EA039642B1 - Method for oligomerization of c2-c10 olefins and complex unit for producing high-octane gasolines, diesel fractions or aromatic hydrocarbons from c1-c10 hydrocarbon fractions of various composition and oxygen-containing c1-c6 compounds using it - Google Patents

Abstract

A method and various embodiments of a complex unit using it are claimed, for producing components of high-octane gasolines, diesel fractions or aromatic hydrocarbons by catalytic oligomerization of individual C2-C10 olefins or their mixtures in any ratio and/or with C1-C10 paraffins, and/or with hydrogen in presence of oxygen-containing C1-C6 compounds selected from C1-C6 alcohols and simple ethers, into gasoline hydrocarbons; it is used to bring the raw materials in contact with a zeolite catalyst for oligomerization of the following composition, wt.%.: zeolite - 40.0-80.0, promoter (promoters) - 1.0-4.0, binder component - the remaining content up to 100, containing zeolite of pentasil group (ZSM-5, ZSM-11 type) with molar ratio SiO2/Al2O3=20-80 mol/mol, and residual content of sodium oxide of not more than 0.2 wt.%, containing the binding component selected from gamma aluminum oxide, silicone dioxide, synthetic aluminosilicate, zirconium dioxide, containing the promoter (promoters) - zinc oxide, or the mixture of zinc oxide and one or several elements, or their oxides, other chemical elements, namely: ferrum(III), magnesium, calcium, gallium, cerium, lanthanum, neodymium, praseodymium, platinum, cobalt, chrome, copper, boron, fluorine, phosphorus. It is characterized in that isobutane is additionally introduced to the raw flow, in the amount (0.5-5.0):1 relative to the olefin-containing fraction (in volume), the ratio of the olefin-containing fraction to the oxygen-containing fraction (by volume) being 1.0:(0.05-0.5). The method is characterized by that the oligomerization process is carried out at temperature of 250-450°C, at pressure of 0.1-4.0 MPa, volume velocity of raw materials feed (by fluid) of 0.1-10.0 h-1 under adiabatic or isothermal conditions.

Description

The invention relates to a method for converting C2-C10 olefins into gasoline hydrocarbons, as well as to a complex installation for the production of synthetic hydrocarbon products (high-octane gasolines, diesel fractions or aromatic hydrocarbons) from C ₁ -C ₁₀ hydrocarbon fractions of various compositions, including those containing C olefins ₂ -C _1o and oxygen-containing compounds Ci-C6 (organic monohydric alcohols and / or their ethers), in particular, to obtain a component of high-octane gasoline from olefin-containing fractions and methanol, diesel fractions from olefin-containing compounds and methanol or dimethyl ether, as well as for obtaining aromatic hydrocarbons from olefin-containing raw materials and methanol or other alcohols or their esters and can be used in oil refining and petrochemistry.

Refining processes produce a large amount of gaseous mixtures of paraffins and olefins under normal conditions, the main source of which is the catalytic cracking process. One of the ways to solve the problem of using olefin-containing gases is to obtain components of motor fuels from them, in particular, components of high-octane motor gasoline or diesel fractions.

Over the past 15-20 years, a possible reduction in world oil production has been increasingly discussed. Therefore, numerous scientific and technical programs have appeared in many countries to develop and implement new technologies for the production of so-called alternative fuels, i.e. non-petroleum fuels such as biomass fuels and other energy sources. Some progress has been made, however, primarily in the production of some low quality fuel oil substitutes and gaseous fuel substitutes for combustion in the burners and nozzles of small boilers to generate heat or electricity. It should be emphasized that these substitutes are not synthetic motor (transport) fuels.

Synthetic hydrocarbons and motor (transport) fuels are commonly referred to as hydrocarbons, their fractions and mixtures that can be obtained from a variety of raw materials (coal, natural gas, oil associated or refinery gases, oil shale, peat, biomass, etc.) using Fisher synthesis processes -Tropsch (obtaining synthetic liquid fuels based on the Fischer-Tropsch synthesis), aromatization processes or upgrading processes such as hydrocracking to improve the quality of intermediate fractions and final synthetic products. The processes for producing synthetic motor fuels include the process of converting olefin-containing gases from various oil refining processes (thermal and catalytic cracking, pyrolysis, alkane dehydrogenation, the Fischer-Tropsch process, etc.) into a component of high-octane gasoline, aromatic hydrocarbons, diesel fuel or oil fractions (polyalphaolefins) on zeolite-containing catalysts based on zeolites of the pentasil group. The processes for obtaining synthetic hydrocarbon products, such as aromatic hydrocarbons, include the processes of aromatization of paraffin- and olefin-containing refinery or associated petroleum gases on zeolite-containing catalysts of the pentasil group.

Zeolites of the pentasil group (such as ZSM-5, ZSM-11 and others) are effective catalysts for the conversion of C ₂ -C ₁₀ olefins into a mixture of C ₅₊ aliphatic and aromatic hydrocarbons. At moderate oligomerization temperatures, a mixture of olefin-containing gasoline and diesel fractions, as well as oil fractions, is obtained. At elevated temperatures, the rate of the dehydrocyclization reaction increases, and the content of aromatic and paraffinic hydrocarbons increases in liquid products. The properties of the catalysts and technology features make it possible to carry out the oligomerization of olefins with high selectivity for a particular product (Oil, gas and petrochemistry abroad, 1985, No. 9, pp. 67-70).

The olefin oligomerization reaction is accompanied by the release of a large amount of heat (heating of the catalyst bed can reach 70-80°C), so the olefin-containing feedstock is diluted and the partially converted feedstock or reaction zone is cooled. As a diluent for olefins, a gasoline fraction is used in the production of diesel fuel or C ₁ -C ₄ paraffins and an inert gas (mainly hydrogen) in the production of gasolines, separated from the product stream or supplied together with the feed stream.

The conversion of C ₃ -C ₄ olefins and their mixtures with paraffins to gasoline hydrocarbons in the MOG process of Mobil Oil Corp. and The Badger Co. carried out in a fluidized bed of a zeolite-containing catalyst under conditions of moderately high temperatures, preferably at 300-500°C, and pressures, preferably up to 3.0 MPa (Oil, gas and petrochemistry abroad, 1993, No. 6, p. 98-99 , US Pat. No. 4746762, US Pat. No. 4899015).

In the MOG process under mild temperature conditions (300-400°C), a high-octane gasoline component with an octane number of 95-98 points according to the research method (r.m.) and 82-84 points according to motor method (m.m.). The typical content of hydrocarbons in this process (vol.%): paraffinic - 34, olefinic - 30, naphthenic - 8, aromatic - 28. % gasoline with an olefin content of about 20 wt.% and aromatic hydrocarbons of more than 40 wt.% with an octane rating of 98 points im.m. and 85 points by m.m. In both cases, the conversion of olefins reaches

- 1 039642 more than 90 wt.%.

C5- _C7 olefins in the CCF (catalytic cracking fluid) gasoline fraction can be used as raw materials in this process, while aromatic and olefinic hydrocarbons and, mainly, light paraffins are formed.

In another embodiment of olefin oligomerization in the Mobil Oil Corp. MOGD process. (US Pat. No. 4150062, US Pat. No. 4227992) or its Russian counterpart (US Pat. RF No. 2135547) in a stationary layer of a zeolite-containing catalyst, under mild conditions, gasoline with an olefin content of up to 90 wt.% and a diesel fraction, almost completely consisting of olefins, and the recycling of gasoline leads to the oligomerization of the olefins contained in it and an increase in the yield of a heavy product. The latter fact characterizes the process mode as extremely mild and indicates the limited acidity of the catalyst compared to the catalyst of the MOG process.

In the patent of the Russian Federation No. 2135547 during the oligomerization of the butane - butylene fraction on a zeolite-containing catalyst in a stationary catalyst bed at a temperature of 420°C, a pressure of 1.6 MPa, a space velocity of the feedstock (by liquid) of 3.4 h ^-1 , the content of olefins in liquid products was 60-66 wt.%. In the same patent, the importance of isobutane contained in the recycle of gaseous reaction products is noted: with an increase in the content of isobutane light olefins in the oligomerization feedstock, the yield of C ₅₊ hydrocarbons and the content of C5 and C6 isoolefins in them increase.

As shown above, liquid light olefin oligomerization products contain either a high amount of olefins (over 40 wt.% for the MOGD process), or more than 40 wt.% aromatics in the MOG process with a moderate content of olefins (20 wt.%). Such gasoline cannot be used as commercial gasoline, since modern requirements for motor gasoline impose restrictions on the content of both olefins and aromatic hydrocarbons and benzene.

So, the document GOST R 51866 - 2002, motor fuels, unleaded gasoline with changes No. 1-4 (EN 228-2004) limits the content of these hydrocarbons in gasoline as follows (vol.%, no more):

olefinic -18.0;

aromatic - 35.0;

benzene -1.0.

