SK7095A3 - Hydrocracking method for selective manufacture of distillate with middle boiling point and hydrocracking catalyst - Google Patents

Abstract

The selectivity of a midbarrel hydrocracking process for middle distillates is significantly increased by using a catalyst containing an amorphous silica-magnesia component in combination with a Y zeolite having a unit cell size below 24.45 angstroms and/or a water vapor sorptive capacity less than 10 weight percent at 25 C and a p/po value of 0.1. LZ-10 zeolite is a preferred zeolite for use in the catalyst.

Description

METHOD OF HYDRO-CRACKING ΝΑ SELECTIVE PRODUCTION OF DISTILLATES

MEDIUM BOILING TEMPERATURE AND HYDRO-CRACK CATALYST

Technical field

The invention relates to a hydrocracking process for the selective production of intermediate boiling distillates. The invention further relates to a catalyst highly selective for use in the process defined above.

BACKGROUND OF THE INVENTION

Oil refineries often produce desirable products such as fuel for combustion turbines, diesel, and other liquid hydrocarbons known as intermediate boiling distillates as well as lower boiling liquids such as naphtha and light petroleum by hydrocracking the hydrocarbon feedstocks derived from crude oil . The raw materials most commonly processed by hydrocracking are gas oils and heavy oils obtained from the crude oil by distillation. A typical gas oil contains a substantial proportion of hydrocarbon components having a boiling point above 371 ° C, usually at least 80% by weight having a boiling point above 371 ° C. A typical heavy gas oil has a boiling point between 315 ° C and 565 ° C.

The hydrocracking is usually accomplished by contact in a suitable reaction vessel, the gas oil or other feedstock is treated with a suitable hydrocracking catalyst at elevated temperature and elevated pressure in the presence of hydrogen to obtain a lower medium boiling product containing the hydrocarbon fraction distribution required by the refiner. Although the operating conditions in the hydrocracking reactor have some influence on the product yield, the hydrocracking catalyst is the primary you do! affecting yields. Currently, there are no high requirements for medium boiling distillates in the United States compared to gasoline requirements, although a marketing survey shows that there will be increased requirements for medium boiling distillates around 2000. For this reason, refineries in the past have focused for hydrocracking catalysts for medium boiling regions which selectively produce medium boiling fractions such as combustion turbine fuel and diesel fuel having boiling points ranging from 149 ° C to 371 ° C.

The three main characteristics to assess the suitability of a hydrocracking catalyst for the production of intermediate boiling distillates are activity, selectivity and stability. The activity can be determined by comparing the temperature at which different catalysts must be used under otherwise stable hydrocracking conditions with the same feedstock to produce a given percentage, normally about 60% by weight of the product with a boiling point below 371 ° C. The lower the active temperature for a given catalyst, the more active such a catalyst is as compared to a catalyst with a higher active temperature. The selectivity of hydrocracking catalysts for intermediate boiling distillates can be determined during said activity test and is measured as a percentage of the boiling point fraction of 371 ° C in the desired range, for example from 149 ° C to 371 ° C for diesel and from 149 ° C to 287 ° C for fuel for combustion turbines. Stability is a measure of how a catalyst maintains its activity for a long period of time in processing a hydrocarbonaceous feed under activity test conditions. Stability is generally measured in temperature change values per day to maintain 60% by volume or other given conversion value.

U.S. Patents 4,062,829 and 4,419,271 disclose two different types of high performance hydrocracking catalysts for the production of intermediate boiling distillates. The catalyst of U.S. Pat. No. 4,062,809 comprises molybdenum and / or tungsten and nickel and / or cobalt on a silica and alumina support dispersed in gamma alumina. U.S. Pat. No. 4,419,271 discloses that the catalyst of U.S. Pat. No. 4,062,809 can be improved by adding aluminosilicate zeolite to the substrate to produce a catalyst comprising molybdenum and / or tungsten and nickel and / or cobalt deposited on a mixture of aluminosilicate zeolite, preferably ultrahydrophobic from zeolite. known as zeolite LZ-10 and a dispersion of silica and alumina in a gamma alumina matrix. The presence of zeolite in this catalyst increases the activity of the catalyst without significantly affecting the selectivity.

Although the catalysts of the above patents are highly active hydrocracking catalysts for the production of intermediate boiling distillates, and this has been proven

in commercial environment are here new requirements hydrocracking catalysts with very high activity selectivity and stability for hydrocracking for production distillates with medium temperature reflux • nature invention In accordance with the invention was now with surprise found that

activity and selectivity of a hydrocracking process catalyst for producing intermediate boiling distillates containing one or more hydrogenation components and a zeolite Y selected from the group consisting of Y zeolites having a cell size of less than 24.25 angstroms and a Y zeolites having a water vapor sorption capacity of less than 10% by weight at 25 DEG C. and a p / p _o of 0.10, can be substantially improved by introducing components of the amorphous silicon dioxide and magnesium oxide into the catalyst. Under typical hydrocracking conditions, i.e. elevated temperature and elevated pressure in the presence of hydrogen, such catalysts are highly effective in converting gas oils and other hydrocarbon feedstocks into a product having a lower mean boiling point and a lower average molecular weight, the product containing a relatively high proportion of components. having a boiling range of 149 ° C to 371 ° C. The hydrocracking catalyst normally comprises one or more hydrogenation components, such as Group VIB or VIII metal components, in combination with zeolite Y, a component of amorphous silica and magnesium oxide, and a porous inorganic refractory oxide binder such as alumina. As used herein, p / p _o means the water vapor partial pressure to which zeolite Y is exposed divided by the water vapor partial pressure at 25 ° C.

