely divided crystalline aluminosilicate material can be mixed well with a hydrosol or hydrogel of the silica material. When a well mixed hydrogel has been obtained, this hydrogel can be dried broken into pieces of the desired shape and size. The hydrogel can also consist of small spherical particles are reshaped by conventional spray-drying processes or other equivalent procedures.

A preferred embodiment of the hydrocarbon conversion process according to the invention FIG. 1 is a single stage hydrocracking process shown in FIG. This embodiment is for explanatory purposes only and is not intended to be within the scope of the limit the present invention.

Referring to Figure 1, fresh Kohlenwasserstoffbeschickungsiiaterial boiling in the range of about 204 ⁰ C (400 ° F) to about 333 ° C (632 ° F) and containing 0.25 wt .-% of sulfur and 159 ppm nitrogen, from the Source 10 received. This particular feed material, a mixture of 30% by volume of light fresh gas oil (LVGO) and 70% by volume of light catalytically treated circulating oil (LOGO) is pumped through the lines 11 and 12 with the aid of the pump 13. Hydrogen-releasing gas is brought through line 14 and into line 12 to become intimate with it

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the fresh hydrocarbon feed that is in it, to mix. The resulting hydrogen / hydrocarbon mixture is through the line 15 »the heat exchanger 16, the Line 17 is introduced into furnace 18. The hot hydrogen / hydrocarbon mixture is then through line 19 into the The top of the hydrocracking reaction zone 20 is introduced. The hydrocracking reaction zone 20 consists of a number of successive catalyst beds, which are replaced by inert, heat-resistant Materials are separated, such as by aluminum oxide balls. Catalyst beds 21, 22 and 23 are shown in FIG. However, should the number of three catalyst beds in no way limit the scope of the present invention and one can imagine that more or fewer catalyst beds in the hydrocracking reaction zone can be used as just three.

The catalyst in each of these catalyst beds is a preferred one ATis embodiment of the catalytic composition according to the invention. The catalyst comprises nickel and tungsten and / or their oxides and / or sulfides on a cocatalytic acidic Cracking carrier, the ultra stable large pore crystalline aluminosilicate material suspended in a porous matrix of silicon dioxide / aluminum oxide with low alumina content. Nickel and tungsten are present in a total amount ranging from about 0.05 to about 0.18 gram-moles NiO and WO, per 100 grams of catalyst and in a tungsten to nickel ratio in the range of

approximately
about 1.2 to / 8.0. The acidic cracking carrier comprises about 35% by weight of the ultra-stable large pore crystalline aluminosilicate material based on the weight of the carrier.

The temperature of the hot hydrogen / hydrocarbon Beschik- · kung mixture which is introduced into the reactor 20 is in the range of about 343 ⁰ C (650 ⁰ F) to about 366 ° C (69O ⁰ F) when the reaction is at Beginning is started and this temperature is increased in stages by the extent to which the batch proceeds in order to compensate for the gradual decrease in catalyst activity. Because the hydrocracking reaction is exothermic, the temperature of the reactants tends to increase as the reactants flow through the catalyst bed. To control the temperature rise and to set the maximum temperature in the

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Reactor "limit" is a liquid back-up stream in the reaction zone 20 introduced at points between the catalyst beds 21, 22 and 23 by means of the lines 24 and 25 · This liquid quench material is fresh feed from feed line 12 and / or recirculated oil from Reflux line 43 »which is described below.

The process conditions include a temperature of about 343 ° C (650 ⁰ F) (850 ⁰ F), a LHSV of about 1 to about 5 and a pressure of about 87.9 atm to about 454 ⁰ C (1250 psig), a hydrogen - 'recycling rate of around 1602 cbm / cbm (9000 SCFB). The effluent from the reaction zone 20 is passed through the lines 26 and 27 through the cooler 28 to the high pressure separator 29. Washing barrels can be introduced into line 26 by means of line 30. Wash water is mixed with the hydrocracked effluent in line 26 and, after passing through cooler 28 and line 27, is separated as an aqueous phase in high pressure separator 29. The wash water dissolves and removes the ammonia and hydrogen sulfide contained in the effluent. This washing water, which contains the dissolved ammonia and hydrogen sulfide, is withdrawn from the high-pressure separator 29 with the aid of the line 31. The gas which separates from the liquid in the high pressure separator 29 is removed from the high pressure separator 29 via line 32, compressed with a gas compressor 33 and recycled via line 34 by means of line 34 to line 14 and introduced into the hydrocarbon stream, by passing it through line 12 for subsequent reintroduction into the hydrocracking reaction zone 20.

The liquid hydrocarbons are extracted from the high pressure separator 29 removed via line 35 and passed to a low-pressure separator 36 with the aid of line 35. The gas phase from the Low pressure separator 36, which contains essentially light hydrocarbons and hydrogen, is via line 37 as Purge gases withdrawn, conveniently used as fuel gases will. The liquid hydrocarbon layer from the low pressure separator 36 is withdrawn therefrom via line 38 and through line 38 from fractionator 39. In the fractionator 39

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the liquid hydrocarbons are fractionated into a light gasoline fraction, a heavy gasoline fraction and soil fractions. The light gasoline fraction is from the fractionator 39 with With the help of the line 40, the heavy gasoline fraction is withdrawn with the help of the line 41, the soil fraction with the help of the lines 42. The soil fraction is via lines 42 and 43 and the return pump 44 is returned to the fresh hydrocarbon feed to be added in line 12 for introduction into the upper part of the Hydrocrackungsreaktionsζone 20 with Using the lines 12, 15, 17 and 19, the heat exchanger 16 and the furnace 18 or in the quenching lines 24 and 25 with Aid of the line 12 introduced. Undesired soil fractions can be removed from the system with the aid of line 45.