The distribution of hydrocarbons in the products of olefin oligomerization depends on the properties of the catalyst and the severity of the process. The rigidity of the regime is determined mainly by the temperature and time of contact of the raw material with the catalyst. By increasing the rigidity of the catalytic process, the rate of cracking, dehydrocyclization and, to a lesser extent, oligomerization of olefins is increased, while the content of aromatic hydrocarbons in gasoline increases and the yield of liquid hydrocarbons decreases. The same result can be obtained by changing the properties of the catalyst, its acidic and hydrogenating/dehydrogenating activity.

Another way to change the composition of liquid oligomerization products is to introduce additional substances into the composition of raw materials, namely, alkylating agents. One of these substances are oxygen-containing organic alcohols or their ethers, in particular methanol and/or dimethyl ether.

A known method (U.S. Pat. RF No. 2208624) for obtaining high-octane gasoline fractions and aromatic hydrocarbons (options) from raw materials containing arenes, and/or naphthenes, and/or olefins, and/or paraffins, boiling up to 250°C, and/or oxygen-containing organic compounds at elevated temperature and pressure in the presence of hydrogen and a zeolite-containing catalyst (pentasil-type zeolite - ZSM-5 or ZSM-11), however, specific examples for the conversion of light olefins are not given in this patent.

Closest to the proposed invention is a method for producing motor gasoline, given in the patent of the Russian Federation No. - 205°C and C ₁ -C ₄ - alcohols and / or dimethyl ether, which are mixed in the system with their separate supply (two pumps) and their volume ratio 20-90: 10-80 at 380-420°C, pressure 0 ,2-5.0 MPa and volumetric feed rate of liquid raw material 0.5-2.0 h ^-1 .

This patent also does not contain examples of the co-conversion of olefin-containing fractions and oxygen-containing compounds (C1-C ₄ alcohols and / or dimethyl ether), and, as will be shown below, the purpose of the present invention, namely, to obtain from olefin-containing hydrocarbon fractions C1-C10 gasoline that meets modern requirements for the content of olefinic and aromatic hydrocarbons.

The purpose of the present invention is a method for the oligomerization of C ₂ -C ₁₀ olefins into high-octane gasoline, the component composition of which meets modern requirements for the content of olefins (not more than 18 vol.%) and aromatics (not more than 35 vol.%), as well as increasing the yield of gasoline 2 039642 new C ₅₊ hydrocarbons from raw olefins.

This goal is achieved by the fact that in the method of catalytic oligomerization of individual C2-C10 olefins or their mixtures with each other in any ratio and/or with paraffins C1-C10 and/or with hydrogen in the presence of oxygen-containing compounds C1-C6, selected from C1-C6 alcohols or their ethers, into gasoline hydrocarbons, in which the raw material is contacted with a zeolite catalyst for oligomerization of the composition, wt.%, zeolite - 40.0-80.0, promoter (s) - 1.0-4.0, binder component - the rest up to 100, containing a zeolite of the pentasil group (type ZSM-5, ZSM-11) with a molar ratio of SiO ₂ /Al ₂ O ₃ =20-80 mol/mol and a residual sodium oxide content of not more than 0.2 wt.%, containing a binder component selected from γ-aluminum oxide, silicon dioxide, synthetic aluminosilicate, zirconium dioxide, containing a promoter (promoters) zinc oxide, or a mixture of zinc oxide and one or more elements or their oxides of other chemical elements, namely: iron (III) , magnesium, calcium, gallium, cerium, lanthanum, neodymium, praseodymium, platinum, cobalt, chromium, copper, boron, fluorine, phosphorus, isobutane is additionally introduced into the feed stream in an amount of (0.5-5.0): 1 in relation to the olefin-containing fraction (by volume ) at a ratio of olefin-containing fraction and oxygen-containing fraction (by volume) of 1.0:(0.05-0.5).

The method is characterized in that the oligomerization process is carried out under adiabatic or isothermal conditions, i. the oligomerization reaction is carried out in an adiabatic (US Pat. RF No. 2206384) or isothermal (US Pat. RF No. 65045) reactors.

The method is characterized in that the oligomerization process is carried out at a temperature of 250-450°C, a pressure of 0.1-4.0 MPa, a volumetric feed rate of raw materials (liquid) of 0.1-10.0 h ^-1 .

The catalyst used in the proposed invention can be prepared by any method and in accordance with any composition described in the patents of the Russian Federation No.

In this case, the catalyst was prepared as follows (US Pat. RF No. 2133640, example 1): 90 g of dry zeolite of the pentasil group with a molar ratio of SiO ₂ /Al ₂ O ₃ =39 in ammonium form, containing less than 0.1 wt.% oxide sodium, exchanged with 1 liter of 0.2n. zinc chloride solution at 55°C, the suspension is cooled, filtered, the zeolite on the filter is washed with distilled water, dried in an oven for 6 h at 110°C, and calcined at 500°C for 4 h (hereinafter, the drying and calcination conditions are the same). Zinc content in zeolite 1.4 wt.% 65.9 g of the resulting zeolite is mixed with 156.3 g of a wet aluminum trihydrate cake (moisture content of the aluminum trihydrate cake is 79 wt.%), the mass is mixed, evaporated and molded by extrusion. The catalyst granules are dried and calcined. The cooled catalyst granules are poured into 140 ml of a solution containing 2.2 g of iron nitrate and 1.8 g of calcium nitrate, kept at room temperature for 2 hours and evaporated to dryness. The catalyst is dried and calcined. The composition of the obtained catalyst, wt.%: zeolite (Na ₂ O content less than 0.1 wt.%) 65.0; zinc 0.9; iron 0.5; calcium 0.4, binder 33.2.

The catalyst in examples 1-10 and 13-24 was tested on the installation with an isothermal heat pipe reactor, described in US Pat. RF No. 65045. Catalyst loading into the reactor 100 cm ³ .

The catalyst in examples 11 and 12 was tested on a flow unit with a single-shelf adiabatic reactor with a catalyst load of 100 cm ³ . Analysis of raw materials and products - chromatographic.

The composition of the raw materials used (butane - butylene fraction) (wt.%): propane - 1.0; propylene - 1.0; isobutane - 8.0; n-butane - 28.0; butenes - 62.0. The amount of olefins - 63 wt.%.

The test conditions and the results obtained are given in table. 1 and 2. Examples 2, 8, 10, 12, 14, 16, 18, 20, 22 and 24 were carried out according to the prototype (isobutane was not supplied to the reactor).

In examples 7-10 varied process conditions (temperature, pressure, volumetric feed rate of raw materials).

In table. Table 3 shows the operating parameters of the catalysts and methods for their calculation. The test time in each experiment on the mode is 8 hours.

In table. 2 shows the results of testing a zeolite-containing catalyst on various types of raw materials according to the proposed invention and according to the prototype (examples 13-24).

In examples 13, 14, instead of BBF, ethylene + methane (0.5 vol.%) + ethane (1 vol.%) + H ₂ (10 vol.%) was used as raw material, 15, 16 - propane - propylene fraction (composition , wt.%: propane - 30, propylene - 70), 17, 18 - a mixture of olefins C ₅ -C ₆ (1:1 by weight).

In examples 19, 20, dimethyl ether (DME) was used instead of methanol, 21, 22 - ethanol, 23, 24 - a mixture of C ₃ -C ₅ alcohols (1:1:1 by weight).

Test conditions: temperature 310°C, pressure 1.8 MPa, volumetric feed rate of raw materials (liquid) 3.3 h ^-1 - for all types of raw materials, except for ethylene. For ethylene (examples 13 and 14) - temperature 350°C, pressure 1.0 MPa, feed space velocity 2 h ^-1 .

As can be seen from the results given in table. 1 and 2, the catalysts prepared according to the proposed invention provide the formation of gasoline hydrocarbons with a high yield, calculated on the amount of initial olefins of the feed, as well as with the required content of olefinic (less than 18 vol.%) and aromatic hydrocarbons (less than 35 vol. %). The content of benzene in all experiments did not exceed 1 vol.%.

- 3 039642

Although isobutane itself is not ultimately consumed in the process, its presence seems to facilitate alkylation reactions of the olefin molecules with each other, with the feed isobutane and methanol or other oxygen-containing compound to form a large amount of isoparaffins (see Tables 1 and 2 ).

This method of oligomerization of olefin-containing raw materials in the presence of oxygen-containing compounds and isobutane can be implemented on a complex plant for the catalytic processing of C ₁ -C ₁₀ hydrocarbon fractions of various compositions.

The claimed installation relates to complex devices for the implementation of catalytic processes for the processing of hydrocarbon fractions C1-C _{10 of} various compositions into synthetic hydrocarbon products, including olefin-containing fractions, and oxygen-containing compounds C ₁ -C ₆ (organic monohydric alcohols and / or their ethers), in particular, for the production of high-octane gasoline from low-octane gasoline and/or olefin-containing fractions and methanol, as well as for the production of aromatic hydrocarbons from paraffin- or olefin-containing feedstock and methanol.