Preliminary tests show that when the catalyst of the invention is used in hydrocracking to produce medium boiling distillates such as fuel for combustion turbines and diesel fuel, it surprisingly has a higher selectivity than other medium boiling distillation catalysts now commercially available for use in hydrocracking processes for producing intermediate boiling distillates. Surprisingly, these tests show that at 60% conversion, more than 86% by volume of the product having a boiling point up to 371 ° C, typically more than 88% by volume, has a boiling point between 149 ° C and 371 ° C, while more than 75% % by volume, often more than 77% by volume, has a boiling point between 149 ° C and 287 ° C.

DETAILED DESCRIPTION OF THE INVENTION

The hydrocracking process of the present invention is directed to high yields of intermediate boiling distillates such as diesel having a boiling point of from 149 ° C to 371 C and a combustion turbine fuel having a boiling point of from 149 ° C to 287 ° C. These high yields are obtained using a catalyst comprising one or more hydrogenation components in combination with zeolite Y having a cell size of less than 24.45 angstroms and / or a water vapor sorption capacity of less than 10% by weight at 25 ° C and a p / p _{o of} 0 , 10th Zeolite Y preferably has a mean silica to alumina molar ratio of less than 6.0, typically from 4.5 to 5.6. The catalyst preferably also comprises a porous inorganic refractory oxide binder.

The zeolite Y contained in the intermediate boiling hydrocracking catalyst has either (1) a cell size of less than about 24.45 angstroms, or (2) a water vapor sorption capacity of less than 10% by weight at 25 ° C and a p / p _o equal to 0.10 of less than 10% by weight, preferably less than 5% by weight. Preferred Y zeolites meet both of the above requirements. As used herein, the term zeolite Y includes crystalline zeolites having either a substantial powder X-ray diffraction pattern as described in U.S. Pat. and calcination, as will be apparent to one of ordinary skill in the art necessary to convert the zeolite into a catalytically effective and stable form. The present invention requires a zeolite Y having any or both of the above characteristics, such zeolites Y being modified zeolites Y as compared to the zeolites Y according to U.S. Patent 3,130,007.

The zeolites Y used in the catalyst of the present invention are large pore zeolites having an effective pore size greater than 7.0 angstroms. Because some of the pores of Y zeolites are relatively large, the zeolites allow the molecules a relatively free access to their internal structure. Y zeolites useful in the compositions of the present invention also have a generally low index of constraints, typically below 1.0, preferably below 0.75, and usually below 0.5.

The preferred zeolite Y used for the hydrocracking catalyst of the invention is zeolite UHP-Y, an ultrahydrophobic zeolite. The composition and properties of UHP-Y zeolites are described in U.S. Patent 4,401,556. See also British Patent Specification 2,014,970 B. UHP-Y zeolites and similar zeolites are essentially produced by a four step process wherein an alkali metal, typically sodium, zeolite Y, typically having a cell size of about 24.65 angstroms, is subjected to a four step process. exchanges of cations with ammonia cations, it is calcined in the presence of water vapor, preferably in the presence of water vapor at a pressure of at least 1.3 kPa, preferably at least 6.66 kPa, more preferably at least 66.6 kPa, and most preferably in an atmosphere consisting essentially or of steam to a cell size of from 24.40 to 24.64 angstroms was obtained, once again in the presence of a consisting size of preferably from 24.42 to undergo cation exchange with sufficient

24.62 angstroms, cations of steam in then ammonia essentially cells between 24.40

24.62 angstroms, with ammonia and

24.42 and cations of a sufficient amount of aqueous or wholly of angstroms, the calcination is carried out once more in the presence of water vapor, preferably in an atmospheric recovery atmosphere and 24.64 then again vapor, preferably between still consisting essentially of water vapor and preferably from steam only, to obtain a cell size below 24.40 angstroms, and preferably not greater than 24.35 angstroms. The first ammonia cation exchange step typically reduces the sodium content of the initial sodium zeolite Y from a value typically greater than 8.0% by weight, typically between 10 and 13% by weight, calculated as Na ₂ O, to a value ranging from 0.6 to 5 %, while the second exchange of cations with ammonia reduces the sodium content to less than 0.5% by weight, usually less than 0.3% by weight. From the production process described above, it is apparent that the UHP-Y zeolites differ from the Y zeolite described in U.S. Pat. No. 3,929,672 by the addition of a final steam calcination step, some of which are known under the Y-82 or LZY- 82 and Y-84 or LZY-84.

The UHP-Y zeolites are disclosed herein as zeolite aluminosilicates having a total silica to alumina molar ratio greater than 4.5, a substantial powder X-ray diffraction pattern of zeolite Y, a cell size of less than 24.45 angstroms, and a surface area of at least 300 m ² / g, a water vapor sorption capacity at 25 ° C and a p / po value of 0,10 of less than 10% by weight and a residual butanol test value of not more than 0,4% by weight. Preferred UHP-Y zeolites have one or more of the following: a total silica to alumina molar ratio of 4.5 to 35, a surface area of at least 350 m ² / g, a water vapor sorption capacity at 25 ° C and a p / 0.10 of less than 5% by weight. Particularly preferred UHP-Y zeolites having an overall molar ratio of silica to alumina from 4.5 to 9 and / or a water vapor sorptive capacity at 25 DEG C. and a p / p _o value of 0.10 of less than 4%. Although UHP-Y zeolites having a silica to alumina molar ratio of less than 6.0 are most preferred, UHP-Y zeolites that have been treated with mineral acid to remove aluminum and thus increase the molar ratio of silica to alumina can also be used . Such acid-treated UHP-Y zeolites are described in detail in U.S. Patent 5,047,139.