The heavy gasoline fraction from the fractionator 39 is with the help withdrawn from line 41 for use in other refining process units such as a reformer. The light gasoline fraction is withdrawn through line 40 for use in gasoline blends. Although the fractionator 39 only as however, it is to be understood that other satisfactory recovery systems can be used and that they are within the scope of the present invention. Such other suitable recovery systems are known to those skilled in the art. Alternatively, a stream of butane can also be drawn off. Fresh hydrogen from source 46 is passed through line 47 and compressed in the compressor 48. The compressed fresh hydrogen is then passed through lines 49 and 14 to to be mixed with the hydrogen / lack of hydrogen mixture and to be introduced into the hydrocracking reaction zone 20.

Quench gas may alternatively be used to control the temperature in the hydrocracking reaction zone 20. This Quench gas can recycle part of the hydrogen containing sgases passing through line 34. It can go out the line 34 can be removed by means of the lines 51 and 52, to the hydrocracking reaction zone 20 at the interstices between the catalyst beds in the H ;, "ocracking reaction zone 20 to be initiated.

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The following examples are given to aid understanding to convey the present invention. It is clear that these examples are for illustrative purposes only and are not intended as a framework of the present invention are intended to limit.

Each distinction in the following examples has been made using a Device carried out, on a laboratory scale, incorporating a tubular stainless steel reactor and conventional product recovery and analyzers used. The reaction vessel was 50.8 cm (20 inches) long and had an inside diameter of 1.58 cm (0.622 inch). The reactor temperature was measured using a Maintaining a hot molten salt bath from DuPont HITEC. The internal reactor temperatures were determined with the aid of a co axial! thermocouple

For the tint requests related to hydrodenitrogenation, a catalyst charge of 15 to 16 g of granular Material used that passed through a 0.84mm (20-mesh U.S. Sieve) screen, but not one 40-mesh U.S. Sieve. The catalyst was supported in the lower third of the reactor on a layer of Pyrex glass beads 4 mm in diameter. The volume the reactor above the catalyst bed was empty. The catalyst bed Occupied about 12.7 to about 19.0 cm (5 to 7.5 inches) of the reactor length depending on the bulk density of the catalyst.

Before the hydrodenitrogenation test, each catalyst was pretreated at 87.9 atmospheres (1250 ρsig) and 260 ⁰ G (500 ⁰ F) for a period of 2 hours with a hydrogen flow of 6.2 cbm / kg (100 SCFHP) . The hydrocarbon feed was introduced at a rate that an LHSV would provide in the range of about 0.8 cc hydrocarbon per hour per cc catalyst to about 1.4 cc hydrocarbon per hour per cc catalyst and a weight hourly rate (VHSV) of 1.9 g of hydrocarbon per hour per g of catalyst. The hydrocarbon feed used in these experiments was a heavy catalytically treated circulating oil (HCOO) and had the properties listed in Table I for the

009842/1825

Feedstock # 1 is shown. Hydrogen was then added to the system once at an amount of about 5204 cbm / cbm (18,000 SCFB). The average catalyst bed temperature was about 571 ° C (70O ⁰ P). The duration of each approach was 7 days.

Tabel I. Charge no. 1 212
515
540
555
574
599 (415)
(594)
(644)
(672)
(704)
(750) 2 205
246
270
286
295
525
555 (598)
W5)
(519)
(546)
(565)
(614)
(652) ASTM distillation ⁰ C ( ⁰ JT)
Initial boiling point
10 % by volume (overhead)
50 »
50 "
70 "
90 "
End boiling point 17.1
1.5507 27.5
1.5026 Loading properties Density, ⁰ API
pn
Refractive index, η η 0.77 0.25 Sulfur, wt% 756
55
27
676 159 Nitrogen, ppm
total (a)
basic (b)
Indoles (c)
Carbazole (d) - 209 Molecular weight 16.9
54.5
48.8 25.5
42.2
54.5 Hydrocarbon type vol-% (e)
Paraffins
Naphthenes
Aromatics

(a) Kjeldahl method

(b) Perchloric acid titration

(c) Spectrophotometric method

(d) by difference

(e) by mass spectrometric analysis

009842/1825

The nitrogen content of a product sample s during the six- & test day were obtained and a product sample that was obtained during the seventh-day of testing were determined. Each sample, a two hour sample of the entire product, was first purged with flowing hydrogen at atmospheric pressure for about 5 minutes to remove hydrogen sulfide and ammonia. It was then subjected to coulometric titration for nitrogen determination. Alternatively, one could analyze each sample for total nitrogen using the Kjeldahl method. The arithmetic mean of the nitrogen values for the two samples was taken as the nitrogen content of the product obtained with each catalyst.

A relative hydrodenitrogenation activity was determined for each catalyst used in each hydrodenitrogenation test. This was accomplished using Figure 2, which depicts a relationship between the nitrogen content of the product of the test and the catalyst activity. The data for this relationship was obtained with a commercially made sulfurized nickel-tungsten-on-fluorinated silica / alumina catalyst. This catalyst, which was obtained from Harshaw Chemical Company (Harshaw Catalyst ΝΪ-4401Ε), contained 5.6% by weight nickel, 14.0% by weight tungsten, 6.1% by weight sulfur, 1.4% by weight % Fluorine and 14.6% by weight aluminum oxide. It had a surface area of 176 square meters / g and was obtained as a 0.254 cm (1/10 inch) extrudate. The study was conducted at an average catalyst bed temperature of about 571 ^° C (700 ^° F), a variable weight throughput (WHSV), a hydrogen addition rate of about 3204 cbm / cbm (18,000 SCFB), and a total pressure of 87.9 atm (1250 psig). A value of 100 was taken as the relative activity of the catalyst for the test run at a V / HSV 1.9 g feed per hour per g catalyst. The results obtained after the sixth and seventh day of the hydrodenitrogenation test were arithmetically averaged, ie the nitrogen contents were averaged and the average catalyst bed temperatures were averaged.

