FI71504B - SAMTIDIG KATALYTISK HYDROKRACKNING OCH -AVVAXNING AV KOLVAETEOLJOR MED ZEOLIT BETA - Google Patents

Description

χ 71504

This invention relates to a process for the catalytic hydrocracking and dewaxing of hydrocarbon feedstocks for the production of low pour point distillates and reduced viscosity heavy fuel oils.

Catalytic dewaxing of hydrocarbon oils to lower the temperature at which waxy hydrocarbons are separated is a known method. A method of this nature is described in The Oil and Gas Journal, dated January 6, 1975, pages 69-73. U.S. Patents 3,668,113 and 3,894,938 also describe dewaxing followed by hydrofinishing.

Reissued Patent 28398 describes a catalytic dewaxing process using a catalyst consisting of a ZSM-5 type zeolite. A hydrogenation / dehydrogenation component may be present.

A process for removing wax from gas oil with a ZSM-5 type catalyst is described in U.S. Patent 3,956,102.

A mordenite catalyst containing a Group VI or Group VIII metal is used to remove wax from a low viscosity distillate obtained from a trivial crude oil, as described in U.S. Patent 4,110,056.

U.S. Patent 3,755,138 describes a mild solvent dewaxing process for removing high quality wax from a lubricating oil feedstock, which is then catalytically dewaxed to the solidification point of the specification.

U.S. Patent 3,923,641 describes a process for hydrocracking naphtha using zeolite beta as a catalyst.

2 71504

Hydrocracking is a well-known process and various zeolite catalysts have been used in hydrocracking processes, but although they may be effective in obtaining distillate yields with one or more properties consistent with the intended use of the distillate, these catalysts have generally been disadvantaged by yielding product yields. good low temperature flow properties, especially reduced pour point and viscosity. The catalysts used for hydrocracking contain an acid component and a hydrogenation component. The hydrogenation component may be a noble metal such as platinum or palladium or a base metal such as nickel, molybdenum or tungsten or a combination of these metals. The acid cracking component may be an amorphous material such as acid clay or amorphous silica-alumina or alternatively zeolite. High porosity zeolites such as zeolites X and Y have been conventionally used for this purpose because the main components of the feedstocks (gas oils, coking plant bottoms, refined crude oils, recycled oils, FCC bottoms) are not therefore no conversion will occur. Thus, if waxy feedstocks such as Amal gas oil are hydrocracked with a high porosity catalyst such as zeolite Y along with the hydrogenation component, the viscosity of the oil is calculated by cracking most of the 343 ° C + material into a material boiling between 343-165 ° C. The remainder of the 343 ° C + material, which is not converted, contains most of the paraffinic components of the feedstock because the aromatics are primarily converted to paraffins. The unconverted 343 ° C + material therefore retains its high pour point, so that the final product also has a relatively high pour point of about 10 ° C. Thus, although the viscosity decreases, the pour point is still not acceptable. Even when conditions are adjusted to produce complete or near-complete conversion, the higher molecular weight hydrocarbons present in the feedstock, mainly polycyclic aromatics, crack, leading to a further decrease in product viscosity.

3 71504

However, cracked products contain a considerable amount of straight-chain components (n-paraffins) which, if they themselves have a sufficiently high molecular weight, as they often have, form a waxy component in the product. The final product may therefore be relatively more waxy than the feedstock and, as a result, may have a pour point that is equally unsatisfactory or even more unsatisfactory. An additional disadvantage of operating under high conversion conditions is that hydrogen consumption increases. Attempts to reduce the molecular weight of these straight-chain paraffinic products only work to produce very light fractions, such as propane, which reduces the desired liquid yield.

On the other hand, the dewaxing process uses a low porosity zeolite or a form-selective zeolite such as ZSM-5 as the acidic component of the catalyst and the normal and slightly branched paraffins present in the feedstock are able to enter the zeolite's inner pore structure. The majority, typically about 70% of the feedstock boiling above 343 ° C, remains unconverted because bulky aromatic components, especially polycyclic aromatics, are unable to enter the zeolite. The paraffinic waxy components are therefore removed, which lowers the solidification point of the product, but the other components are retained, so that the final product has too high a viscosity to be acceptable, although the solidification point may be satisfactory.