The most suitable zeolites of the invention have and preferably the most preferred zeolite LZ-10

UHP-Y for use in cells less than 24.40 Zeolite not more than 24.35 angstroms.

zeolite UHP-Y, zeolite LY-10 typically has a cell size of preferably between 24.20 and between 24.25 and 24.35 angstroms and angstroms, and most preferably the water vapor capacity at the partial pressure of water presented by the angstroms LZ-10 is available from UOP . at or above 24.20 24.40 the steam of angstroms has a sorption 4.6 mm at ´C of less than 8.0% by weight of zeolite. See U.S. Patent 4,419,271.

As mentioned above, the zeolites Y used in the catalyst of the invention are typically produced by a process that involves a dual exchange of cations with ammonia to reduce the sodium or other alkali metal content of the starting zeolite Y to less than 0.5% by weight of sodium, usually less than about 0.3% by weight, calculated as Na ₂ O. These zeolites of reduced sodium content possess catalytic cracking activity and can be used as components of hydrocracking catalysts. The selectivity of these Y zeolites for the production of medium boiling distillates can be substantially increased by exchanging the cations of reduced sodium content of Y zeolites with rare earth cations.

If such an ion exchange is desired, the rare earth metals selected for the exchange of cations with zeolites may be lanthanides or combinations thereof having atomic numbers in the Periodic Table of the Elements 57 to 71. Thus, the metals suitable for ion exchange are lanthanum, cerium, prazodium, neodymium, samarium , europium, gadolinium, terbium, dysprosium, holmium, erbium, tiumium, yterbium and lutium. In a most preferred embodiment of the invention, the rare earth cation mixture is introduced into the zeolite, a rare earth metal containing mixture in a distribution similar to that in rare earth ores such as bastnazite, monazite, xenotime and the like from which the metals were obtained.

There are known ways in which rare earth cations can be exchanged for sodium and other cations, especially hydrogen ions, in crystalline aluminosilicate zeolite Y. The most common method is to contact zeolite with an aqueous solution containing multivalent cations of the rare earth element (s) to be exchanged. introduced into the zeolite. Most often, the solution contains more than 20g per liter of rare earth cations, calculated as RE ₂ O ₃ , where RE is the sum of all rare earth metals considered, regardless of whether one or more of these metals actually constitute a trioxide or compound of equivalent formula and contact usually by immersing the zeolite in an ion exchange solution and stirring at or above room temperature, but not higher than 100 ° C. If desired, the solution may also contain ammonia ions and the solution may further comprise: some of the anions that do not interfere with cation exchange, such as chloride, nitrate, sulfate, etc.

For best results, the ion exchange is carried out such that the rare earth exchanged zeolite contains at least 1%, preferably 2%, and usually 4 to 6% by weight of the rare earth metals, calculated as RE ₂ O ₃ . Although the small amount of rare earth metals introduced by the zeolite exchange will replace residual sodium ions at the zeolite exchange sites, the largest proportion will be exchanged for hydrogen ions and / or ammonia ions with respect to their high concentration versus low concentration, usually below 0.3% by weight, calculated as N ₂ O, sodium cations present. Sometimes a single immersion of the zeolite into the exchange solution will be sufficient for the necessary exchange. However, in some cases, it may prove necessary to exchange by immersing it in a solution containing rare earth cations or gradually immersing in several solutions containing various rare earth elements, or using other methods to introduce the rare earth cations into the zeolite.

The zeolite Y to be used in the catalyst according to the invention, whether or not it has been subjected to ion exchange with rare earth cations, is compounded with a binder material to form particles, the substrate which serves to support one or more hydrogenation components. Preferably, this is achieved by combining zeolite with exchanged rare earth ions (1) with a material such as alumina hydrogel or peptized Catapal alumina, which upon calcination yields a porous inorganic refractory oxide binder, or (2) with a material itself is a porous inorganic refractory silica silica and clay oxide such as magnesium,

Sufficient binder, for example alumina, alumina mixture, silica-kaolin mixture, as well as combinations of these materials, zeolite Y is normally used, up to 35% by weight, preferably from preferably 5 to 10% rare earth cations.

The integration of the oxide binder of the oxide binder contained from 2 wt.% and Y with the exchanged physical refractory to the carrier 3 to 20 wt.% zeolite. and the size, shamrock or quatrefoil of extruded material to a suitable 19.0 mm, by drying the extrudates at a temperature of 426 ° C to 649 ° C for use in hydrocracking. , by crushing or cutting a length, for example from 1.5875 to subsequent calcination, to produce a material suitable for high temperature reactions, requires that the substrate be cylindrical in shape. However, as noted above, other cross-sectional shapes, such as the shamrock or quatrefoil shape shown in FIG. 8 and FIG. 10 in U.S. Patent 4,028,227.

It is clear from the foregoing description that a porous inorganic refractory oxide is used as a binder to maintain zeolite particles with rare earth cations in the substrate. Other materials, such as amorphous diluents, which have some type of catalytic activity may have been mixed in the mixture. It has now been discovered by the invention that the Y-containing zeolites described above may not be substantially increased with the present catalysts by incorporating the amorphous silica component. or

Accordingly selectivity and magnesium oxide to the catalysts. Tests have shown that the use of such an amorphous component gives much better fuel selectivity for combustion turbines and diesel than other amorphous components such as alumina, silica-alumina and gamma alumina dispersion of silica-alumina mixture.