The hydrodenitrogenation activity was determined by obtaining

009842/1825

- * r>

ORIGINAL

of the activity value from FIG. 2, which corresponds to the corresponding nitrogen content of the product, which was determined from product samples from the sixth and seventh day in the course of the test. This value is equivalent to one hundred times ratio of the activity of the catalyst was studied of this test in the average catalyst bed temperature to the activity of the comparative catalyst for the average catalyst bed temperature of $ 71 ⁰ G (70o ⁰ P). This relative activity value has been corrected to an average catalyst bed temperature of 371 ⁰ C (700 ⁰ P) using the following equation:

A - A _o e -

1 1
T "T

, whereby

A = the activity of the investigated catalyst in the

Test temperature,
A ₀ = the activity of the catalyst, which at 371 ⁰ C (700 ° P)

was investigated,
T = temperature of the test in ⁰ K
T _o = 644 ° K (371 ° C (700 ⁰ P))
Δ Ε = 11,000 calories per gram-mole and E = 1.987 calories per gram-mole per ⁰ K.

In each of those studies aimed at hydrocracking a catalyst charge of 19 g of a granulating material was used, which was passed through a sieve with a mesh size of 1.68 mm (20-mesh U.S. Sieve) passed but not through a 0.84 mm (20-mesh U.S. Sieve) screen. The catalyst was supported in the lower third of the reactor by a layer of 4 mm thick Pyrex glass beads. The volume of the reactor above the catalyst bed was empty. The catalyst bed filled about 15 by 3 to about 22.9 cm (6 to about 9 inches) of the length of the reactor depending on the bulk density of the catalyst.

Before this Hydrocrackungstest each catalyst was 87.9 atm (1250 psig) at 260 ⁰ C (500 ⁰ P) during a time period of 16 to 20 hours with hydrogen at a Pluss of about 3.1 cubic meters / kg

009842/1825 bad orig _{/ NA} l

SCFHP) pretreated. The hydrocarbon feed was started at 260 ° C (500 ° F) and raised the temperature over a period of several hours until the desired conversion sion level was reached. Thereafter, the temperature was adjusted to maintain about a 77 wt% conversion. Other Process conditions included a total pressure: of 87> 9 atm (1250 psig), a fluid flow rate in the range of about 0.7 ecm of carbon hydrogen per hour per ecm of catalyst up to about 1.0 ecm of hydrocarbon per hour per ecm of catalyst, a weight throughput of about 1.38 grams of hydrocarbon per hour per gram of catalyst and a hydrogen feed rate of about 2136 obm per cbm (12,000 SCFB). The hydrocarbon feed, that was used was a low sulfur mixture from LVGO and LCCO and is hereinafter referred to as loading material No. 2 designated. Its properties are given in Table I.

For each By & rocracking test, data was obtained from 1 to 13 days during the execution of the procedure. Weight determinations were obtained on two hour samples taken at intervals of at least 24 hours. Product recoveries were generally in excess of 99 weight percent based on the weight of the hydrocarbon feed. Gas and liquid analyzes were combined and normalized to 100% to obtain the conversion rate. The product distributions were calculated to be 103 weight percent total based on the hydrocarbon feed to compensate for hydrogen consumption.

As used herein, the term conversion is defined as that portion of the total reactor effluent that is both liquid or is gaseous boiling below a true boiling point of 193 ö (38O ° F). This percentage was determined by gas chromatography. The hydrocarbon product was taken in the form of analysis samples at intervals of not less than 24 hours. The collection period was 2 hours during this time the liquid product was collected in a dry ice acetone condenser, "to avoid the condensation of pentanes and heavy hydrocarbons to ensure. During this time the water

009842/1825

Substance-rich exhaust gas is collected in samples and immediately analyzed for light hydrocarbons using isothermal gas chromatography. The liquid product was weighed and analyzed using a two-column temperaturprogramniierten gas chromatograph (χ 6 feet 1/4 inch) 1.83 cm 0.635 m χ equipped columns / SF 96 on firebrick and thermoresistance detectors. Individual compounds were determined using methylcyclopentane. The minimum in the chromatographs just before the n-undecane peak was designated as the 193 ° C (380 ⁰ P) point. The distance between the light and the heavy naphtha (82.2 ⁰ C - 180 ⁰ F) was arbitrarily selected as a specific minimum within the co-paraffin-naphthene group, which agreed with the fraction obtained by a Oldershaw least Ilia ion of the product was obtained.

The temperature requirements for the 77 % conversion were calculated from the observed data using zero order kinetics and an activation energy of 35 kilocalories. Adjustment of the temperature requirements was also made to a constant hydrogen to oil ratio of 2136 cbm / cbm (12,000 SCFB) using the following equation:

where R is the gas flow rates in 178 cbm / cbm (1000 SCFB).

The temperature required for 77 % conversion was chosen as a means of expressing the hydrocracking activity of the catalyst under study. In order to eliminate irregular values that could occur at the beginning of the run, an estimated value was obtained for the temperature required for the 77% conversion after 7 days while the process was being carried out. In order to estimate these values, a curve was made with the temperature necessary for the 77% conversion plotted on the ordinate and the test days were plotted on the abscissa, then the value of the temperature after 7 days was plotted Execution of the procedure read from the smooth curve of this drawing_. This latter value was used to determine the activity of the catalyst used in the

009842/1825

investigation was used from which the recorded data was obtained were to determine.

The relative hydrocracking activity was obtained using the following equation:

/ * s E

A = 100e R

, whereby

A = the relative activity of the catalyst studied. E = 35,000 calories per gram-mole

R = 1.987 calories per gram-mole per ⁰ K T - the temperature of the Testos on the seventh day in ⁰ K, and T ₀ = 646 ⁰ K.

The yield of heavy naphtha, ie the yield of product formed between 82.2 ⁰ C (180 ° ^ and 193 ° C (38O ° F) boiling was corrected with respect to the temperature and conversion. The following equation was used to calculate the heavy naphtha yield under normal conditions of 385 ° C (725 ⁰ P) and 77% conversion.