It has now been found that heavy hydrocarbon oils can be simultaneously hydrocracked and dewaxed by hydrogenation to form a liquid product with a satisfactory pour point and viscosity. This desired result is obtained by using a catalyst containing zeolite beta as an acidic component to cause cracking reactions. The catalyst preferably contains a hydrogenation component to cause hydrogenation reactions. The hydrogenation component may be a noble metal or a base metal and is suitably of the conventional type, e.g. nickel, tungsten, cobalt, molybdenum or combinations of these metals.

4 71504 According to the present invention, there is provided a process for cracking and dewaxing heavy hydrocarbon oil, which process comprises contacting the oil with a zeolite beta-containing catalyst.

In the process of the invention, the hydrocarbon feedstock is heated with a catalyst under conversion conditions suitable for hydrocracking. During the conversion, the aromatics and naphthenes in the feedstock undergo hydro-cracking reactions such as dealkylation, ring opening and cracking, followed by hydrogenation. In addition, the long-chain paraffins present in the feedstock, together with the paraffins formed by hydrocracking the aromatics, are converted to products that are less waxy than the straight-chain n-paraffins, resulting in simultaneous dewaxing. The use of zeolite beta is thought to be unique in this respect, as it not only causes a decrease in the viscosity of the product by hydrocracking, but at the same time also lowers the solidification point by catalytic hydro-wax removal.

This process makes it possible to convert heavy feedstocks, such as gas oils boiling above 343 ° C, into distillate products boiling below 343 ° C, but in contrast to previous processes using high-porosity catalysts such as zeolite Y. , hydrogen consumption decreases even if the product is consistent with the desired specifications for pour point and viscosity. In contrast to dewaxing processes using shape-selective catalysts such as zeolite ZSM-5, total conversion occurs, including cracking of aromatic components, which guarantees an acceptably low viscosity for the distillate product.

Thus, this process is able to achieve total conversion along with simultaneous dewaxing. Besides, this is achieved with reduced hydrogen consumption compared to other process types. It is also possible to operate with a partial version, which provides economy in hydrogen consumption while still meeting the pour point and viscosity requirements. The process also achieves improved selectivity for the production of distillate materials; the yield of products boiling below the gas and distillate range decreases.

As mentioned above, the process combines hydrocracking and dewaxing features. The catalyst used in the process has an acidic component and a hydrogenation component, which may be of the conventional type. The acidic component consists of the zeolite beta described in U.S. Patents 3,303,069 and Re 28341, and these patents are incorporated herein by reference for this zeolite and its preparation details.

Zeolite beta is a crystalline aluminosilicate zeolite with a pore size of more than 5 Angstroms. The composition of this zeolite as disclosed in U.S. Patents 3,303,069 and Re 28341 in its synthesized form may be expressed as follows: /xNa(i,o-o,l-x)tea7aio2.YSi02.wh2o wherein X is less than 1, preferably less than 0.7; TEA represents tetraethylammonium ion; Y is greater than 5 but less than 100 and W is at most 60 (it has been found that the degree of hydration may be greater than originally determined when W was determined to be at most 4) depending on the degree of hydration and the metal cation present. The TEA component is calculated as one of the differences between the analyzed value of sodium and the theoretical cation and structural aluminum.

In the fully base-exchanged form, beta has the composition /"^M(1-0,1-X)H?.A102.YSi02.wh 2o where X, Y and W have the values listed above and n is the valence of the metal M .

In a partially base-exchanged form obtained from the original sodium form of the zeolite by ion exchange without calcination, the zeolite in beta has the formula: 6 71504

/ ~ γΗ (Ι ^ Ο, 1-Χ) TEα7α10. YSiO .WH O

When zeolite is used in catalysts, it is at least partially in hydrogen form to provide the desired acidic functionality for the cracking reactions that are to occur. Normally, it is preferred to use the zeolite in a form having sufficient acidic functionality to give it an alpha value of 1 or more. The alpha value, which is a measure of the acid functionality of a zeolite, is described in conjunction with the details of its measurement in U.S. Patent 4,016,218 and J. Catalysis, Vol. VI, pages 278-287 (1966) and these publications are referred to for these details. The acid functionality can be adjusted by base exchange of the zeolite, especially with alkali metal cations such as sodium, by evaporation or by adjusting the silica / alumina ratio of the zeolite.