The silica and magnesium oxide component used in the catalyst of the invention may be prepared by generally known methods. According to one such method, a silica hydrogel is first prepared by precipitating the silica from an aqueous solution of a silicate salt such as an alkali or alkaline earth silicate or silicate, by lowering the pH of the solution, usually by adding inorganic acids such as hydrochloric, nitric and sulfuric acid. Thereafter, the resulting silica hydrogel is washed and soaked in a magnesium salt solution in proportions sufficient to provide the desired relative amounts of magnesium and silica in the resulting mixture of silica and magnesium oxide. Examples of water-soluble magnesium salts that can be used to form magnesium-containing solutions are magnesium chloride, magnesium nitrate, magnesium sulfate, magnesium acetate, magnesium bromide, magnesium iodide, and the like. After the silica hydrogel has been combined with the magnesium salt solution, the magnesium oxide is precipitated in situ by the addition of a base such as ammonia, ammonium hydroxide and sodium hydroxide to increase the pH of the solution.

The resulting hydrogel is then washed with water to remove soluble salts, dried and thermally activated by calcination at a temperature above 315 ° C to form the active ingredient of silica and magnesium oxide.

In some cases, the oxide and magnesium oxide component may be formed by mixing the oxide with the dried hydrogel of the oxide, mixing more efficiently and magnesium oxide if the water is coherent to form and calcined to form magnesium.

silica. When the magnesium magnesium is used, the silica co-treatment with sufficient amounts of paste present. The paste is then a silica component and dried with an oxide

The amount of silica and magnesium oxide in the silica and magnesium oxide component may vary considerably, but at least 5% by weight of magnesium oxide, based on the total weight, is usually present. mixtures of silica and magnesium oxide. The silica-magnesium component preferably contains from 65 to 75% by weight of silica and from 25 to 35% by weight of magnesium oxide.

It is further understood that the production of a substrate for a catalyst in the form of prints is not the only way possible for those skilled in the art. The catalyst support can also be made in the form of tablets, grains, beads and pellets as desired, in some known manner, for combining zeolites with porous inorganic refractory oxide components. Regardless of how the substrate particles are made, they typically contain from 5 to 40, preferably from 10 to 25% by weight of binder and from 25 to 90, usually from 40 to 80% by weight of amorphous silica and magnesium oxide.

After production of the catalyst support particles, these are converted into catalyst particles by compounding, for example, by impregnating the particles with one or more precursors of at least one catalytically active hydrogenation metal. The impregnation can be carried out by a method known in the art in which a solution containing hydrogenated metal precursors in dissolved form is sprayed onto the substrate particles. Another method involves wetting the substrate particles in a large volume of impregnating solution. Another method is a pore saturation technique in which the substrate particles are introduced into an impregnating solution of a volume just sufficient to fill the pores of the substrate. Occasionally, the saturation technique may be modified to use an impregnating solution having a volume 10% smaller or greater than the volume that would just fill the pores. If the active metal precursors are introduced by impregnation, subsequent or second calcination at, for example, 371 ° C to 649 ° C converts the metals into their oxides. In some cases, calcination may follow each impregnation with the individual active metals.

Alternative methods of introducing the active metals into the catalyst support include (1) mixing a suitable solid or liquid material containing the metal components with the materials prior to extrusion, and (2) impregnating the materials to be extruded with the desired metal components before extrusion. These methods can be cheaper and more convenient than the impregnation methods described above and also result in a thorough mixing of the active hydrogenation components with the components of the support.

The hydrogenation components suitable for introduction into the extruded catalyst support comprise metals selected from Group VIII and / or Group VIB of the Periodic Table of the Elements. The term used

The Periodic Table in the Handbook published in 1984 Elements means the version published by Chemistry and Physics, 65th edition by the Chemical Rubber Company,

Cleveland, Ohio. Preferred hydrogenation components include metals from the group consisting of platinum, palladium, cobalt, nickel, tungsten, chromium and molybdenum. The catalyst preferably comprises at least one Group VIII metal and at least one Group VIB metal, with cobalt or nickel and molybdenum or tungsten being the preferred combinations of active ingredients and nickel and tungsten being most suitable. The catalyst typically contains up to 15%, usually from 1 to 10% by weight, preferably from 2 to 8% by weight of Group VIII non-noble metal, calculated as a monoxide and up to 30% usually from 2 to 28% by weight, preferably from 10 to 25%. % of Group VIB metal calculated as trioxide. If the hydrogenation component contains a noble metal such as platinum or palladium, it is generally required that the catalyst comprises from 0.2 to 10% by weight, preferably from 0.30 to 2.0% by weight, calculated as metal.

The hydrogenated metal catalysts present in the form of oxides are prepared as particles by the above methods or equivalent. The completed hydrocracking catalysts typically have a surface area of from 100 to 350 m @ 2 / g. When used for the production of intermediate boiling distillates, the catalysts typically contain (1) from 2 to 25% by weight, preferably from 3 to 15% by weight, and most preferably from 4 to 8% by weight of zeolite Y having a cell size of less than about 24.45 angstroms; / or a sorptive capacity for water of less than 10% at 25 DEG C. and a p / p _o equal to 0.1, (2) 4-35% by weight of a porous refractory oxide binder, preferably 7-18 wt%, (3) from 15 to 65% by weight of a mixture of amorphous silica and magnesium oxide, preferably from 30 to 70% by weight, most preferably from 50 to 65% by weight, (4) from 2 to 28% by weight of hydrogenation group VIB metal and (5) from 0 2 to 15% by weight of a Group VIII hydrogenation metal, preferably from 2 to 8% by weight.