H ₀ = H ₊ 15 χ 10 ⁴ (^ I ₅ - ^ f ^) ₊ 9 x 10 ⁸ ^ ^ ή

where H ₀ = heavy haphtha yield at 385 ° C (725 ° F) and 77 wt%,

iger conversion,

H = the observed heavy haphtha yield in% by weight, T = the observed temperature in ⁰ ° C., and C = the observed conversion in% by weight.

The heavy naphtha yield was used to determine the catalyst selectivity to express.

example 1

Various catalysts were prepared in this example and on their hydrodenitrogenation activities, hydrocracking

009842/1825

activities and their hydrocracking selectivities.

Unless otherwise specified, each catalyst was prepared by impregnating the hydrogenation metals identified in detail into a catalyst support comprising 35 weight percent ultra-stable crystalline aluminosilicate material suspended in a matrix of low alumina silica / aluminum oxide. This catalyst support was not a physically discrete mixture of aluminosilicate material and silica / alumina particles. Rather, the carrier was made by mixing finely divided, ultra-stable, large-pore crystalline aluminosilicate material in a sol or gel of the silica / alumina before drying and / or calcining. The hydrogenation metals were deposited on the catalyst base, either alone or in various combinations of solutions of heat decomposable salts. In each case, the catalyst base was impregnated with a pin aqueous solution of the desired metal salts, but only enough solution was used to fill the pore volumes of the catalyst support material. If necessary, the pH of the solution was adjusted to avoid precipitation. Each catalyst composition was then per a period of time ranging from about 1 to 2 hours at a temperature of about 121 ⁰ C (25O ⁰ F) dried in air, with an amount of about 0.0425 m³ (1.5 cubic feet) 1 hour flowed and formed into pellets 0.635 x 0.635 ca (1/4 "χ 1/4") in size with Sterotex (at about 4% by weight) and held at a temperature of about 3 to 4 hours for a period of time ranging from about 3 to 4 hours calcined about 538 ^° C (1000 ^° F) in flowing air at a rate of about 0.0425 cbm (1.5 cubic feet) per hour.

In the case of a catalyst A, a commercially prepared one was used Catalyst selected. The base or support of this catalyst was a low silica / alumina material Aluminum oxide content (about 13% by weight of aluminum oxide), which contained 35% by weight of ultra-stable, large-pore crystalline aluminosilicate material suspended in its porous matrix. Cobalt and molybdenum were in this catalyst support with the help of a solution of cobalb acetate or a solution: of ammonium molybdate,

ΟΟ9842 / 18125

BATH ORIGINAL

leads. It was found that the catalyst A contained 2.52% by weight of CoO and 9 »46% by weight of MoO ₅ . , _§

As part of the catalyst B, 94-6 catalyst carrier materials were impregnated with a solution which was prepared by adding a solution of 9%. g of nickel nitrate, dissolved in 50 ml of hot distilled water (about 71 ° C (160 ° F)), with a solution of 12.1 g of ammonium molybdate in 50 μl of hot distilled water (about 71 ° 0 (16O ⁰ J ¹ )) combined. It was found that the catalyst B contained 2.5% by weight NiO and 10% by weight MoO ₅ .

3dn the case of catalyst C was added 90 g of the catalyst support material iatprügniert with a solution of 8.4 g cobaltous acetate and 10.9 g Aamoniiam-metatungstate dissolved in 110ml of hot distilled water (about 71 ⁰ C (160 ° F)). Catalyst C was prepared to contain 2.7 wt% CoO and 10.6 wt% WO.

In the case of catalyst D, 97 g of the catalyst support material β were impregnated with a solution which was prepared by dissolving 10.9 g of ammonium aeta-tungstate and 9.7 g of nickel-II nitrate in 110 ml of hot distilled water (approx. 71 ⁰ C (160 ⁰ F)). Catalyst D was prepared to contain 2.5 wt% NiO and 10.1 wt% wOx.

In the case of catalyst E, the ultrastable, large-pore crystalline aluminosilicate material was first cation-exchanged in a finely divided state in order to reduce its sodium content to 2.20 wt at 90 ⁰ O (19A- ⁰ F). The solution was prepared by dissolving 157B ammonium sulfate in 1.5 liters of distilled water. The contact was made with stirring. The cation-exchanged aluminosilicate material was filtered and washed with about 3.78 liters (1 gallon) of hot distilled water (about 71 ^° C) (160 ^° F) in 500 ml portions. The cation exchange procedure was repeated twice. After the final exchange, the aluminosilicate material was washed free of sulfate anions, dried in air, at a temperature of about 121 ^° C (250 ^° F) for about two

0098A 2 / T 825

Hours at an air flow rate of about 0.0425 Cbin per hour and for two hours at a temperature of 81O ⁰ C (149O ⁰ JF) calcined in flowing air at a flow rate of about 0.0425 cubic meters per hour (1.5 cubic feet per hour).

Then 65 g of the aluminosilicate material, which contained only 0.65% by weight of sodium, were impregnated with a solution which was prepared by adding 0.57 B of diaminopalladium nitrite (57% by weight of palladium) in 150 ml of distilled water which contained 2 drops of concentrated nitric acid. The catalyst mass was then for 2 hours at a temperature of 121 ⁰ C (25O ⁰ F) with an air flow rate of about 0.0425 m³ (1.5 cubic feet) per hour in dried pellets with 0.635 x 0.635 cm (1/4 χ 1/4 inch) to about 4 wt .-% Sterotex transferred and 4 hours in air at a temperature of about 538 ⁰ C (1000 ⁰ F) and an air flow rate of about 0.0425 m³ (1.5 cubic feet) per hour calcined. The catalyst E was prepared in such a way that / contained 0.5% by weight of palladium.