Developed by Mobil Oil Corp. (USA) in the 1970s of catalysts based on zeolites of the pentasil group (such as ZSM-5, ZSM-11 and others), active and stable in the reactions of dehydrocyclooligomerization of aliphatic hydrocarbons and the conversion of oxygenates into hydrocarbons, gave rise to the rapid development of this direction. There are a number of methods for producing high-octane gasolines and aromatic hydrocarbons from hydrocarbon fractions and/or alcohols and their ethers. The features of the technology are associated with high thermal effects of these processes: exothermic - during the conversion of oxygenates or olefin fractions into gasoline hydrocarbons, endothermic - during dehydrocyclooligomerization or dehydrocyclization of paraffinic hydrocarbons of the feedstock. Methods for producing high-octane gasolines from aliphatic hydrocarbons and oxygenates and complex devices for their implementation solve the problems of supplying heat to the reactor necessary for chemical reactions, removing heat generated from the reactor, optimizing conditions in the reaction zone, and simplifying and reliability of the technological process.

A special position is occupied by the processes of obtaining aromatic hydrocarbons and high-octane gasoline during the joint processing of hydrocarbon fractions and methanol. In this case, the catalytic process does not have a high thermal effect, since exo- and endothermic reactions occur simultaneously and, in principle, a heat-balanced process is possible. At the same time, technical solutions for processes using one type of raw material and a high thermal effect of its conversion, for example, the use of tubular reactors with catalyst loading into pipes (RF Pat. No. 2069227, RF Pat. No. 2098173) or into the annulus (RF Pat. No. 65045 , Pat. RF No. 2429910), or a cascade of reactors with the ability to control the temperature in each reactor with heat exchange of intermediate flows in remote heat exchangers or with the introduction of cold flows of raw materials and circulating heat carriers are redundant in the joint processing of hydrocarbon fractions and methanol, lead to process complication and high capital and operating expenses.

So, the method and system for the conversion of methanol to gasoline according to US Pat. US No. 5602289 includes the implementation of the conversion of raw materials in more than two contact zones with a catalyst (a number of sequential reactors, five in the example), the volume of which is limited by the requirement of an adiabatic temperature increase in the zone of not more than 50 ° C, with the introduction of fresh raw materials and diluent into each layer - hydrocarbons not heavier than C ₅ .

In a method for producing hydrocarbons from aliphatic alcohols using several separate contact zones of the raw material with the catalyst according to US Pat. US No. 4542252 carry out temperature control, cooling the reaction stream before entering each subsequent zone using a built-in heat exchanger.

At least two reaction zones with intermediate heat exchange of the intermediate reaction flow between the zones provides for a method for producing high-octane gasoline fractions from hydrocarbon feedstock boiling up to 250°C and/or oxygen-containing organic compounds according to the method (US Pat. RF No. 2208624). Hydrocarbon gases, gasoline fraction and/or aromatic hydrocarbons, as well as a heavy fraction boiling over above 180-215°C, are separated from the product stream by the separation and rectification method.

In the method for co-processing low-octane hydrocarbon fractions and aliphatic alcohols and/or dimethyl ether (U.S. Pat. RF No. 2429910), an isothermal reactor with heat pipes is used, which allows contact of the raw material with the catalyst under close to optimal conditions, but has all the disadvantages of a tubular reactor, including the complexity of loading the catalyst, which is especially undesirable with a relatively short catalyst life.

Close to isothermal mode of conversion of raw materials can be obtained in the process of converting methanol to olefins or gasoline according to the method (U.S. Pat. No. 5191142) in a homogeneous reaction mixture, which is a mixture of methanol and olefin in a reactor with a fluidized bed of a catalyst or in a reactor with a fixed bed with reverse mixing.

Installation of catalytic production of high-octane gasoline fractions and aromatic hydrocarbons according to US Pat. RF No. 2098173 includes a reactor block with a shell-and-tube type reactor, in

- 4 039642 tubes of which are loaded with a catalyst, and a gas coolant with a given temperature is supplied to the annular space, which is obtained by mixing air with flue gases generated in the heat generator during the combustion of fuel gas. The coolant after the reactor is used to heat the feed streams in a tubular heat exchanger before being fed into the tubular reactor. The installation includes two reactor blocks operating alternately, as well as technologically connected separators, distillation columns, heat exchange, capacitance, and pumping equipment.

Integrated installation for the production of high-octane gasoline from gasoline fractions and methanol according to US Pat. RF No. 138334 (prototype) includes reactors, technologically connected with them heating, heat exchange, separating, capacitive and pressure equipment for heating raw materials, cooling, partial condensation, separation and rectification of reaction products, differs in that each reactor includes at least two stationary catalyst beds, with the possibility of feeding into the raw mixture, as well as into the second and each subsequent layer of the part of the gas heated in the fire heater, separated in the three-phase separator from the flow of reaction products after their partial condensation.

We offer a complex installation for the production of synthetic hydrocarbon products (high-octane gasolines, diesel fractions or aromatic hydrocarbons) from hydrocarbon fractions C ₁ -C _{10 of} various compositions, including olefin-containing fractions C1-C10, and oxygen-containing compounds C ₁ -C ₆ (organic monatomic alcohols and/or their ethers), in particular, for the production of high-octane gasoline from low-octane gasoline and/or olefin-containing fractions and methanol, combines well-known approaches to solve the technical problem - providing a preferred temperature regime in reactors and can be used for simple and reliable implementation a method for producing high-octane components of gasoline, for example, according to US Pat. US No. 5019663, or US Pat. RF No. 2103322, 2372988.

The complex installation according to the present invention is intended for the production of synthetic hydrocarbon products (high-octane gasolines, diesel fractions or aromatic hydrocarbons) from a mixture of C ₁ -C ₁₀ hydrocarbons of various compositions (low-octane gasoline fractions n.k. 180°C, 90-160°C or more narrow fractions, pentane - heptane (hexane) fractions, propane - butane fractions, NGL - broad fraction of light hydrocarbons - a product of gas processing plants, and / or C ₂ -C ₁₀ olefins and / or their mixtures with each other and / or with C paraffins ₁ -C ₁₀ and / or with hydrogen) in the presence of oxygen-containing compounds (alcohols C ₁ -C ₆ and / or their ethers), including one or more parallel sectioned adiabatic reactors, consisting of one or more stationary layers (sections) of zeolite-containing catalyst with supply or removal of heat between the layers (sections) of the catalyst, or one or more parallel isothermal reactors with thermal logs, and / or coils, and / or pipe heat exchangers, and / or panels with heat supply or removal with a zeolite-containing catalyst with the possibility of feeding into the raw mixture, as well as into the second and each subsequent layer (section) with a zeolite-containing catalyst in an adiabatic reactor part of the gas (H ₂ and C ₁ -C ₄ hydrocarbons) heated in a fire or electric heater, separated in a three-phase separator from the flow of reaction products after their partial condensation, in order to circulate it through a catalyst for supplying or removing heat in an adiabatic reactor, converting those contained in of unsaturated hydrocarbons and an increase in the regeneration run of the catalyst, heating, heat exchange, separating, capacitive, injection and mass transfer equipment technologically connected with the reactor for heating the raw material, cooling, partial condensation, separation and rectification of the reaction products, while isobutane is also used as a raw material stream, which is preliminary Relatively heated can be supplied alone and / or in a mixture with oxygen-containing compounds C ₁ -C ₆ and / or in a mixture with olefin-containing raw materials containing olefins C ₂ -C ₁₀ as an inlet to the reaction part of an isothermal reactor or to the first layer (section) of the adiabatic reactor, and in the second and / or each subsequent layer (section) of the adiabatic reactor in the following volumetric ratio of components: hydrocarbons C ₁ -C ₁₀ : oxygen-containing compounds C ₁ -C ₆ - 1: (0.01-10 ,0); hydrocarbons C ₁ -C ₁₀ : isobutane - 1:(0.01-5.0).

The plant differs in that it includes a block consisting of at least one reactor for removing sulfur compounds using hydrogen, including its own hydrogen-containing gas formed in the process, from at least part of the C ₁ -C ₁₀ hydrocarbon feedstock, into which the heated feedstock is sent , a stream of raw materials purified from sulfur is removed from the reactor and sent to the raw material heater of the isothermal reactor or the first layer (section) of the adiabatic reactor.

The installation differs in that it includes a distillation column for separating C ₅₊ components from hydrocarbon feedstock (NGL) in the case of aromatization of this feedstock, and a vapor-phase stream containing C ₁ -C ₄ hydrocarbons is sent to the feed heater of the isothermal reactor or the first layer (section) of the adiabatic reactor top of the column after reflux release.

The installation differs in that fractions C ₃ , C4, C ₃ -C ₄ , C ₅ -C ₆ or C ₅ -C ₇ are separated from NGL by rectification, and these fractions are sent separately from each other for aromatization to different isothermal reactors or to different shelves (sections) with zeolite-containing catalyst in an adiabatic reactor

- 5 039642 under different temperature conditions.

The installation differs in that it includes at least one adsorber filled with an adsorbent and selectively adsorbing C ₅₊ hydrocarbons, including aromatic hydrocarbons, into which the vapor phase stream separated in the three-phase separator from the product stream is directed, and the vapor phase depleted in C ₅₊ - hydrocarbons is removed a stream containing hydrogen and hydrocarbons Ο ₁ -€ ₄ .