Catalysts prepared in the form of an oxide as described above are usually converted into hydrogen sulphide for hydrocracking purposes. This may be accomplished by converting the catalyst to sulfide prior to use at high temperature, for example from 149 ° C to 371 ° C, with a mixture consisting, for example, of 10% v / v and 90% v / v hydrogen. The catalyst can be converted to sulphide ex situ in various ways. As an example, Sulficat R may be mentioned: Off-Site Presulfiding of Hydroprocessing Catalysts from Eurecat by J. H. Wilson and G. Berrebi, Catalysis 87. Studies in Surface Science and Catalysts, Elsevier Science Publishers B.V., 1988, p. Alternatively, the conversion to sulfide is performed in situ, for example, using an oxide catalyst for hydrocracking a hydrocarbon feedstock containing sulfur compounds under hydrocracking conditions including elevated temperature and pressure and the presence of hydrogen.

The catalysts described above are useful in converting a large range of hydrocarbon feedstocks by hydrocracking into more valuable lower boiling and lower molecular weight hydrocarbon products, typically having a boiling point of from 149 ° C to 371 ° C. The raw materials to be subjected to hydrocracking according to the process of the invention are mineral oils and synthetic oils such as shale oil, tar sands oil, liquefied coal and the like. Examples of suitable hydrocracking raw materials are gas oils and other hydrocarbon fractions having at least 50% by weight and usually more than 70% by weight of components having a boiling point above 371 ° C. Heavy hydrocarbon oils such as heavy crude oil, reduced crude oil, vacuum distillation residues and similar heavy materials are not normally suitable raw materials for the process of the invention.

Typically, the above raw materials are treated with hydrogen before they have been subjected to the hydrocracking process of the invention. Hydrogen treatment is carried out in conjunction with hydrocracking, usually in a manner referred to as an integral operation. In this process, the hydrocarbon feedstock is introduced into the catalytic hydrogen treatment zone, where in the presence of a suitable catalyst and under suitable conditions, i. at elevated temperatures, such as from about 204 ° C to about 538 ° C, and at elevated pressures, such as from about 0.66 to about 33.33 MPa, and with the hydrogen reagent, the organo-nitrogen components and organosulfur components contained in the feedstock are converted to ammonia and hydrogen sulfide. Suitable hydrogen treatment catalysts include zeolites or non-permeable granular catalysts containing Group VIII metals and Group VIB metals on a porous inorganic refractory support, most commonly of alumina. The entire effluent withdrawn from the hydrogen treatment zone is then treated in a hydrocracking zone maintained under suitable conditions at elevated temperature, pressure and partial hydrogen pressure and containing the hydrocracking catalyst of the invention. Typically, the hydrogen treatment zone and the hydrocracking zone in an integral operation are maintained in separate reaction vessels, or it may be advantageous to use a single downstream reaction vessel containing one or more upper beds of hydrogen processing catalyst particles and one or more lower beds of hydrocracking catalyst particles. Examples of integral operation are in U.S. Patents 3,159,564, 3,655,551, 4,040,944, and 4,584,287. In some cases, the hydrocracking zone effluent is subjected to hydrogen treatment in a manner similar to that described above to remove mercaptans from the article.

The catalyst of the present invention is typically used as a fixed catalyst bed in a hydrocracking reactor into which hydrogen and feedstock are fed and downstream.

The reactor vessel is held to the desired product, containing and powered

Generally, under conversion conditions, which is normally a hydrocarbon substantial fraction of fuel for diesel fuel having a boiling point between 149, the reaction vessel temperature is maintained at 454 ° C, preferably between 260 ° C and 426 ° C. The pressure is ´C and the raw material product of turbine 371 ’C.

232 C between normal between

MPa and about 23.33 MPa, preferably between 6.66 MPa and 20 MPa. The hourly space velocity of the liquid is from 0.3 to 5.0, preferably from 0.5 to 3.0 h ^-1 . The ratio of hydrogen gas to feedstock used is generally from 28.31 m ³ to 283.1 m ³ per barrel, preferably from 56.63 m ³ to 226.5 m ³ per barrel, measured at 60 ° C and a pressure of 0.1

Bar.

A typical raw material, gas oil, does not contain more than 35% by volume, usually less than 15% by volume, of boiling points between 149 ° C and 371 ° C. Where medium boiling distillates are required, the hydrocracking conditions shall be chosen such that at least 80% by volume, preferably 86% by volume, most preferably 87% by volume of the product having a boiling point of up to 371 ° C has a boiling point between 149 ° C and 371 ° C. Typically, an article having a boiling point of up to 371 ° C contains more than 75% by volume, preferably at least 80% by volume, most preferably 83% by volume of hydrocarbons having a boiling point of from 149 ° C to 287 ° C.

Based on the data currently available, the catalyst of the invention, compared to a commercial hydrocracking catalyst for the production of medium boiling distillates containing rare earth cation exchange LZ-10 zeolite and having a cell size of 24.30 angstroms, produces improved results when used for selective production of fuel for combustion turbines and diesel fuels. In particular, the catalyst of the invention has a higher activity and significantly increases the yield of hydrocarbon distillates having a boiling point of from 149 ° C to 287 ° C and from 149 ° C to 371 ° C. The results obtained and other facts are shown in the following example, which is intended to illustrate and not in any way limit the scope of the invention defined in the claims.

Example

Catalyst 1

Catalyst 1 is a catalyst according to the invention and was prepared by mixing 5% by weight of zeolite LZ-10, 75% by weight of a mixture of amorphous silica and alumina (a mixture of silica and alumina, manufactured by SM-30, Davison Chemical Division of WR Grace) Co.), containing 70% by weight silica and 30% by weight alumina and about 20% by weight of a peptized Catapal alumina binder. The UOP-derived LZ-10 zeolite had a cell size of 24.30 angstroms, an effective pore size of greater than 7.0 angstroms, and a total silica to alumina molar ratio of about

5.2. The mixture was stirred and then passed through a 1.5875 diameter cylindrical die to produce cylindrical prints that were cut to a length of 3.175 mm to 12.7 mm. The prints were dried at 100 ° C and then calcined at 482 ° C. The dried and calcined prints were then impregnated by the method of saturating the pores with an aqueous solution containing nickel nitrate and ammonium tungstate in sufficient amounts such that after drying at 100 ° C and calcining at 482 ° C the resulting catalyst particles contained about 5% Ni by weight calculated from NiO. 22% by weight of tungsten, calculated as WO ₃ , on a substrate consisting of 5% by weight of zeolite LZ-10, 75% by weight of a mixture of silica and magnesium oxide and 20% by weight of an alumina binder.