Once the zeolite beta is synthesized in the alkali metal form, it can be converted to the hydrogen form by forming the ammonium intermediate as a result of ammonium ion exchange and calcining the ammonium form to obtain the hydrogen form. In addition to the hydrogen form, other forms of zeolite in which the original alkali metal has been reduced to less than about 1.5% by weight may be used. Thus, the original alkali metal of the zeolite can be replaced using ion exchange with other suitable metal cations, such as nickel, copper, zinc, palladium, calcium, and rare earth metals.

In addition to having a composition as defined above, zeolite beta can be characterized by its X-ray diffraction data as disclosed in U.S. Patents 3,308,069 and Re. 28341. The relevant d-values (Angstroms, radiation: copper K-alpha doublet, Geiger counter spectrometer) are shown in Table I below: 7 71504

Table I

D-values of reflections in zeolite beta 11.40 + 0.2 7.40 + 0.2 6.70 + 0.2 4.25 + 0.1 3.97 + 0.1 3.00 + 0.1 2, Preferred forms of zeolite beta for use in this process are silica-rich forms with a silica / alumina ratio of at least 30: 1. In fact, it has been found that zeolite beta can be prepared with silica / alumina ratios above the maximum ratio of 100: 1 defined in U.S. Patents 3,308,069 and Re. 28341 and these forms of zeolite provide the best performance in the process. Ratios of at least 50: 1 and preferably at least 100: 1 or even greater, for example 250: 1, 500: 1, may be used.

The silica / alumina ratios mentioned herein are structural or bulk ratios, i.e., the ratio between the SiO 2 and AlO 4 quaternaries, which together form the structure of which the zeolite is composed. It is to be understood that this ratio may vary from the silica / alumina ratio determined by various physical and chemical methods. For example, the overall chemical analysis may include aluminum present in the form of cations associated with the acidic sites of the zeolite, resulting in a low silica / alumina ratio. Similarly, if the ratio is determined by thermogravimetric analysis (TGA) of ammonia desorption, a low ammonia titration can be obtained if cationic aluminum prevents the exchange of ammonium ions at acidic sites. These differences are particularly problematic when certain treatments are used, such as the demineralization process described below, which results in the presence of ionic aluminum free of zeolite structure. Therefore, due care must be taken to ensure that the silica / alumina ratio of the carcass is correctly determined.

The silica / alumina ratio of the zeolite may be determined by the nature of the starting materials used in its preparation and their relative proportions. Some variation in ratio can therefore be obtained by varying the relative concentration of the silica precursor compared to the alumina precursor, but clear limits on the maximum achievable silica / alumina ratio of the zeolite can be observed. For zeolite beta, this limit is usually about 100: 1 (although higher ratios can be obtained) and at ratios above this value, other methods are usually necessary to produce the desired silica-rich zeolite. In one such process, aluminum is removed by extraction with an acid and in this process the zeolite is contacted with an acid, preferably a mineral acid such as hydrochloric acid. Aluminum removal is easy at ambient and slightly elevated temperatures and occurs with minimal loss of crystallinity to form zeolite beta-rich silica forms with silica / alumina ratios of at least 100: 1, with ratios of 300: 1 or higher being readily achievable.

Zeolite is conventionally used in the hydrogen form for the aluminum removal process, although other cationic forms, e.g., the sodium form, may be used. If these other forms are used, sufficient acid must be used to allow the original cations of the zeolite to be replaced by protons. The amount of zeolite in the zeolite / acid mixture should generally be between 5-60% by weight.

The acid may be a mineral acid, i.e. an inorganic acid or an organic acid. Typical inorganic acids that can be used include mineral acids such as hydrochloric, sulfuric, nitric and phosphoric acids, peroxydisulfonic acid, dithionic acid, sulfamic acid, peroxymonosulphuric acid, amidodisulfonic acid, pyrosulfonic acid, nitrosulfonic acid, nitrosulfonic acid Typical organic acids that can be used include formic acid, trichloroacetic acid and trifluoroacetic acid.