The installation is characterized in that a sectioned adiabatic reactor, consisting of two or more layers (sections) of the catalyst, has an increasing volume distribution of the catalyst over the layers (sections) from the first to the last, namely, 1:2:3:4:6 and so on. .

The plant differs in that it has a separate independent circulation circuit for regeneration of the zeolite-containing catalyst with a circulating compressor, return flow heat exchange, separation equipment, a fire or electric heater for circulating a nitrogen-air mixture through the reactors with nitrogen and air feed with a mixture flow rate of 60-1800 nm ³ / m ³ catalyst and blowing regeneration gases.

The installation is characterized in that to increase the selectivity of the process for liquid products (aromatic hydrocarbons or diesel fraction) in the feed mixture (after the desulfurization reactor) of the adiabatic or isothermal reactors, as well as in the second and / or each subsequent layer (section) of the zeolite-containing catalyst in the adiabatic reactor for the purpose of their circulation through the catalyst, unreacted components of the liquid feedstock heated in a fire or electric heater (non-aromatic hydrocarbons in the case of the aromatization of hydrocarbons C1-C ₁₀ ) or non-target liquid products (olefin-containing gasoline fraction in the case of obtaining a diesel fraction from olefin-containing feedstock) separated from the stream are fed through the catalyst products by absorption and/or adsorption and/or extractive distillation (distillation) (in the first case) or rectification (in the second case).

The installation is characterized by the fact that the process of processing raw materials (hydrocarbons C1-C ₁₀ , oxygen-containing compounds C1-C6 and isobutane) is carried out in the presence of catalysts based on zeolites of the pentasil group in hydrogen form (such as ZSM-5, ZSM-11 and others) with a molar ratio SiO ₂ /Al ₂ O ₃ \u003d 20-80 mol / mol and the content of Na ₂ O is not more than 0.2 wt.% at a temperature of 250-600 ° C, a pressure of 0.1-10.0 MPa, a volumetric feed rate of raw materials ( by liquid) 0.1-10 h' ¹ .

The raw materials of the process are straight-run oil or gas condensate low-octane gasoline fractions, preferably boiling up to 180 ° C, propane-butane fractions, NGL, and / or olefin-containing fractions, as well as isobutane and oxygen-containing organic compounds (C1-C6 alcohols or their ethers), for example methanol of any degree of purification, including raw methanol or dimethyl ether.

A complex installation for ensuring continuous processing of raw materials may include two or more reactors operating alternately in the reaction-regeneration cycle. Use isothermal or shelf adiabatic reactors with switchgear at least between two layers of catalyst to enter the circulating gas - coolant. Known catalysts based on zeolites of the pentasil group in hydrogen form (Na ₂ O content is not more than 0.2 wt.%) such as ZSM-5, ZSM-11 and others can be used, namely, catalysts for paraffin dehydrocyclization, olefin oligomerization, methanol conversion or dimethyl ether into gasoline hydrocarbons, in the presence of which, at a temperature of 250-600 ° C, a pressure of 0.1-10.0 MPa, a feed space velocity (for liquid feedstock) of 0.1-10 h ^-1 , predominantly gasoline hydrocarbons are formed from paraffin- or olefin-containing raw materials, from alcohols, their ethers or their mixtures, aromatic hydrocarbons - from aliphatic hydrocarbons and oxygen-containing raw materials, diesel fraction - from olefin-containing raw materials, and cracking, alkylation and isomerization of hydrocarbons take place. Catalysts prepared according to RF patents No. 2087191, 2098455, 2100075, 2133640, 2165293, 2169043, 2284343, 2440189 or EAPO patent No. 002139 can be used.

The nitrogen-air oxidative regeneration of the catalyst for the conversion of hydrocarbon feedstock in the reactors is provided by the regeneration circuit.

The plant includes technologically connected recuperation heat exchangers and a fire heater for heating the feedstock, a recuperation heat exchanger and a fire heater of the recycle supplied both to the reactor inlet through the feedstock line or through its own line, and into the second and subsequent catalyst layers in the adiabatic reactor, as well as air and a water cooler for condensing the components of the product stream, a three-phase separator for separating uncondensed components from it, mainly hydrocarbons C1 and C2, an aqueous phase containing traces of hydrocarbons and oxygen-containing compounds, and a liquid unstable catalyzate, which is stabilized in a distillation column to obtain a stable high-octane gasoline or concentrate aromatic hydrocarbons (mainly benzene - toluene - xylene fraction) or stable gasoline and diesel fraction. If necessary, the gasoline fraction or BTX fraction can be separated into aromatic and non-aromatic hydrocarbons by absorption and/or adsorption and/or extractive rectification (distillation) methods. In the case of targeted production of diesel

- 6 039642 fractions from olefin-containing raw materials, the resulting gasoline olefin-containing fraction is sent for recycling from a distillation column, and in the case of targeted production of aromatic hydrocarbons, the resulting non-aromatic hydrocarbons are sent for recycling after separation from aromatic hydrocarbons by absorption and/or adsorption and/or extractive rectification methods ( distillation). After separating non-aromatic and aromatic hydrocarbons, the latter can be separated in a separate distillation column into individual hydrocarbons (benzene, toluene, xylenes, ethylbenzene, C ₉ aromatics, naphthalene, and so on).

If the paraffin-containing raw material (propane-butane fraction) needs to be subjected to full aromatization and it contains C ₅₊ (NGL) components, they are separated in a distillation column. In this case, the installation additionally includes a distillation column and the equipment necessary for its operation: a raw material heater, a reboiler or an evaporator for heating the bottom of the column, a refrigerator, as well as a reflux tank and a pump for supplying to the cold reflux column. The vapor-phase stream containing hydrocarbons C1-C4 is sent for processing from the top of the column after reflux separation.

If it is necessary to aromatize propane, butane, pentane-hexane or pentane-heptane fractions separately, then C ₃ , C ₄ , C ₃ -C ₄ , C ₅ -C ₆ or C5-C7 fractions are isolated from NGL by distillation and these fractions are sent separately for aromatization (to different isothermal reactors operating under different temperature conditions, or to different shelves of an adiabatic reactor).

If the raw material is represented by separate fractions of saturated hydrocarbons and olefin-containing fractions, the latter can be fed separately to each layer (section) of the zeolite-containing catalyst in the adiabatic reactor. This also applies to the supply of isobutane and oxygenated compounds. Such a supply of raw materials in an adiabatic reactor makes it possible to obtain conditions for the course of a chemical process that are close to isothermal. Preferably, the olefin- and oxygen-containing feedstock is mixed with the paraffin- and isobutane-containing feedstock before each layer (section) to form a mixture of hydrocarbons, 30-40% by weight of which is olefins and oxygenated compounds.

For a more complete extraction of hydrocarbons С ₅₊ from gaseous hydrocarbons С1-С ₄ separated in a three-phase separator, some of which is subsequently sent for mixing with the raw hydrocarbon stream of the installation (after the desulfurization reactor), without the use of high pressure and deep cold, the installation may include at least one or more adsorbers - to ensure continuous purification of the vapor stream separated from the product stream (U.S. Pat. RF No. 2277527, 57278). For adsorption of hydrocarbons C ₅₊ from a mixture of hydrocarbons C1-C4 known adsorbents can be used: activated carbon, silica gel, zeolites. It is preferable to use activated carbon as an easily regenerated sorbent with a high adsorption capacity. Adsorption is carried out in conventional adsorbers. The conditions for the adsorption of C5+ hydrocarbons depend on the properties of the sorbent used and, in the preferred case, coincide with the characteristics of the vapor phase stream being purified.

The vapor-phase stream from the separator, if necessary, is sent to the adsorber located after the compressor on the line of circulation of gaseous reaction products, in the preferred case without preliminary preparation, and the vapor-phase stream depleted in C ₅₊ hydrocarbons containing hydrogen and C1-C ₄ hydrocarbons is removed from the adsorber. The degree of extraction of hydrocarbons WITH ₅₊ from the vapor stream in the preferred case is not lower than 95% for a complete cycle of adsorption.

Especially needed block adsorption of hydrocarbons C ₅₊ (primarily benzene and toluene) from hydrocarbons C1-C ₄ in the process of obtaining aromatic hydrocarbons.

The adsorbent saturated with C5+ hydrocarbons is regenerated by raising the temperature in the adsorber while simultaneously purging the adsorbent with a part of the C5+-hydrocarbon depleted stream from another adsorber (stripping gas) or with an inert gas from the side. Desorbed C5+ hydrocarbons are recovered from the saturated stripping gas with a smaller volume of non-condensable components than in the adsorption stream, and their condensation and separation can be carried out with less loss.