Catalyst 2

Catalyst 2, the comparative catalyst, was prepared similar to Catalyst 1 except that amorphous gamma alumina was used instead of amorphous silica and magnesium oxide. The finished catalyst contained nickel and tungsten in the proportions listed above for Catalyst 1 on a support consisting of 5 wt% zeolite LZ-10, 75 wt% amorphous gamma alumina and 20 wt% alumina as a binder.

Catalyst 3

Catalyst 3, another comparative catalyst, similar to silica and silica except that the magnesium oxide was an alumina containing 25% by weight of nickel and the tungsten catalyst was higher for the silica and alumina by weight and 20 was prepared instead of amorphous amorphous 75% alumina.

Catalyst 1 based on zeolite LZ-10, 75% oxide oxide oxide Finished in proportions consisting of 5% by weight amorphous oxide% by weight alumina as a binder.

Catalyst 4

Catalyst 4, another comparative catalyst, was prepared similar to Catalyst 1 except that the amorphous silica and magnesium oxide were replaced by a dispersion of silica and alumina particles in a gamma alumina matrix (Aero 5545 product from Criterion Catalyst Company LP). . The dispersion was amorphous and contained 55% by weight alumina and 45% by weight silica. The finished catalyst contained nickel and tungsten in proportions indicated for catalyst 1 on a support consisting of 5% by weight of LZ-10 zeolite, 75% by weight of an amorphous dispersion of silica and alumina in gamma alumina and 20% by weight of alumina as a binder.

Catalyst 5

Catalyst 5 is a commercial catalyst suitable for hydrocracking to produce intermediate boiling distillates obtained from UOP. It was prepared similarly to Catalyst 4 except that it contained 7 wt% nickel, calculated as NiO, 10 wt% LZ-10 zeolite and 70 wt% amorphous dispersion of silica and alumina particles in the gamma alumina matrix used in catalyst 4 . the catalyst thus contained 7% nickel, calculated as NiO, and 22% by weight of tungsten, calculated as W0 _3, the support consisting of 10% w LZ-10, 70% by weight of a dispersion of silica-alumina in a matrix of gamma alumina and 20% alumina as a binder.

Catalyst 6

Catalyst 6 was prepared similarly to Catalyst 1 except that 10% by weight of zeolite LZ-10 and 70% by weight of the same mixture of silica and magnesium oxide were used. The finished catalyst thus contained 5% by weight of nickel, calculated as NiO, and 22% by weight of tungsten, calculated as W0 _3, the support consisting of 10% w LZ-10, 70% by weight of silica and magnesium oxide and 20% by weight of alumina binder.

Each of the catalysts described above was pretreated with a flow of gas containing 10% by weight hydrogen sulfide and the rest of the hydrogen through the catalyst bed at an initial temperature of 149 ° C slowly raised to 371 ° C and then held for 1 hour.

The six catalysts were then tested for activity and selectivity in hydrocracking to produce medium boiling distillates using a non-hydrogenated light Arabian vacuum gas oil having an API weight of 37 ° C and an initial boiling point of 224 ° C, a final boiling point of 578 ° C and 50%. a boiling point of 434 ° C at about 20% by volume having a boiling point below 370 ° C and 5% by volume having a boiling point below 309 ° C as determined by modified distillation according to ASTM D1160. The gas oil, which contained about 0.085% by weight of nitrogen calculated as element, was blown once through an isothermal reactor containing about 140 ml of a quartz mixed catalyst having a particle size of 6 to 8 mesh. The reactor was operated at a liquid hourly space velocity of 1.0 h -1, a total pressure of 13.33 MPa and one hydrogen flow rate of 283.1 m ³ per barrel. The reactor temperature was set to ensure the conversion of 60% by volume to materials having a boiling point below 371 ° C. Thereafter, tert, butylamine and thiophene were added to the reactor in amounts proportional to the amounts of ammonia and hydrogen sulfide that should be present from the treatment of the gas oil with the hydrogen from which the feedstock was derived. The conditions under which the catalysts were tested simulated those conditions that could be expected in a hydrocracking vessel used in the integral hydrogen treatment and hydrocracking process, where the total amount exiting the hydrogen treatment reactor and the hydrogen added were transferred to the hydrocracking reactor for further refining, to convert into a medium boiling distillate. The test results are given in Table 1.

Table 1

Cat. Activity (° C) Reactor temperature 60% Conversion selectivity % By volume to 371 ° C Turbine 149-287 ° C Diesel 149-371-C product base into 371 ° C 1 75% of amorphous oxide S i-Mg 20% spoj iva 387 79.4 89.8 5% zeolite LZ-10 2 75% of amorphous oxide A1 20% binder 387 76.7 84.1 5% zeolite LZ-10 3 75% of amorphous Si-Al oxide 20% spoj iva 389 74.2 85.0 5% zeolite LZ-10 4 75% of amorphous Si-Al oxide in oxide A1 20% spoj iva 389 71.1 83.0 5% zeolite LZ-10 5 70% of amorphous Si-Al oxide in oxide A1 20% binder 381 71.3  83.4 10% zeolite LZ-10 6 70% of amorphous oxide S i-Mg 20% spoj iva 381 70.3 82.1 10% zeolite LZ-10

As can be seen from the data in Table I, catalysts 1 to 4, which differ only in the amorphous component of their supports, have almost the same activity, i. from 387 ° C to 389 ° C. However, the catalyst 1, i.e. the catalyst according to the invention, has a considerably high and unexpected selectivity for fuel for combustion turbines and for diesel fuel. Differences in selectivity for the range of fuel for combustion turbines vary from high 8.3% (79.4% by volume).