Each of the catalysts made in this example, with the exception of the palladium-containing catalyst was tested for hydrodenitrogenation activity examined using the procedure described above. A relative hydrodenitrogenation activity value was determined for each catalyst examined in this way. Each catalyst was also tested for its hydrocracking activity / ^ selectivity investigated according to the procedure given above. The results of these tests are given in Table II specified.

009842/1825

Table II Catalyst performance

catalyst 'Kr. O σ A. CD OO
■ Ρ- B. ro -> * σ co D. IO σι E.

Pd CoO NiO

2.52

2.5 2.7 -

2.5 0.5

nent - Wt% activity HC EC MoO,
0 - WHERE, HUN 100 selectivity 9.46 86 152 59.1 10 - 99 115 60.5 10.6 94 165 57.9 10.1 154 115 58.2 60.6

• S3 • Ρ-

The hydrocracking (HG) selectivity is expressed as the percentage of Schwernaphthafraktion (boiling between about 82.2 ⁰ C (180 ° F) and 193 ° C (38O ⁰ P)) of the product in wt .-% of the product. The accuracy of such data from these small scale test units is H ₁ 1.5 wt%.

These results indicate that Catalyst D had superior hydrodenitrogenation activity (HDIi), superior hydrocracking activity and has a satisfactory selectivity.

Example 2

In this example, additional catalysts were prepared and tested for their corresponding hydrodenitrogenation activities, Hydrocracking Activities and Hydrocracking Selectivities. Each catalyst was prepared by adding a certain amount of the catalyst support described in Example 1 is impregnated with certain hydrogenation metals. Again, the hydrogenation metals were on the catalyst support material applied from solutions of heat-decomposable salts. The technique of impregnation used in Example 1 was used applied to this example. ' This technique involved drying and subsequent calcining as described in Example 1.

In this example three catalytic compositions were investigated. These compositions were designated as the F..G and H Catalysts.

In the case of Catalyst F, 100 g of catalyst support material was impregnated with a solution which was prepared by adding 38.8 g of nickel H nitrate and 18.4 g of ammonium meta-tungstate ini? * '* Ml of hot distilled water ( about 71 ⁰ C (160 ⁰ F)). The catalyst F was prepared so that it was 8.6 wt .-% NiO and 14, ρ

Wt .-% WO _x contained. P.

In the case of the catalyst G was added 80 g of the catalyst support material impregnated with a solution, which was prepared by reacting 31 g of cobalt nitrate and 20.9 g of ammonium molybdate (about 71 ⁰ C (160 ⁰ F)) dissolved in about 100 ml of hot distilled water .

009842/18 2 5

BATH ORIGINAL

The catalyst G was prepared so that it contained 8.1% by weight of CoO and 17.4 % by weight of MoO.

Bb case of the catalyst H was added 80 g catalyst support material impregnated with a solution which was prepared so that 31 g Hicke lnitrat and 20.9 g Aiaaoniummolybdat (about 160 ⁰ P) 71 ° C () was dissolved in about 100 mlheissem distilled water. The catalyst H was prepared so that it contained 8.1% by weight NiO and 17.4% by weight WoO.

Each of these catalysts was tested for hydrodenitrogenation activity, hydrocracking activity, and hydrocracking selectivity by the methods given above. The results of these tests are given in Table III.

009842/1825

Table III Catalyst Performance

σ catalyst ο CD as PO P. _ »
00 ro G cn

Hydrogenation component wt% activity

CoO

NOK

8.6

8.1

8.1

MoO-

17.4-

17.4

WO, HDH "HO Z

14-5 132

79

125 EC
selectivity

59.2

53.4

60.2

CD -J • S3

The catalyst F was an embodiment of the invention catalytic composition. Although the amounts of the hydrogenation components of these catalysts were larger than those in the catalysts of Example 1, the catalyst F showed at least just as effective as catalysts G and H.

Example 3

In this example, embodiments of the invention were Catalytic compositions have been studied to reduce their activity in hydrocracking a petroleum hydrocarbon feed and to compare this decline in activity with that of other hydrocracking catalysts. A portion of Catalyst A was used as the other hydrocracking catalyst while Catalysts I and J the embodiments according to the invention of the catalytic compositions represent.

In the case of Catalyst I, 194 S catalyst support material used in the previous examples was impregnated with a solution prepared by dissolving 12.4 g of nickel nitrate and 21.8 g of ammonium meta-tungstate and 260 ml of hot distilled water (approx ° C (16O ⁰ F)). The catalyst I was prepared so that it contained 1.5 wt .-% NiO and 9.5 wt .-% WO ^.

In the case of Catalyst J, 172 g of the catalyst support material used in the previous examples was impregnated with a solution prepared by dissolving 38.8 g of nickel nitrate and 36.8 g of ammonium meta-tungstate in 260 ml of hot / distilled Water (about 71 ⁰ G (160 ⁰ F)). This catalyst was prepared to contain 4.7 wt% NiO and 16.0 wt%

WO-, contained. 3

Each of these catalysts was in the unit described above on a small scale 'investigated. The operating conditions that given above were used; H. a total pressure of 87.9 atmospheres (1250 psig), a weight flow of 1.38 grams of hydrocarbon per hour per gram of catalyst, a hydrogen addition rate ranging from 1958 cbm / cbm (11,000 SCFB) to about 2436 cbm / cbm

009842/1825

BATH ORIGINAL

SGFB) and a temperature that would result in a conversion, which is close to 77% by weight. Feedstock # 2 was used. The activity was expressed in each PaI-I as the temperature that was required for the 77 weight

% conversion. The data obtained from these tests are in of Figure 3 indicated.

Referring to Figure 3, both embodiments show the catalytic compositions according to the invention, d. H. the catalysts I and J, an activity related to the activity of the other hydrocracking catalyst, i.e. H. the catalyst A, were superior. The embodiments according to the invention catalytic composition required a much lower temperature to achieve 77 wt% conversion. There was also a drop in activity these two embodiments of the invention catalytic compositions, as demonstrated by the slope of the corresponding curves, are significantly lower than the decrease in activity of the other hydrocracking catalyst.