In the regeneration mode, the adsorber is supplied with a heated stripping gas - a part of a stream depleted in C ₅₊ -hydrocarbons from another adsorber or an inert gas (for example, nitrogen) if there is only one adsorber in the circuit, and stripping gas saturated with C ₅₊ hydrocarbons is removed. To regenerate the adsorbent, the installation includes a heater for heating the stripping or inert gas and heat exchangers and refrigerators connected in series with the outlet of the adsorber for cooling and condensing C5+ hydrocarbons from a saturated stripping or inert gas, as well as a separator for separating the vapor-liquid mixture, from which a liquid-phase stream containing C ₅₊ hydrocarbons, which is sent to a distillation column for stabilization, and a vapor phase stream, which is mixed with a C 5+ depleted hydrocarbon stream from another adsorber, or an inert gas.

Adsorbers and adsorbent regeneration equipment form an adsorption unit. The input stream containing C5+ hydrocarbons is a vapor phase stream from a three-phase reactor product stream separator, the output streams are a vapor phase stream depleted in C ₅₊ hydrocarbons, or an inert gas and a liquid phase stream containing C ₅₊ hydrocarbons.

- 7 039642

If there is no need to clean the vapor-phase flow from the three-phase separator from C ₅₊ hydrocarbons, a part of this flow is blown into the fuel network on the line before the circulating compressor.

An unstable liquid product from a three-phase separator through heat exchangers, where it is heated by the heat of a stable product, enters a distillation column for stabilization. Vapors are removed from the top of the column, which are cooled and condensed in an air cooler and a water cooler, and then uncondensed gas is released in the separator, which is sent to the fuel network, and liquefied gas, part of which is pumped to the column as irrigation, and the balance part is removed from plant as a commercial product (in some cases this gas is used as plant feedstock). The bottom product of the column enters the reboiler, where light fractions are stripped from it, which then go under the bottom plate of the column, and a stable liquid product (high-octane gasoline, aromatic hydrocarbon concentrate or diesel fraction in the case of olefin-containing feedstock oligomerization - in this case, the gasoline olefin-containing fraction is removed from side stream and sent for recycling for further processing, not shown in the diagrams) is cooled in the annulus of the heat exchanger and delivered to the warehouse (in the case of obtaining high-octane gasoline) or for further separation of aromatic and non-aromatic hydrocarbons (in the case of obtaining a concentrate of aromatic hydrocarbons) using standard methods in their separation block.

In the latter case, aromatic and non-aromatic hydrocarbons (usually not more than 1.0 wt.%) are separated in block A-2 by absorption, and / or adsorption, and / or extractive rectification (distillation) methods (Sulimov A.D. Production of aromatic hydrocarbons from oil raw materials, Moscow, Chemistry, 1975, pp. 36-72). The content of non-aromatic hydrocarbons in the composition of aromatic hydrocarbons after the separation unit does not exceed 1000 ppm. After the separation unit, aromatic hydrocarbons can be separated in a separate distillation column (block) K-3 into individual hydrocarbons (benzene, toluene, xylenes, C ₉ aromatics, naphthalene, etc.), and liquid non-aromatic hydrocarbons are removed from the unit as a product.

In FIG. 1 (with two shelf four-section adiabatic reactors), fig. 2 (with two isothermal reactors for carrying out exothermic reactions) and FIG. 3 (with two isothermal reactors for carrying out endothermic reactions) schematic diagrams of a complex installation for the production of high-octane gasoline, diesel fraction or aromatic hydrocarbons from C ₁ -C ₁₀ hydrocarbons of various compositions, oxygen-containing compounds and isobutane are presented. The installation includes the following devices: E-1, E-2 - raw materials tanks; N-1, N-2, N-3, N-4, N-5 - pumps; P-1, P-2 - ovens; P-3 - fire or electric heater; R-1, R-2 - synthesis reactors; T-1, T-2, T-3, T-4, T-5, T-6, T-7 - heat exchangers; PC-1, PC-2 - compressors; XV-1, XV-2, XV-3 - air coolers; X-1, X-2 - water coolers; С-1 - three-phase separator, С-2, С-3, С-4 - separators; RB - reboiler; K-1 - distillation column. The diagrams also show: R-3 - desulfurization reactor; K-2 - distillation column for the separation of NGL into fractions (C ₁ -C ₂ , C ₃ , C ₄ , C ₃ -C ₄ , C ₅ -C ₆ , C ₅ -C ₇ , C ₅ -C ₁₁ , C ₁₂₊ ); A-1 - block adsorption of hydrocarbons With ₅₊ from hydrocarbons With ₁ -C ₄ ; A-2 - unit for separating aromatic and non-aromatic hydrocarbons; K-3 - unit for the separation of aromatic hydrocarbons. As necessary, based on the nature of the feedstock, these blocks may or may not be involved in the operation of a complex installation. The diagrams indicate the flows: I - raw hydrocarbon fraction C ₁ -C ₁₀ , II - raw oxygen-containing fraction C ₁ -C ₆ , III - raw mixture, IV products of the catalytic process, V - hydrogen-containing gas, VI - water condensate, VII - unstable gasoline, VIII - liquefied gas, IX - stable product (high-octane gasoline or diesel olefin-containing fraction or concentrate of aromatic hydrocarbons or individual aromatic hydrocarbons), X - C5 + non-aromatic hydrocarbons - component of the gasoline fraction, XI - C5 + hydrocarbons.

The installation works as follows. Raw materials - raw hydrocarbon fraction C ₁ -C ₁₀ I (straight-run oil or gas condensate gasoline fraction or pentane-hexane (heptane) fraction or propane-butane fraction or NGL (or part of it) and / or olefin-containing fraction) and isobutane - from the raw tank E-1 enters the intake of the raw pump H-2, where the raw oxygen-containing fraction C ₁ -C ₆ II is supplied by the pump H-1 from the raw tank E-2. WFLH can be previously separated in the K-2 distillation column into fractions C ₁ -C ₄ , C ₅ -C ₇ , C ₅ -C ₆ , C ₅₊ , C ₆₊ or C ₇₊ , and as a raw material for the installation it can be any of these fractions is used. If necessary, the scheme may include a desulfurization reactor (R-3) for raw gasoline, pentane-hexane (heptane), olefin-containing fractions, NGL, as well as any fraction separated from NGL, using the hydrogen-containing gas obtained in the process. The raw mixture passes through the tube space of the T-3 and T-1 heat exchangers, where it is heated by the gas-product stream coming from the R-1 or R-2 reactor operating in the synthesis mode. Next, the feed stream III is heated in the P-1 furnace to the required reaction start temperature (250-510°C, depending on the feedstock used) and enters the R-1 or R-2 reactor operating in the synthesis mode. Simultaneously into the first or each subsequent catalyst layer or between these layers in the adiabatic reactor, as well as at the inlet to the isothermal

The PK-1 circulating compressor supplies the reactor with gas heated in the T-2 recuperation heat exchanger and P-3 furnace to 410-480°C, as well as a hot recycle stream from K-1 (not shown in the diagrams). In the reactor, depending on the raw material used, at a temperature of 250-510°C and a pressure of 0.1-10.0 MPa, a catalytic process of raw material transformation is carried out. The product stream IV is removed from the reactor, which sequentially passes through the annular space of the heat exchangers T-1, T-2 and T-3, where it gives off heat to the feed stream and circulating gas, then the T-4 heat exchanger, where it gives off heat to the unstable product flow, then cools in air cooler XB-1 to a temperature of 55°C and then - recycled water in a water cooler X-1 to a temperature of 35°C. Further, the gas-liquid mixture of products enters the three-phase separator C-1, where it is separated into a gas stream V, an aqueous condensate VI and an unstable liquid product VII.

The main part of the gas stream separated in the three-phase separator enters the PK-1 circulation compressor unit and at a pressure of 1.9-2.0 MPa (in the preferred case when producing high-octane gasolines) and 4.0-7.0 MPa (in the preferred case at obtaining diesel fractions from olefin-containing raw materials) enters the coil of the T-2 heat exchanger, then to the P-3 furnace and then to the R-1 or R-2 reactor operating in the synthesis mode.

Water condensate containing acidic components and unreacted alcohol (ether) (usually traces) is sent to a water treatment system or to an oil desalination unit.

The unstable product VII (gasoline or diesel fraction, or aromatic hydrocarbon concentrate) is fed by the H-3 pump to the coil of the T-4 heat exchanger, then to the coil of the T5 heat exchanger, where it is heated by the heat of the stable product, and then enters the K-1 distillation column for stabilization . Vapors are removed from the top of the column, which are cooled and condensed in the XV-2 air cooler and X-2 water cooler, and then in the S-2 separator, uncondensed gas is released, which is sent to the fuel network, and liquefied gas, part of which is pumped by the H-4 pump sent to the column as irrigation, and the balance part VIII is removed from the installation. The bottom product of the column enters the RB reboiler, where light fractions are stripped from it, which then go under the lower plate of the column, and the stable product IX - high-octane gasoline or diesel fraction - is cooled in the annulus of the T-5 heat exchanger and enters the warehouse. If the target stable product of the K-1 column is the diesel fraction obtained by oligomerization of the olefin-containing feedstock, then the olefin-containing gasoline fraction formed in the process is removed by the side stream of the K-1 column and is recycled to the R-1 or R-2 reactor (not shown in the diagrams) .