- 71.1) between catalysts 1 and 4 to a low 5.2% by volume (79.4 - 74.2) between catalysts 1 and 3. The differences in selectivity for diesel fuel change from a high 6.8% by volume (89.8

- 83.0) between catalysts 1 and 4 to a low 4.8% by volume (89.8 - 85.0) between catalysts 1 and 3. The fact that the presence of: an amorphous component of silica and magnesium oxide in catalyst 1 would result in to such a significant increase in selectivity for both turbine fuel and diesel fuel compared to similar catalysts in which the amorphous component was alumina, a mixture of silica and alumina or a dispersion of silica and alumina in gamma alumina is surprising and unexpected .

Such an increase in selectivity is extremely significant in view of the fact that an increase of 4.0% is considered commercially significant because there is significantly more raw material converted to the desired product with fewer less valuable products. For example, in a one-time hydrocracking process that yields 14,000 barrels per day of diesel fuel, an increase in selectivity of 4.8% associated with Catalyst 1 over Catalyst 4 yields 672 barrels more per day and about 210,000 barrels more per year of the desired product. Assuming a price of one barrel of $ 40, this increase in selectivity is worth $ 8.5 million per year for the refiner, which is quite a lot of money.

Comparison of Catalyst 1 with Catalyst 5, a commercial catalyst that differs from Catalyst 1 only in that it contains 5 wt% more zeolite LZ-10 and 70 wt% dispersion of a mixture of silica and alumina in alumina instead of 75 wt% oxide mixture of silica and magnesium oxide, shows that Catalyst 1 is 5.55 ° less active (387 ° C - 381 ° C) than Catalyst 5. This lower activity is expected because Catalyst 1 contains less zeolite than Catalyst 5. Catalyst 1 is however, much more selective. According to the data, Catalyst 1 gives 8.1% by volume (79.4-71.3) more fuel for combustion turbines with a boiling point between 149 ° C and 287 ° C and 6.4% by volume (89.9-83). , 4) More diesel fuel with a boiling point between 149 ° C and 371 ° C. These differences in selectivity are commercially very significant. If an analysis similar to the above is performed due to the selectivity differences for diesel fuel between catalyst 1 and the commercial catalyst, i. Catalyst 3, use of Catalyst 1 instead of a commercial catalyst, i. Catalyst 5, results in an income of about $ 11.3 million a year higher for the refiner.

A comparison of Catalyst 1 with Catalyst 6, which differs from Catalyst 1 in that it contains 5% by weight more zeolite LZ-10 and 5% by weight less silica-magnesium oxide mixture, shows what could be expected: with regard to the increased zeolite content according to claim 1, wherein catalyst 1 is 5.55 ° C less active (387 ° C - 381 ° C). A comparison of the selectivity data for both catalysts shows that Catalyst 1 is much more selective for both fuel for combustion turbines (79.4% v / v 70.3% v / v) and diesel (89.8% v / v 82.1% v / v). Since the only differences between the two catalysts are in the amount of zeolite and the amount of amorphous silica and magnesium oxide, these data appear to indicate that it is desirable to maintain the zeolite content in the substrate below about 10% by weight and the amount of silica mixture to maintain high selectivities. and magnesium oxide above about 70% by weight. Thus, it appears that the preferred catalysts of the invention will have. substrates containing from 2 to 90% by weight of zeolite Y, preferably from 3 to 8% by weight and from 70 to 90% by weight of a mixture of silica and magnesium oxide, preferably from 75 to 85% by weight.

Although the invention was originally described in connection with one example and with reference to an embodiment thereof, it is obvious that it may have been made a number of variations and variations that will be apparent to one skilled in the art without departing from the spirit of the invention.

Claims (32)