Example 4

In this example a number of embodiments of the catalytic compositions according to the invention were examined in order to determine the optimal amounts of the hydrogenation component. Each of the catalysts was prepared by impregnating a certain amount of the catalyst support material used in the previous examples with a particular solution, as the specified amounts of nickel-2-nitrate (25.8% by weight NiO) and ammonium meta-tungstate (92,2Gew .-% W0 ^) contained, which were dissolved in a certain amount of hot distilled water (about 71 ° C (160 ⁰ F)).

Catalysts K through Z and AA through HH were used in this example and were prepared using the method described above the tightness of connections and Kafcalysatortträgermaterial, -like it is given in Table IV, and to create the amounts of NiO and WO, also given in Table IV are.

BATH ORIGINAL

009842/1825

Table IV ^ "" Zj.'2 4 13 520 Cat carrier
(free from flee
gen components) Cat
te . -Compon
(Wt%) Catalyst Manufacture 10.9 (ml) (6) NOK WHERE,
t> r _n 4- ! ^ bindings (g)
WH ( ΪΪΩ ^ fvm Well 10.9 150 90 0.0 10.0 D. Clv » JSX \ a \ J-rJ ρ 57.9 110 86.4 2.5 10.2 K 37.0 200 177.5 5.0 15.8 Il 9.7 18.5 190 165.7 2.7 16.6 M. 45.0 18.5 190 174.6 2.8 . 8.7 If 21.8 13.7 250 174.5 4.2 8.5 O 21.8 27.9 250 183.5 1.4 6.3 P. 52 * 2 13.7 250 172.5 1.4 12.8 Q 10.9 36.8 250 183.5 2.1 6.3 S. 10.9 21.8 260 194.0 4.2 16.4 S. 16.5 21.8 260 183.5 3.0 9.6 JD 58.8 36.8 260 188.1 1.5 9.5 U 24.8 56.8 260 167.0 4.7 16.0 V 12.4 18.4 260 170.6 2.4 16.2 W. 58.8 18.4 150 90.0 8.6 14.5 X 19.4 57.0 125 79.0 5.0 16.8 ϊ. 58.8 55.9 200 162.5 1.6 17.1 Z 19.4 58.2 200 146.0 . 1.6 25.7 IJL 12.4 54.7 200 139.5 5.4 26.8 SB 12.4 15.0 250 161.0 1.0 16.4 UO 26.4 47.8 250 180.4 1.0 6.2 BD 7.8 69.4 250 148.0 1.6 22.6 IT 7.8 28.2 220 128.7 1.6 32.7 ff 12.4 250 165.7 4.0 13.0 00 12.4 HH 51.0

009842/1825

Each of these catalysts was tested for its hydrocracking selectivity and activity of the laboratory unit described above. The normal operating conditions for these tests were as follows: a pressure of 87.9 atmospheres (1250 psig), a weight throughput of 1.38 grams of hydrocarbon per hour per gram of catalyst; a hydrogen addition rate of 19-58 cbm / cbm (11,000 SCFB) to about -24-72 cbm / cbm (13,200 SCFB); and a temperature that provided a conversion equivalent to about 77 weight percent conversion. The hydrocracking and HydroerackungsSelektivität, expressed as% by weight of heavy naphtha (boiling between 82.2 ⁰ C (18O ⁰ F) and 193 ° C (380 ⁰ F)) in the product are available for each of these catalysts in the Table V given.

009842/1825

Table V Catalyst Performance

Hydrocracking;

catalyst activity selectivity K 17th 64.2 L. 167 58.2 M. 167 62.6 Ή 210 61.8 O 149 60.8 P. 140 61.0 Q 152 61.1 E. 220 63.9 S. 149 60.2 T 172 62.4 U 164 61.9 V 205 61.3 W. 172 60.8 X 227 59.7 Y 110 59.2 Z 160 59.2 AA 239 60.3 BB 195 62.0 CC 205 58.5 DD 227 62.4 EE 164 60.9 FF 200 60.1 GG 160 61.0 HH 180 59.8

009842/1825

The hydrocracking activity is the activity of the catalyst during the seventh day. The Hydrocrackungsselektivität is an arithmetic mean of the Schwernaphthaausbeuten, the strength at 77% conversion and a temperature of 385 ⁰ C (725 ⁰ P) have been corrected.

The values contained in this example and given in Tables IV and V above were used to prepare the activity contour plot shown in FIG. This activity contour diagram shows the response of the activity to the catalyst composition and was constructed by taking as the ordinates the amount of nickel in terms of gram-moles of nickel oxide (NiO) per 100 g of catalyst and as the abscissa the amount of tungsten in terms of gram-moles of tungsten oxide (WO,) used per 100 g of catalyst. The NiO and WO4 concentrations are given in terms of gram-moles per 100 grams of catalyst to give an indication of surface coverage. Starting from the origin, lines of constant atomic ratio of the hydrogenation metals are drawn. The line shown corresponds to an atomic ratio of 3 atoms of tungsten per atom of nickel. The activity contours shown in Figure 4 range from 1.2 to 2.4. Since the activity is based on the temperature requirements for constant conversion, the activity contours in fact isotherms (696 ⁰ F) range from 369 ° C at 1.2 activity to 353 ° C (668 ⁰ F) at 2.4 Activity . For comparison, the catalyst A used in Example 1 had a relative activity of 100 and required a temperature of 372 ⁰ C (704 ° F) under the testing conditions used. It is determined by Figure 4 that the most active hydrocracking catalyst composition is one containing about 19 weight percent metal oxides with a tungsten to nickel atomic ratio of about three. The above activity contour plot provides much useful information relating to the effective amounts of the metals of the hydrogenation components of the catalytic composition of the present invention. For example, the diagram indicates the broad advantageous preferred ranges of these metals. The graph shows that one embodiment of the catalyst should have a minimum efficiency of about 140 when nickel and tungsten are present in a total amount ranging from about 0.03 to about 0.20 gram-moles NiO