If the stable product of the column is a concentrate of aromatic hydrocarbons in the case of aromatization of hydrocarbon feedstock, then further processing of this product occurs as follows.

Aromatic and non-aromatic (usually not more than 1.0 wt.%) hydrocarbons are separated by absorption and/or adsorption and/or extractive distillation (distillation). The content of non-aromatic hydrocarbons in the composition of aromatic hydrocarbons after the separation unit does not exceed 1000 ppm. After the separation unit, aromatic hydrocarbons can be separated in the distillation column (block) K-3 into individual hydrocarbons (benzene, toluene, xylenes, C9 aromatics, naphthalene, etc.), and liquid non-aromatic hydrocarbons are removed from the unit as a commodity, as well as in some cases, they can be recycled to the aromatization process.

As the catalyst coking in the synthesis reactor P-1 or P-2, the temperature of the incoming feed stream rises. After reaching the maximum processing temperature for each type of feedstock, the reactor with the catalyst that has lost activity is disconnected from the feed stream and switched to the oxidative regeneration mode with a nitrogen-air mixture through the regeneration circuit. The scheme for supplying nitrogen and air to the C-4 separator and the nitrogen-air mixture line from C-4 to the intake of the PK-2 circulation compressor are launched. At the beginning of regeneration, the reactor is purged of combustible compounds with pure nitrogen. Nitrogen after the PK-2 circulation compressor, the tube space of the T-6 and T-7 heat exchangers through the P-2 furnace enters the reactor with a coked catalyst. After leaving the reactor, nitrogen, and subsequently the nitrogen-air mixture, passes through the annulus of the T-7 and T-6 heat exchangers, where it releases heat to the fresh nitrogen-air mixture. Next, nitrogen (nitrogen-air mixture) through the XB-3 cooler enters the C-3 separator, where part of it is blown off to the candle, and the other part through the C4 separator enters the regeneration circuit again. After lowering the temperature in the reactor to 280°C and the content of combustibles less than 0.5 vol.% begin a gradual dosage of air (up to 100%) into nitrogen in the C-4 separator and a gradual rise in temperature in the furnace P-2 to 510°C. The duration of the regeneration cycle is up to 120 hours.

The proposed complex installation can be illustrated by the following examples (examples 25-29). Test time on the mode - 8 hours.

Example 25.

According to the described technological scheme (Fig. 1) under the following conditions: temperature 380°C, pressure 0.8 MPa, volumetric feed rate of raw materials (in terms of liquid) 1.5 h ^-1 out of 14800.0 g/h of the gasoline fraction boiling within 38-165°C, 200.0 g/h of isobutane and 3750.0 g/h of methanol in contact with a zeolite-containing catalyst in four zones of an adiabatic reactor at a temperature at the exit from each zone of only

- 9 039642 4-5°C below the temperature at the entrance to the zone receive 14550.0 g/h of stable gasoline with a benzene content of 0.85 vol.% and an octane number of 94 points according to the research method, 1391.0 g/h of the fraction

With _s -C ₄ , 2109.4 g/h of water condensate and 699.6 g/h of gas purge sent to the fuel network.

Example 26.

According to the described technological scheme (Fig. 1) under the following conditions: temperature 300°C, pressure 1.8 MPa, volumetric feed rate of raw materials (liquid) 2 h ^-1 of 14600.0 g/h butane - butylene fraction (composition, wt %: propane - 1.0; propylene - 1.0; isobutane - 8.0; n-butane - 28.0; butenes - 62.0; total olefins - 63.0), 400.0 g/h of isobutane and 3750.0 g/h of dimethyl ether in contact with a zeolite-containing catalyst in four zones of an adiabatic reactor at a temperature at the outlet of each zone only 5-8°C higher than the temperature at the inlet to the zone, 12020.0 g/h of stable gasoline containing benzene less than 0.1 vol.% and an octane number of 97.5 points according to the research method, 3427.0 g / h of the C3-C4 fraction, 1467.4 g / h of water condensate and 1835.6 g / h of gas blowing sent to the fuel network.

Example 27.

According to the described technological scheme (Fig. 2) under the following conditions: temperature 260°C, pressure 4.2 MPa, volumetric feed rate of raw materials (in terms of liquid) 1 h ^-1 of 14900.0 g/h propane - propylene fraction (composition, wt %: propane - 30.0, propylene - 70.0), 100.0 g/h of isobutane and 3750.0 g/h of a mixture of methanol and dimethyl ether (1:1 by weight) in contact with a zeolite-containing catalyst in an isothermal reactor with heat removal from the reaction space at a temperature at the outlet of the reactor only 2-3°C higher than the temperature at the inlet to the reactor, 12480.0 g/h of diesel fraction (170-270°C) with a cetane number of 34.0 are obtained (before hydrotreatment ) and 53.0 (after hydrotreating), 3350.0 g/h of gasoline fraction (38-170 _° C), _920.0 g/h of C3-C4 fraction, 1790.4 g/h of water condensate and 209, 6 g/h of gas purge directed to the fuel network.

Since in this example the target stable product of the K-1 column is the diesel fraction obtained by oligomerization of the olefin-containing feedstock (propane-propylene fraction), the gasoline fraction formed in the process is removed by the side stream of the K-1 column and, after heating, is recycled to the R- reactor. 1 or P-2 (not shown in the diagram) or can be displayed as a product.

Example 28.

According to the described technological scheme (Fig. 3) under the following conditions: temperature 520°C, pressure 0.8 MPa, volumetric feed rate of raw materials (liquid) 1 h ^-1 out of 11000.0 g/h propane-butane fraction (composition, wt .%: propane - 50.0, butane - 50.0), 3000.0 g / h of isobutane and 4750.0 g / h of a mixture of alcohols C ₂ -C ₄ (ethanol: isopropanol: butanol - 1: 1: 1 according to mass) in contact with a zeolite-containing catalyst in an isothermal reactor with heat supplied to the reaction space at a reactor outlet temperature only 3-4°C lower than the reactor inlet temperature, 11550.0 g/h of aromatic hydrocarbon concentrate, 320.0 g /h of the C ₃ -C ₄ fraction, 1425.0 g/h of water condensate and 5455.0 g/h of gas purge sent to the fuel network.

Since the target product of this example is an aromatic hydrocarbon concentrate, the typical scheme (Fig. 3) is being finalized as follows:

the vapor-phase flow from the three-phase separator C-1 is sent to the adsorber located after the compressor on the circulation line of the gaseous reaction products. From the adsorber, a vapor-phase stream depleted in C ₅₊ hydrocarbons (mainly aromatic hydrocarbons, in particular, benzene and toluene) containing hydrogen and C1-C4 hydrocarbons is removed, which is recycled to the R-1 or R-2 reactor. The degree of extraction of hydrocarbons C ₅₊ from the vapor stream in the preferred case is not lower than 95% for the full cycle of adsorption;

aromatic and non-aromatic (usually not more than 1.0 wt.%) hydrocarbons (stream IX from column K-1) are separated by absorption and / or adsorption and / or extractive distillation (distillation) in the separation unit of aromatic and non-aromatic hydrocarbons (A -2). The content of non-aromatic hydrocarbons in the composition of aromatic hydrocarbons after the separation unit does not exceed 1000 ppm. After the separation unit, aromatic hydrocarbons can be separated in the distillation column K-3 into individual hydrocarbons (benzene, toluene, xylenes, C ₉ aromatics, naphthalene, etc.), and liquid non-aromatic hydrocarbons are removed from the unit as a commodity.

In this example, from the formed 11550.0 g/h of the concentrate of aromatic hydrocarbons in the column K-3 can be obtained, wt.%: benzene - 20-24; toluene - 32-34; sum of xylenes - 24-27; ethylbenzene 2-3; aromatic C ₉ - 3-4; naphthalene - 5-7; methylnaphthalenes - 5-7; the rest is 1.

Example 29.

According to the described technological scheme (Fig. 3) under the following conditions: temperature 500°C, pressure 0.5 MPa, volumetric feed rate of raw materials (liquid) 2 h ^-1 out of 14500.0 g/h pentane - hexane fraction (composition, wt %: pentane - 50.0, hexane - 50.0) isolated from NGL, 500.0 g/h of isobutane and 3750.0 g/h of methanol in contact with a zeolite-containing catalyst in an isothermal reactor with heat supply to the reaction space at temperature at the outlet of the reactor only 2-3°C lower than the temperature at the inlet to the reactor, 12300.0 g/h of aromatic hydrocarbon concentrate, 1100.0 g/h of the C ₃ -C ₄ fraction, 2109.4 g/h of water condensate are obtained and 3240.6 g/h of gas purge sent to the fuel network.