PATENT CLAIMS A hydrocracking process for the selective production of intermediate boiling distillates comprising contacting a hydrocarbonaceous feedstock with a hydrocracking catalyst under elevated temperature and pressure conditions in the presence of hydrogen to produce a lower medium boiling product, characterized in that the catalyst comprises (a) one or more and (b) a zeolite Y selected from the group consisting of Y zeolites having a size. cells less than 24.25 angstroms and Y zeolites having a water vapor sorption capacity of less than 10% by weight at 25 ° C and a ρ / ρθ value of 0.1; and (c) an amorphous silica and magnesium oxide component. The method of claim 1, wherein the Y zeolite has a cell size of less than 24.45 angstroms. 3. The method according to claim 1, • wherein said Y zeolite has a water vapor sorptive capacity less than about 10% by weight at 25 DEG C. and a p / p _o value of 0.10. 4. The method of claim 1, wherein the Y zeolite has a unit cell size below 24.45 angstroms and a water vapor sorptive capacity less than about 10% by weight at 25 DEG C. and a p / p _o value of 0.10. The process of claim 1, wherein the catalyst further comprises an amorphous inorganic refractory oxide as a binder. The method of claim 5, wherein said inorganic refractory oxide is alumina. A process according to claim 1, wherein the medium boiling distillate comprises components boiling below 371 ° C and more than 86% by volume of these components has a boiling point between 149 ° C and 371 ° C. 8. The process of claim 1 wherein the zeolite Y is UHP-Y zeolite. 9. The process of claim 1 wherein the zeolite Y has undergone ion exchange with rare earth cations. The process of claim 9, wherein the zeolite Y contains at least 2% by weight of rare earth metals, calculated as RE ₂ O ₃ . The method of claim 1, wherein the medium boiling distillate comprises components boiling below 371 ° C and more than 75% by volume of these components have boiling points between 149 ° C and 287 ° C. The method of claim 1, wherein the zeolite Y has a total silica to alumina molar ratio of less than 6.0. The process of claim 6, wherein the catalyst comprises a hydrogenation component of nickel and a hydrogenation component of tungsten and the zeolite is zeolite LZ-10 and more than 86% by volume of the hydrocarbon fraction boiling up to 371 ° C has a boiling point between 149 ° C and 371 C. 14. The process of claim 1 wherein the zeolite Y is prepared by a process comprising (a) partial cation exchange between ammonia and sodium zeolite Y, (b) calcining the resulting zeolite in the presence of water vapor, (c) a second cation exchange between ammonia and zeolite, (d) calcining the zeolite from step (c) in the presence of water vapor. The process of claim 14, wherein the zeolite Y is prepared by reducing sodium zeolite Y in step (a) to a sodium content of from 0.6 to 5% by weight, calculated as Na ₂ O, using a vapor atmosphere in step (b) and zeolite cell size it is reduced to between 24.40 and 24.64 angstroms, in step (c) the ion exchange with ammonia reduces the sodium content to below 0.5% by weight, calculated as Na ₂ O, and in step (d), calcination is carried out in the presence of steam and the size of the zeolite cells after calcination is less than 24.40 angstroms. The method of claim 15, wherein the zeolite in step (d) has a final cell size of between 24.20 and 24.35 angstroms. 17. The process of claim 1 wherein the catalyst comprises a Group VIII metal hydrogenation component comprising nickel, cobalt and oxides and sulfides thereof and a Group VIB metal hydrogenation component selected from molybdenum, tungsten, and oxides and sulfides thereof. Process according to claim 17, characterized in that at least 70% by volume of the hydrocarbonaceous feedstock components have a boiling point above 371 ° C and the process is carried out under conditions such that the distillate produced having a boiling point lower than 371 ° C contains at least 87%. % by volume of the components having a boiling point between 149 ° C and 371 ° C. The method of claim 18, wherein the lower intermediate boiling product comprises boiling components. lower than 371 ° C and more as 76 % by volume of these ingredients my boiling point between Mp 149 ° C and 287 ° C. 20. Method by claim 18, characterized by by that the catalyst comprises from 4 up to 8% % zeolite Y • 21. Method by claim 20, characterized by by that the catalyst contains from 50 to 65% by weight of silica and magnesium oxide. A hydrocracking process which comprises contacting a hydrocarbonaceous feedstock with a catalyst under elevated temperature and pressure conditions and in the presence of hydrogen to produce a lower intermediate boiling distillate, wherein the catalyst comprises one or more hydrogenation components in combination with a support comprising (and (b) a Y zeolite having a total silica to alumina molar ratio of less than 6.0, a cell size between 24.20 and 24.35 angstroms, and a water vapor sorption capacity of less than 10% by weight at 25 '; C and a p / p value _of 0.10; and (c) a component of amorphous silica and magnesium oxide. The method of claim 22, wherein the zeolite Y has a water vapor sorptive capacity less than about 5% by weight of zeolite at 25 ° C and a p / p _o value of 0.10. 24. The method of claim 23, wherein the hydrogenation component comprises components selected from the group consisting of metals and metal compounds of Groups VIB and VIII. The method of claim 24, wherein the Group VIB metal hydrogenation component is selected from the group consisting of molybdenum, tungsten and oxides and sulfides thereof and the Group VIII metal hydrogenation component is selected from the group consisting of nickel, cobalt and oxides thereof, and sulphides. The method of claim 22, wherein the lower intermediate boiling distillate comprises components boiling below 371 ° C and more than 88% by volume of these components have boiling points between 149 ° C and 371 ° C. 27. The process of claim 23 wherein the zeolite Y has undergone ion exchange with rare earth cations. 28. The method of claim 23, wherein the porous inorganic refractory oxide binder comprises alumina. 29. The process of claim 28 wherein the catalyst comprises a hydrogenation component of nickel and a hydrogenation component of tungsten. 30. process according to claim 29 characterized the by that zeolite Y is zeolite LZ-10 First 31st process according to claim 22 characterized the by that support It contains from 2 to 9% by weight zeolite Y. 32nd process according to claim 27 characterized the by that zeolite Y contains at least 4% by weight of rare earth metals, calculated as RE2O3. 33. The method of claim 30, wherein the lower middle boiling distillate comprises components boiling below 371 [deg.] C and more than 77% by volume of these components boiling from 149 [deg.] C to 287 [deg.] C. 34. The process of claim 22, wherein the hydrocarbon feedstock is treated with hydrogen prior to contact with the catalyst. 35. The method of claim 31, wherein said substrate comprises from 75% to 85% by weight of an amorphous silica and magnesium oxide component. 36. A hydrocracking catalyst comprising (a) one or more hydrogenation components and (b) a zeolite Y selected from the group consisting of Y zeolites having a cell size of less than 24.45 angstroms and Y zeolites having a water vapor sorption capacity of less than 10% by weight at 25 DEG C. and a p / p _o of 0.10, and (c) component of the amorphous silica and magnesium oxide. Hydrocracking catalyst according to claim 36, characterized in that it comprises a porous, inorganic refractory oxide binder.