The concentration of the added acid should be such that it does not lower the pH of the reaction mixture to an undesirably low level that could affect the crystallinity of the zeolite being treated. The acidity that the zeolite can tolerate depends at least in part on the silica / alumina ratio of the starting material. In general, it has been found that zeolite beta is able to withstand concentrated acid without undue loss of crystallinity, but as a general rule, the acid should be 0.1-4.0-N, usually 1-2-N. These values remain good regardless of the silica / alumina ratio of the zeolite beta starting material. Stronger acids tend to provide a relatively higher degree of aluminum removal than weaker acids.

The aluminum removal reaction proceeds rapidly at ambient temperatures, but slightly elevated temperatures, e.g. up to 100 ° C, can be used. The duration of the extraction affects the silica / alumina ratio of the product, as the extraction, which is controlled by diffusion, depends on the time. However, since the zeolite gradually becomes more resistant to loss of crystallinity as the silica / alumina ratio increases, i.e., it becomes more stable when aluminum is removed, higher temperatures and more concentrated acids can be used towards the end of treatment than at the beginning without the associated risk of loss of crystallinity.

After the extraction treatment, the product is washed with water free of impurities, preferably distilled water, until the pH of the leaving wash water is approximately between 5-8.

The crystalline dealuminated products obtained by the process of this invention have substantially the same crystallographic structure as the starting aluminosilicate zeolite, but with increased silica / alumina ratios. The formula of dealuminated zeolite 10 71 5 0 4 beta is therefore (1-0.1-X) h7aio2.YSiO2.WH2o where x is less than 1, preferably less than 0.15, Y is at least 100, preferably at least 150 and W is at most 60. M is a metal, preferably a transition metal or a metal of Groups IA, 2A or 3A or a mixture of metals. The silica / alumina ratio Y is generally between 100: 1 and 500: 1, more usually between 150: 1 and 300: 1, for example 200: 1 or more. The X-ray diffraction pattern of the dealuminated zeolite is substantially the same as that of the original zeolite, as shown in Table 1 above.

If desired, the zeolite can be vaporized prior to acid extraction to increase the silica / alumina ratio and make the zeolite more stable to the acid. Evaporation can also act to improve the ease with which the acid is removed and to promote the retention of crystallinity during the extraction process.

Zeolite beta is preferably used in combination with a hydrogenation component, usually derived from the metal of groups VA, VIA or HATE of the Periodic Table. Preferred base metals include tungsten, vanadium, molybdenum, nickel, cobalt, chromium and manganese, and preferred noble metals include platinum, palladium, iridium and rhodium.

Base metal combinations such as cobalt-molybdenum, cobalt-nickel, nickel-tungsten or cobalt-nickel-tungsten are exceptionally useful with many feedstocks and in a preferred combination the hydrogenation component contains 0.7-7% by weight nickel and 2.1-21% by weight. % by weight of tungsten expressed as metal. The hydrogenation component can be replaced on the surface of the zeolite, absorbed into it or physically mixed with it. If the metal is to be absorbed into the zeolite or even on its surface, it can be done, for example, by treating the zeolite with an ion containing platinum metal. Suitable platinum compounds include chloroplatinic acid, platinum chloride and various compounds containing a platinum-amine complex.

11 7150 4

The catalyst can be treated by conventional pre-sulfidation treatments, for example by heating in the presence of hydrogen sulfide to convert metal oxide forms such as CoO or NiO to their corresponding sulfides.

Metal compounds can be either compounds in which the metal is present as a cation of the compound or compounds in which it is present as an anion of the compound. Both types of compounds can be used. Platinum compounds in which the metal is in the form of a cation or cationic complex, for example Pt (NH3) 4Cl2, are particularly useful, as are anionic complexes such as vanadate and metawolframic ions. The cationic forms of other metals are also very useful, as they can be exchanged on or absorbed into the surface of the zeolite.