Since the raw material of this example is NGL, and the target product of the process is the concentrate of aromatic hydrocarbons, then the typical scheme (Fig. 3) is being finalized as follows:

if NGL contains sulfur compounds, then the unit includes a block (R-3), consisting of at least one reactor, for removing sulfur compounds using hydrogen, including its own hydrogen-containing gas formed in the process, from at least part of the hydrocarbon feedstock Ci -Su, into which the heated raw material is directed, the flow of the raw material purified from sulfur is removed from the reactor and sent to the raw material heater of the isothermal reactor. In this case, it is preferable to purify the entire NGL stream so that all further NGL processing products are purified from sulfur compounds;

the pentane-hexane fraction, which must be subjected to full aromatization, is isolated as a side stream in the K-2 distillation column, and this fraction is sent to the R-1 or R-2 reactor for aromatization.

The purification of the exhaust gas from the reactor and the purification and separation of the obtained aromatic hydrocarbons are carried out in the same way as in example 28.

The vapor-phase flow from the three-phase separator C-1 is directed to the adsorber located after the compressor on the circulation line of the gaseous reaction products. From the adsorber, a vapor-phase stream depleted in C ₅₊ hydrocarbons (mainly aromatic hydrocarbons, in particular benzene and toluene) containing hydrogen and C ₁ -C ₄ hydrocarbons is withdrawn, which is recycled to the R-1 or R-2 reactor. The degree of extraction of hydrocarbons With ₅₊ from the vapor stream in the preferred case is not less than 95 wt.% for a full cycle of adsorption.

Aromatic and non-aromatic (usually not more than 1.0 wt.%) hydrocarbons (stream IX from column K-1) are separated by absorption and / or adsorption and / or extractive rectification (distillation) in the separation unit of aromatic and non-aromatic hydrocarbons (A -2). The content of non-aromatic hydrocarbons in the composition of aromatic hydrocarbons after the separation unit does not exceed 1000 ppm. After the separation unit, aromatic hydrocarbons can be separated in a distillation column (block) K-3 into individual hydrocarbons (benzene, toluene, xylenes, C ₉ aromatics, naphthalene, etc.). Liquid non-aromatic hydrocarbons are removed from the plant as a commodity.

In this example, from the formed 12300.0 g/h of the concentrate of aromatic hydrocarbons in the column (block) K-3 can be obtained (wt.%): benzene 18-25; toluene 30-40; sum of xylenes 14-22; ethylbenzene 1-2; aromatic C ₉ 4-5; naphthalene 2-5; methylnaphthalenes 3-6; the rest 1.

The plant described above is suitable for processing an olefin-containing feed mixture containing C ₂ -C ₁₀ olefins and/or hydrogen and/or C ₁ -C ₁₀ paraffins and/or isobutane and/or oxygen-containing compounds (C1 alcohols -C6 or their ethers), into high-octane gasoline containing in its composition no more than 18 vol.% olefinic hydrocarbons and no more than 35 vol.% aromatic hydrocarbons (i.e. gasoline that meets the requirements of GOST R 51866 - 2002 for motor gasolines ), both in the presence of isobutane and oxygen-containing compounds (C1-C6 alcohols or their ethers), and in their absence.

In this case, this installation, which includes a sectioned adiabatic conversion reactor, consisting of two or more stationary layers (sections) with a zeolite-containing catalyst with heat removal between the layers (sections) of the catalyst, with the possibility of feeding into the raw mixture, as well as into the second and each a subsequent layer (section) with a zeolite-containing catalyst of the part of the gas (H ₂ , C ₁ -C ₄ hydrocarbons, oxygen-containing compounds) heated in a fire or electric heater, separated in a three-phase separator from the reaction product stream after their partial condensation, in order to circulate it through the catalyst for supply or removal of heat in the reactor, conversion of unsaturated hydrocarbons contained in it and increase in the inter-regeneration run of the catalyst, heating, heat exchange, separating, capacitive, injection and mass transfer equipment technologically connected with the reactor for heating raw materials, cooling, partial condensation, separation and rectification of products reaction, while in the first section (layer, shelf) in the direction of movement of the raw mixture, or in the first and second, or in the first, second and third, a zeolite-containing catalyst for the oligomerization of olefin-containing components of the raw material is loaded, and in the last (last) section in the direction of movement of the raw mixture (layer, shelf) a zeolite-containing catalyst for aromatization of aliphatic hydrocarbons is loaded, and sections (layers, shelves) with zeolite-containing oligomerization and aromatization catalysts can alternate with each other, starting with a section with an oligomerization catalyst.

This plant is different in that the layers (sections) with the aromatization catalyst are equipped with a system for supplying heat with recycle flows to compensate for the amount of heat missing for the aromatization reactions of the raw material.

The installation also differs in that the temperature of the raw mixture at the entrance to each section (layer, shelf) with a zeolite-containing aromatization catalyst is at least 15 ° C higher than at the entrance to the previous section (layer, shelf) with a zeolite-containing oligomerization catalyst.

It is known that in order to increase the degree of oligomerization of lower olefins into a high-octane component of motor gasoline and reduce the formation of aromatic hydrocarbons, and, consequently, increase the regeneration run of the catalyst, the process of converting lower olefins

- 11 039642

C2-C10 must be carried out under mild conditions (the preferred temperature is from 280°C to 400-450°C, gradually increasing as the catalyst coking) using a zeolite-containing catalyst of the pentasil group with the lowest acidity, and therefore, without the content of rare earth elements (lanthanum , cerium, neodymium, praseodymium) as promoters that increase the aromatizing ability of zeolite-containing catalysts. An example of such catalysts are zeolite-containing catalysts, which, in addition to zinc oxide, contain iron (III) and calcium oxides (RF patent No. 2133640, 1998, Bull. No. 21, 1999) or iron oxides (III), calcium and magnesium (RF patent No. 2169043, 2000, Bull. from. No. 17, 2001 and Eurasian patent No. 002139, 2000). It was this catalyst that we used as an oligomerization catalyst. The specific composition of such a catalyst is given above (US Pat. RF No. 2133640, example 1).

The introduction of rare earth elements into zinc-containing zeolite-containing catalysts facilitates the aromatization reactions of lower paraffins and olefins. It is known (US patent No. 4128504, 1978, RF patent 2092240, 1997, Bull. from. No. 28, RF patent No. 2087191, 1996, Bull. from. No. 23, 1997 (together with fluorine) , RF patent No. 2098455, 1996, Bull. No. 34, 1997 (together with boron oxide), RF patent No. 2100075, 1996, Bull. No. 36, 1997) (together with phosphorus), RF patent No. 2172212, 1999, Bull. from. No. 23, 2001 (together with iron(III) and magnesium oxides), RF patent No. 2165293, 2000, Bull. from. No. 11, 2001 (together with oxides of calcium, iron (III) and magnesium) introduction of rare earth elements (lanthanum, cerium, neodymium, praseodymium) into the composition of zinc-containing zeolite-containing catalysts. It was this catalyst that we used as an aromatization catalyst, namely, _the catalyst according to example 1 of the patent of the Russian Federation No. SiO ₂ /Al ₂ O ₃ =35) - 65; ZnO - 2; sum (REE) ₂ O ₃ - 2 (REE - cerium, lanthanum, neodymium, praseodymium); oxides of iron (III), magnesium and calcium - 0.5 each; γ-alumina (binder) - the rest up to 100 wt.%

The proposed complex installation can be illustrated by the following examples (examples 30-31). Test time on the mode - 8 hours.

Example 30.

According to the described technological scheme (Fig. 1) under the following conditions: temperature at the inlet to the first shelf (catalyst zone) 300°C, pressure 1.8 MPa, feed space velocity (liquid) 2 h ^-1 out of 14600.0 g/ hours of butane-butylene fraction (composition, wt.%: propane - 1.0; propylene - 1.0; isobutane - 8.0; n-butane - 28.0; butenes - 62.0; total olefins - 63.0 ) and 4000.0 g/h of methanol, upon contact with zeolite-containing oligomerization and aromatization catalysts in four zones of the adiabatic reactor, a gasoline fraction was obtained, according to the group chemical composition corresponding to the requirements of GOST R 51866 - 2002 for motor gasoline, i.e. containing in its composition no more than 18 vol.% olefinic hydrocarbons and no more than 35 vol.% aromatic hydrocarbons. The first, second and third catalyst zones were loaded with a zeolite-containing oligomerization catalyst (according to US Pat. RF No. 2133640, example 1), and the zeolite-containing aromatization catalyst prepared as described above was loaded into the fourth catalyst zone. When the temperature at the outlet of each oligomerization zone is only 5-8°C higher than the temperature at the inlet to the same zone, while maintaining the following temperature values at the inlet to each catalyst zone: the second - 310°C, the third - 320°C, the fourth - 360°C, 11242.0 g/h of stable gasoline is obtained with a benzene content of less than 1.0 vol.% (0.85 vol.%) and a research octane number of 98.0 points (composition, vol.%: n - paraffins 11.5; _{isoparaffins 53.7} ; olefins 7.0; naphthenes 5.3; aromatic hydrocarbons 22.5); 0 g/h of dry hydrogen-containing gas (gas purge) sent to the fuel network.

Example 31.