009842/1825 bad o _R , g, nal

and WO, per 100 g of catalyst and at a tungsten to nickel ratio in the range of about 0.6 to about 12.0; a minimum activity of about 180 when nickel and tungsten are present in a total amount ranging from about 0.05 to about 0.18 gram-moles HiO and WO, per 100 grams of catalyst and with a tungsten to nickel ratio in the range of about 1.2 to about 8.0 and one minimum activity of about 220 when nickel and tungsten are present in a total amount ranging from about 0.07 to about 0.13 Graaa moles NiO and WO, per 100 g of catalyst and for one Tungsten to Hickel ratio ranging from about 2 to about 5 · Das

Diagram suggests »that the optimal tungsten to nickel ratio 5.0.

It is clear and has been sufficiently demonstrated by the foregoing Examples that the inventive catalytic composition and certain hydrocarbon conversion processes, who use this catalytic composition, in particular one hydrodenitrogenation process and a hydrocracking process, are superior to other catalytic compositions and methods using these catalytic compositions.

BATH ORIGINAL

009842/1825

Claims (1)

Claims 1. Catalytic composition for the conversion of petroleum hydrocarbons, characterized in that the catalytic Composition a hydrogenation component and a cocatalytic one acidic cracking carrier, wherein the hydrogenation component comprises a member of the following group that consists of 1.) elemental nickel and elemental tungsten, 2.) their oxides, 30 their sulfides and 40 mixtures thereof, the Carrier comprises large pore crystalline aluminosilicate material and a refractory inorganic oxide. 2. Use of the catalytic composition according to Claim 1 for the conversion of petroleum hydrocarbons, characterized characterized in that one converts the hydrocarbons into a hydrocarbon Konyersionzone under hydrocarbon conversion conditions and contacted with the catalytic composition in the presence of hydrogen. 3. Catalytic composition according to claim 1, characterized in that that the large pore crystalline aluminosilicate material is ultra stable large pore crystalline aluminosilicate material is characterized by a maximum cubic unit cell dimension of 24.55 ^ -units, by hydroxylin- —Ί —1 infrared bands around 3700 cm and around 3625 cm and due to an alkali metal content less than 1%, and where the refractory inorganic oxide is silica / alumina. 4. Catalytic composition according to claim J, characterized in that that the ultra stable large pore crystalline aluminosilicate material is suspended in a matrix of silicon dioxide / aluminum oxide, this carrier being produced ind3in the ultra-stable, large-pored crystalline aluminosilicate material in a finely divided state with a sol or gel of silicon dioxide / silicon oxide mixed and then dries. 5 · Catalytic composition according to claim 3, characterized characterized in that nickel and tungsten are present in a total amount ranging from about 0.03 to about 0.20 gram-moles of KiO 009842/1825 ORIGIMAL BATHROOM and VO, per 100 g of catalyst and at a tungsten to nickel ratio ranging from about 0.6 to about 12.0. 6. - Catalytic composition according to claim J, characterized in that that nickel and tungsten are present in a total amount ranging from about 0.05 to about 0.18 gram-moles NiO and WO, per 100 g of catalyst and in a tungsten to nickel ratio ranging from about 1.2 to about 8.0. 7. Catalytic composition according to claim 3? characterized, that nickel and tungsten are present in a total amount ranging from about 0.07 to about 0.13 gram-moles NiO and WO, per 100 g of catalyst and in a tungsten to nickel ratio in the range from about 2.0 to about 5 * 0. 8. Use of the catalytic composition according to claim 3 for the conversion of mineral oil hydrocarbons, characterized in that that these hydrocarbons are in a hydrocarbon conversion zone under hydrocarbon conversion conditions and contacted with the catalytic composition in the presence of hydrogen. 9. Catalytic composition according to claim 4, characterized in that that nickel and tungsten are present in a total amount ranging from about 0.03 to about 0.20 gram-moles of NiO and WO, per 100 g of catalyst and in a tungsten to nickel ratio ranging from about 0.6 to about 12.0. 10. Catalytic composition according to claim 4, characterized in that that nickel and tungsten are present in a total amount ranging from about 0.05 to about 0.18 gram-moles NiO and WO, per 100 g of catalyst and in a tungsten to nickel ratio ranging from about 1.2 to about 8.0. 11. Catalytic composition according to claim 4, characterized in that that nickel and tungsten are present in a total amount ranging from about 0.07 to about 0.13 gram-moles NiO and WO, per 100 g of catalyst and in a tungsten to nickel ratio in the range of about 2.0 to about 50. 009842/1825 12. Use of the catalytic composition according to Claim 4 for the conversion of petroleum hydrocarbons, characterized that these hydrocarbons are in a hydrocarbon conversion zone under hydrocarbon conversion conditions contacted with the catalytic composition in the presence of hydrogen. . 13. Use of the catalytic composition according to claim 4 for the hydrocracking of petroleum hydrocarbons, thereby characterized in that in a hydrocracking reaction zone the hydrocarbons under hydrocracking conditions and in the presence of hydrogen contacted with the catalytic composition. 14. Use of the catalytic composition according to claim 4 for hydrodenitrogenation of a petroleum hydrocarbon fraction, characterized in that in a hydrodenitrogenation reaction zone the hydrocarbon fraction under hydrodenitrogenation conditions in the presence of hydrogen contacted with the catalytic composition. 15 catalytic composition according to claim 9, characterized in that that the ultra stable large pore crystalline aluminosilicate material is present in an amount ranging from about 5 to about 70 percent by weight based on the weight of the carrier. 16. Catalytic composition according to claim 10, characterized in that the ultra-stable, large-pore crystalline aluminum silicate material is present in an amount ranging from about 5% to about 70% by weight based on the weight of the Carrier. 17 · Catalytic composition according to claim 11, characterized in that the ultra-stable, large-pored crystalline aluminosilicate material is present in an amount ranging from about 5% to about 70% by weight based on the weight of the Carrier. 18. Use of the catalytic composition according to claim 15 for the conversion of petroleum hydrocarbons, thereby characterized in that the hydrocarbons are in a hydrocarbon conversion zone under hydrocarbon conversion 009842/1825 "conditions and in the presence of hydrogen contacted with the catalytic composition,