Prior to use, the zeolite should be at least partially dehydrated. This can be done by heating to a temperature between 200-600 ° C in air or in an inert atmosphere such as nitrogen for 1-48 hours. Dehydration can also be performed at lower temperatures using vacuum alone, but a longer time is required to achieve a sufficient amount of dehydration.

It may be desirable to incorporate the catalyst into another material that can withstand the temperature and other conditions used in the process. Such matrix materials include synthetic and naturally occurring substances such as inorganic materials such as clay, silica and metal oxides. The latter can be either naturally occurring or in the form of gel-like fines or gels containing mixtures of silica and metal oxides. Naturally occurring clays can be mixed with zeolites, such as zeolites of the montmorillonite and kaolin groups. The clays can be used in the originally mined impure form or initially subjected to calcination, acid treatment, or chemical modification.

The zeolite can be mixed with a porous matrix material such as alumina, silica-alumina, silica, silica, zirconia, silica-thorium oxide, silica-alumina, silica, silica-titanium dioxide such as silica with magnesium oxide and silica-magnesium-oxide-zirconia. The matrix may be in the form of a keragel. The relative amounts of zeolite component and inorganic oxide gel matrix, calculated on the anhydrous basis, can vary widely, with the zeolite content ranging from 10 to 99, more usually from 25 to 80% by weight of the dry mixture. The matrix itself may have catalytic properties that are generally acidic in nature.

The feedstock for the process of the invention includes heavy hydrocarbon oil such as gas oil, coke tower bottom fraction, crude oil purified from lower boiling fractions, vacuum tower bottom fractions, asphalt purified vacuum residues, FCC tower bottom fractions or recycled oils. Oils derived from coal, shale or tar sands can also be treated in this way. Such oils generally boil above 343 ° C, although the process is also useful for oils with an initial boiling point of only 260 ° C. These heavy oils contain high molecular weight long chain paraffins and high molecular weight aromatics, much of which are condensed ring aromatics. During the processing, the condensed ring aromatics and naphthenes are cracked with an acid catalyst, and the paraffinic cracking products, together with the paraffinic components of the original feedstock, are converted to isoparaffins and to some extent to lower molecular weight materials. The hydrogenation of the unsaturated side chains in the monocyclic cracking residues of the original polycycles is catalyzed by the hydrogenation component to form substituted monocyclic aromatics, which are highly desirable end products.

The heavy hydrocarbon oil feedstock normally contains a significant amount of fraction boiling above 230 ° C and normally has an initial boiling point of about 290 ° C, more usually about 340 ° C. Typical boiling ranges are about 340-565 ° C or about 340-510 ° C, but 71504 oils with a narrower boiling range can, of course, be processed, for example those with a boiling range of about 340-455 ° C. Heavy gas oils are often such as are recycled oils and other non-residual materials. It is possible to simultaneously process materials boiling below 200 ° C, but the conversion of such components is lower. Feed materials containing such lighter tails normally have an initial boiling point above 150 ° C.

Highly paraffinic feeds are of particular benefit to the process because such feeds can provide the greatest improvement in pour point. However, most feeds contain a certain concentration of polycyclic aromatics.

The process is carried out under conditions similar to those used for conventional hydrocracking, although the use of a high-silicon zeolite catalyst makes it possible to lower the overall pressure requirements. Process temperatures of 230-500 ° C can be conventionally used, although temperatures above 425 ° C are not normally used because the thermodynamics of hydrocracking reactions become unfavorable at temperatures above this point. Generally, temperatures of 300-425 ° C are used. The total pressure is usually between 500 and 20,000 kPa and higher pressures in this range of more than 7,000 kPa are normally preferred. The process is used in the presence of hydrogen and the partial pressures of hydrogen are normally 600-6000 kPa. The ratio of hydrogen to hydrocarbon feedstock (hydrogen recycling rate) is 3 3 normally 10-3500 m / m. The volume flow rate of the feedstock is normally 0.1-20 LHSV, preferably 0.1-10 LHSV. At low conversions, the n-paraffins in the feedstock are primarily converted to isoparaffins, but at higher conversions, under more severe conditions, isoparaffins are also converted. The product contains few fractions boiling below 150 ° C and in most cases the product has a boiling range of 150-340 ° C.

14 71504

The conversion can be performed by contacting the feedstock with a solid, stationary bed, a solid fluidized bed or a transfer bed of a catalyst. A simple form is a flow bed operation in which the feed is allowed to flow through a stationary bed. With such a form, it is desirable to initiate the reaction with fresh catalyst at a reasonable temperature, which is naturally raised as the catalyst ages to maintain activity. The catalyst can be regenerated by contacting it at elevated temperature with, for example, hydrogen gas or by combustion in air or another oxygen-containing gas.

A preliminary hydrotreating step to remove nitrogen and sulfur and saturate the aromatics to naphthenes without a substantial change in boiling range usually improves catalyst performance and allows the use of lower temperatures, higher volume flow rates, lower pressures, or combinations of these conditions.

The method of the invention is illustrated by the following examples.

All parts and ratios in these examples are by weight unless otherwise indicated.

Example 1 This example illustrates the preparation of a catalyst.

A mixture of zeolite beta (SiO 2 / Al 2 O 3 = 30) with a crystallite size of less than 0.05 microns and an equal amount of gamma alumina calculated to be anhydrous was extruded to form 1.5 mm granules. The granules were calcined at 540 ° C under nitrogen, magnesium exchanged and then calcined in air.

One hundred grams of air-calcined compact was impregnated with 13.4 g of ammonium metatungstate (72.3% by weight) in 60 ml of water, followed by drying at 115 ° C and calcination in air at 540 ° C. The compact was then impregnated with 1.5 g of 71504 nickel nitrate hexahydrate in 60 ml of water and the wet granules were dried and calcined at 540 ° C.

The final catalyst had a nickel content of about 4% by weight as NiO and a calculated tungsten content of about 10.0% by weight as WO. The sodium content was less than 0.5% by weight as sodium oxide.

Example 2 This example illustrates the preparation of a silica-rich zeolite-beta.

A zeolite beta sample in its synthesized form with a silica / alumina ratio of 30: 1 was calcined in flowing nitrogen at 500 ° C for 4 hours and then in air at the same temperature for 5 hours. The calcined zeolite was then refluxed with 2-N hydrochloric acid at 95 ° C for 1 hour to form Deal-mininated, silica-rich zeolite beta with a silica / alumina ratio of 280: 1, an alpha value of 20, and an crystallinity of 80% compared to the original, which was assumed to be 100% crystalline.

The zeolite was converted to the ammonium form with 1-N ammonium chloride solution at reflux at 90 ° C for one hour, followed by exchange with 1-N magnesium chloride solution at reflux at 90 ° C for 1 hour. Platinum was added to the zeolite by ion exchange of the tetramine complex at room temperature. The metal-exchanged zeolite was washed thoroughly and dried in an oven by calcination in air at 350 ° C for 2 hours. The finished catalyst contained 0.6% platinum and was granulated, crushed and screened to a size of 0.35-0.50 mm.

Examples 3-5

The catalyst of Example 1 was evaluated for the catalytic conversion of Arab Light gas oil (HVGO) with an oil boiling range of 354-580 ° C. For comparison, the magnesium-exchanged zeolite Y (SiO 2 / Al 2 O 3 = 5) catalyst was also alloyed by extruding an equal amount of gamma alumina and impregnating it to contain 4% by weight of nickel and 16 71504 by weight of tungsten.

The composition of the feedstock, the conditions used, and the product analysis are shown in Table 2 below.