19 * Use of the catalytic composition according to In-. Proverb 15 on the hydrocracking of petroleum hydrocarbons, man
characterized in that / in a hydrocracking reaction zone the hydrocarbons under hydrocracking conditions. and contacted with the catalytic composition in the presence of hydrogen. 20. Use of the catalytic composition according to claim 15 for the hydrodenitrogenation of a petroleum hydrocarbon st off fraction, characterized in that one in a Hydrodenitrogenation reaction zone the hydrocarbon fraction in the presence of hydrogen and under hydrodenitrogenation conditions contacted with the catalytic composition, 21. 'Use of the catalytic composition' according to claim 16 for the hydrocracking of petroleum hydrocarbons, characterized in that in a hydrocracking reaction zone the hydrocarbons under hydrocracking conditions and in the presence of hydrogen with the catalytic composition contacted. 22. Use of the catalytic composition according to claim 16 for hydrodenitrogenation of a petroleum hydrocarbon -Fraction, characterized in that in a hydrodenitrogenation reaction zone, the hydrocarbon fraction in the presence of hydrogen under hydrodenitrogenation conditions contacted with the catalytic composition. 23. Use of the catalytic composition according to claim 17 for the hydrocracking of petroleum hydrocarbons, characterized in that in a hydrocracking reaction zone the hydrocarbons under hydrocracking conditions and in the presence of hydrogen with the catalytic composition contacted. 24. Use of the catalytic composition geraäss Claim 17 for the hydrodenitrogenation of petroleum hydrocarbon fractions, characterized in that in a hy- , drodenitrogenation reaction zone the hydrocarbon fraction in the presence of hydrogen and under "hydrodenitrogen 009842/1825 conditions with the catalytic! Composition contacted, 25 · Application claim 19> characterized in accordance that the Hydrocrackungsbedingungen are the following: a total pressure in the range of about 49 atmospheres gauge (700 psig) to about 281 atm (4000 psig), an average catalyst bed temperature in the range of about 543 ⁰ C. (650 ° F) · to about 454 ⁰ C sdurchsat (850 ⁰ F) and a hydrogen to oil ratio in the range of about 890 cubic meters / cubic meter (5000 SCFB) to about 3560 cubic meters / cubic meter (20,000 SCFB), and a liquid »Range from about 0.5 to about 10 volumes of hydrocarbon per hour per volume of catalyst. 26. Use according to claim 20, characterized in that the hydrodenitrogenation conditions are the following: an average catalyst bed temperature in the range of about 204 ° C (40O ⁰ F) to about 427 ⁰ C (800 ⁰ F) and a hydrogen to oil Ratio in the range of about 890 cbm / cbm (5000 SCFB) to about 7120 cbm / cbm (40,000 SCFB), a total pressure in the range of about 35 atg (500 psig) to about 352 atg (5000 psig) and a liquid flow rate in the range from about 0.1 to about 20 volumes of hydrocarbon per hour per volume of catalyst. Use according to claim 21, characterized in that the hydrocracking conditions are: a total pressure in the range of about 49 atm (700 psig) to about 281 atm (4000 psig) an average catalyst bed temperature in the range of about 343 ° C ( 650 ⁰ F) to about 454 ⁰ C (850 ⁰ F), a hydrogen to oil ratio in the range of about 890 cubic meters / cubic meter (5OOO SCFB) to about 3560 cubic meters / cubic meter (20,000 SCFB), and a liquid flow rate in the range of about 0.5 to about 10 volumes of hydrocarbon per hour per volume of catalyst. 28. Use according to claim 22, characterized in that the Hydrodenitrogenierungsbedingungen include the following: an average catalyst bed temperature in Bereicli of about 204 ⁰ C (40ö ° F) to about 427 ° C (800 ⁰ F), a hydrogen to oil Ratio in the range of about 890 cbm / cbm (5,000 SCFB) to about 7120 cbm / cbm (40,000 SCFB), a total pressure in the range of about 009842/1825 35 atmospheres (500 psig) to about 352 atmospheres (5000 psig) and a liquid flow rate in the range of about 0.1 to about 20 volumes of hydrocarbon per hour per volume of catalyst, Use according to claim 23 characterized in that the hydrocracking conditions comprise the following conditions ^{: a} total pressure in the range of about 49 atmospheres (700 psig) to about. 281 atm (4000 pßig), an average catalyst bed temperature in the range of about 34-3 ⁰ C (650 ⁰ F) to about 454- ⁰ C (850 ⁰ F), a hydrogen to oil ratio in the range of about 890 cbm / cbm (5000 SCFB) to about 3560 cbm / cbm (20,000 SCFB) and a liquid throughput ι in the range from about 0.5 to about 10 volumes of hydrocarbon per hour per volume of catalyst. 30. Use according to claim 24, characterized in that the hydrodenitrogenation conditions. the following include, an average catalyst bed temperature in the range of about 204 ⁰ C (400 ⁰ F) to about 427 ° C (800 ⁰ F), a hydrogen to oil ratio in the range of about 890 cbm / cbm (5000 SCFB) to about 7120 cbm / cbm (40,000 SCFB), a total pressure in the range of about 35 atg (500 psig) to about 352 atg (5000 psig) and a liquid flow rate in the range of about 0.1 to about 20 volumes of hydrocarbon per hour per volume of catalyst. 009842/1825