Table 2

Hydrocracking / dewaxing

Example Input 3 4 No.__

Catalyst - Mg-beta MgY

Conditions: Temperature, ° C - 423 414

Pressure, kPa - 7000 7000 LHSV, h "1 - 0.54 0.71 H2, n.1 / 1 - 1674 1318 H2 consumption, n.1 / 1 - 125 193 343 ° C + conversion,% by weight - 62.2 56.6

Features:

Dry gas + - 3.5 4.1 e ,. - 165 ° C fuel oil,% by weight - 16.6 24.7 165-343 ° C, distillate,% by weight - 40.6 26.3 343 ° C +,% by weight 100 37.8 44.0 343 ° C + pour point, ° C 40 -1 35 343 ° C + 95% TBP, ° C 552 468 528 343 ° C + properties

Density, ° API 22.0

Hydrogen,% by weight 12.07 12.7 13.7

Sulfur 2.45 0.04 0.03

Nitrogen 600 80 18 Freezing point, ° C 40 -1 35

Paraffins,% by volume 24.0 31 40

Naphthenes,% by volume 25.3 28 35

Aromatics,% by volume 50.7 41 25

As shown in Table 2 above, with a relatively high conversion of about 60%, the beta catalyst significantly reduced the freezing point of the 17 71 504 343 ° C + product, while the products obtained with the zeolite Y catalysts remained waxy. In addition, the beta catalyst converted significantly more of the high boiling components in the feed, resulting in an end point of 343 ° C + product that was about 55 ° C lower than that obtained with the catalyst of Example 4. Hydrogen consumption was also significantly lower whether it was an absolute value or a ratio to conversion.

For comparison, an Arab Light HVGO gas oil similar to Example 3 with a boiling range of 370-550 ° C was hydrocracked over rare earth-exchanged ultrastable zeolite Y (SiO 2 / Al 2 O = 75). The zeolite was prepared by steam calcination and dealumination of zeolite Y in a SiO 2 / Al 2 O 2 ratio of the hull of 75: 1, followed by exchange with rare earth metal, extrusion with an equal amount of gamma alumina and impregnation to contain 2% by weight of nickel and 7% by weight of tungsten.

The composition of the feedstock, the conditions used and the product analysis are shown in Table 3 below.

___ - TT- ie 715 0 4

Table 3

Hydrocracking over REY catalyst

Example Input 5 No._ _ _

Conditions: Temperature, ° C - 416

Pressure, kPa - 7000 LHSV, h-1 - 0.67 H2, approx. 1/1 - 1338 ^ consumption, approx. 1/1 - 143 343 ° C + conversion,% by weight - 60.4

Features:

Dry gas + C. - 3.6

Cc - 165 ° C fuel oil, wt% - 14.2 165 ° -343 ° C, distillate, wt% - 41.4 343 ° C + wt% 100 39.6 343 ° C freezing point, ° C 43 32 343 ° C + 95% TBP, ° C 540 504 343 ° C + properties:

Density, ° API 21.7

Hydrogen,% by weight 12.17 13.26

Sulfur 2.41 0.01

Nitrogen 550 33 Freezing point ° C 43 32

Rules:

Paraffins,% by volume 19 41

Naphthenes,% by volume 27 26

Aromatics,% by volume 54 33

Claims (10)

19 71 504 A process for cracking and dewaxing heavy hydrocarbon oil, characterized in that the oil is contacted with a catalyst containing zeolite beta. Process according to Claim 1, characterized in that the oil is contacted in the presence of hydrogen with a catalyst containing (i) zeolite beta as an acidic component and (ii) a hydrogenation component. Process according to Claim 2, characterized in that the zeolite beta silica / alumina ratio is more than 50: 1. Process according to Claim 2 or 3, characterized in that the hydrogenation component contains nickel, tungsten, cobalt, molybdenum or a mixture of any two or more such metals. Process according to Claim 4, characterized in that the hydrogenation component contains nickel and tungsten. Process according to Claim 2 or 3, characterized in that the hydrogenation component contains platinum, palladium, iridium, rhodium or a combination of two or more such metals. Process according to any one of Claims 1 to 6, characterized in that the initial boiling point of the oil is greater than 290 ° C. Process according to Claim 7, characterized in that the initial boiling point of the oil is above 340 ° C. Process according to Claim 8, characterized in that the boiling point of the oil is 340-565 ° C. 71504 Process according to any one of Claims 2 to 9, characterized in that the oil is contacted with the catalyst in the presence of hydrogen gas at a temperature of 230 to 500 ° C, a pressure of 500 to 20,000 kPa, a volume flow rate of 0.1 to 20 and hydrogen. at a recycling speed of 10-3,500 m / m.