According to the described technological scheme (Fig. 1) under the following conditions: temperature at the inlet to the first shelf (catalyst zone) 300°C, pressure 1.8 MPa, feed space velocity (liquid) 2 h ^-1 of 18600.0 g/ h of butane-butylene fraction (composition, wt.%: propane - 1.0; propylene - 1.0; isobutane - 8.0; nbutane - 28.0; butenes - 62.0; total olefins - 63.0) at contact with zeolite-containing oligomerization and aromatization catalysts in four zones of the adiabatic reactor, a gasoline fraction was obtained, according to the group chemical composition corresponding to the requirements of GOST R 51866 - 2002 for motor gasolines, i.e. containing in its composition no more than 18 vol.% olefinic hydrocarbons and no more than 35 vol.% aromatic hydrocarbons. A zeolite-containing oligomerization catalyst (according to US Pat. RF No. 2133640, example 1) was loaded into the first and third catalyst zones, and a zeolite-containing aromatization catalyst prepared as described above was loaded into the second and fourth catalyst zones. When the temperature at the outlet of each oligomerization zone is only 5-8°C higher than the temperature at the inlet to the same zone, while maintaining the following temperature values at the inlet to each catalyst zone: the second - 340°C, the third - 350°C, the fourth - 365°C, 10230.0 g/h of stable gasoline is obtained with a benzene content of less than 1.0 vol.% (0.95 vol.%) and an octane number of 98.5 research points (composition, vol.%: n - paraffins 8.7; _{isoparaffins 55.5} ; olefins 6.3; naphthenes 5.0; aromatic hydrocarbons 25.5); purge) sent to the fuel network.

- 12 039642

Table 1

Example No.: Indicators: one 2 prot. 3 4 5 6 7 8 prot. nine 10 prot. eleven* 12* prot. 1. Composition of raw materials (% vol.): BBF 45.5 91.0 60.6 15.2 48.85 38.6 45.5 91.0 45.5 91.0 45.5 91.0 ethylene + H ₂ + C _{1 +} C ₂ - - - - - - - - - - - - PPF olefins С ₅ -С ₆ methanol DME С ₂ Н ₅ OH alcohols С ₃ -С ₅ isobutane total: 2. Process conditions: temperature, °C pressure, MPa volumetric feed rate (by liquid), hour ¹ : olefin-containing fraction, ml/h olefins, g/h oxygenate, ml/h g/h olefin-containing fraction + isobutane ml/h g/h isobutane: olefin-containing fraction (by volume) 4. Results obtained: 9.0 45.5 100 310 1.8 3.3 150 55.6 30 23.7 300 168 1:1 9.0 100 310 1.8 3.3 300 111.2 30 23.7 9.0 30.4 100 310 1.8 3.3 200 74.1 30 23.7 300 168 1:0.5 9.0 75.8 100 310 1.8 3.3 fifty 18.5 30 23.7 300 168 1:5 2.3 48.85 100 310 1.8 3.3 161.2 59.8 7.6 6 322.4 180.5 1:1 22.8 38.6 100 310 1.8 3.3 127 47 76 60 254 142.2 1:1 9.0 45.5 100 280 0.1 0.5 150 55.6 30 23.7 300 168 1:1 9.0 100 280 0.1 0.5 300 111.2 30 23.7 9.0 45.5 100 450 4.0 10.0 150 55.6 30 23.7 300 168 1:1 9.0 100 450 4.0 10.0 300 111.2 30 23.7 9.0 45.5 100 310 1.8 3.3 150 55.6 30 23.7 300 168 1:1 9.0 100 310 1.8 3.3 300 111.2 30 23.7 olefin conversion, % wt. 99.3 98.3 99.1 98.7 99.3 98.6 97.4 96.5 84.2 88.5 99.5 98.9 oxygenate conversion, % wt. 100 100 100 100 100 100 88.7 90.2 100 100 100 100 isobutane conversion, % wt. 1.2 3.4 1.6 0.3 1.1 1.4 1.6 4.8 20.1 22.4 1.3 3.7 n-butane conversion, % wt. 4.6 26.6 5.1 2.8 4.8 4.7 3.9 22.4 23.7 25.5 4.9 27.0 amount of C ₅₊ hydrocarbons after debutanization, g/h 61.3 118.7 82.4 20.0 65.6 53.0 60.3 115.1 57.5 112.5 62.1 119.1 the output of debutanized gasoline on olefins of raw materials, % wt. 110.3 106.7 111.2 108.3 109.7 112.8 108.5 103.5 103.4 101.2 111.7 107.1 Composition of hydrocarbons С ₅₊ (gasoline), % vol.: n-paraffins 13.7 28.5 15.6 16.7 14.2 17.9 17.6 29.5 10.5 25.6 13.2 27.8 isoparaffins 56.7 36.7 54.2 58.1 57.8 54.6 55.4 41.0 49.6 35.0 57.9 36.3 olefins 8.3 19.6 9.1 5.9 10.3 4.8 12.5 17.9 1.4 2.3 8.3 22.5 naphthenes 5.6 6.3 5.9 6.3 9.9 5.2 5.1 6.2 3.8 4.7 5.3 5.0 aromatic hydrocarbons 15.7 8.9 15.2 13.0 7.8 17.5 9.4 5.4 34.7 32.4 15.3 8.4 Octane number of debutanized gasoline, points: motor method 84.0 82.0 83.1 83.5 83.0 84.5 82.5 81.0 83.0 82.1 83.5 81.5 research method 97.5 96.0 96.5 96.2 95.5 97.4 95.1 94.3 96.5 94.5 97.0 96.5 * - adiabatic reactor

- 13 039642

table 2

Example No.: Indicators: thirteen 14 prot. fifteen 16 prot. 17 18 prot. nineteen 20 prot. 21 22 prot. 23 24 prot. 1. Composition of raw materials (% vol.): BBF - - - - - - 45.5 91.0 45.5 91.0 45.5 91.0 ethylene + H ₂ + C _{1 +} C ₂ 45.5 91.0 - - - - - - - - - - PPF - - 45.5 91.0 - - - - - - - - olefins C ₅ -C ₆ - - - - 45.5 91.0 - - - - - - methanol 9.0 9.0 9.0 9.0 9.0 9.0 - - - - - - DME - - - - - - 9.0 9.0 - - - - C ₂ H ₅ OH - - - - - - - - 9.0 9.0 - - alcohols С ₃ -С ₅ - - - - - - - - - - 9.0 9.0 isobutane 45.5 - 45.5 - 45.5 - 45.5 - 45.5 - 45.5 - total: 100 100 100 100 100 100 100 100 100 100 100 100 2. Process conditions: temperature,°C 350 350 310 310 310 310 310 310 310 310 310 310 pressure, MPa 1.8 1.8 1.8 1.8 1.8 1.8 1.8 1.8 1.8 1.8 1.8 1.8 volumetric feed rate of raw materials (by liquid), hour ¹ 2.0 2.0 3.3 3.3 3.3 3.3 3.3 3.3 3.3 3.3 3.3 3.3 3. Filed raw materials: olefin-containing fraction, ml/h 35l/h 70l/h 150 300 150 300 150 300 150 300 150 300 olefins, g/h 36.7 73.4 55.7 111.4 98.6 197.2 55.6 111.2 55.6 111.2 55.6 111.2 oxygenate, ml/hour thirty thirty thirty thirty thirty thirty thirty thirty thirty thirty thirty thirty g/hour 23.7 23.7 23.7 23.7 23.7 23.7 20.0 20.0 24.4 24.4 24.1 24.1 olefin-containing fraction + isobutane ml/hour - 300 - 300 - 300 - 300 - 300 - g/hour 121.0 - 163.5 - 182.6 - 168 - 168 - 168 - isobutane: olefin-containing fraction (by volume) 1:1 1:1 1:1 1:1 1:1 4. Results obtained: olefin conversion, % wt. 76.3 73.9 87.6 85.9 97.6 96.3 99.2 98.4 99.3 98.2 99.4 98.5 oxygenate conversion, % wt. 100 100 100 100 100 100 100 100 100 100 100 100 isobutane conversion, % wt. 1.1 - 1.4 - 1.1 - 1.2 3.2 1.2 3.3 1.1 3.1 n-butane conversion, % wt. - - - - - - 4.5 26.3 4.6 26.1 4.4 26.7 amount of C ₅₊ hydrocarbons after debutanization, g/h 33.6 58.9 60.7 90.8 104.3 189.8 61.8 119.1 62.0 119.6 66.4 124.3 the output of debutanized gasoline on olefins of raw materials, % wt. 91.6 80.1 108.9 81.5 105.8 96.2 111.2 107.1 111.5 107.6 119.4 111.8 Composition of hydrocarbons С ₅₊ (gasoline), % vol.: n-paraffins 14.0 28.2 13.2 16.8 15.4 17.9 16.9 27.4 14.5 26.5 14.1 27.5 isoparaffins 55.4 36.9 54.6 37.2 51.8 31.1 54.1 35.1 55.0 36.1 56.0 35.8 olefins 8.1 18.5 7.9 20.2 7.5 19.1 6.1 17.8 8.6 22.9 8.2 19.1 naphthenes 5.0 4.9 5.8 5.7 5.2 5.4 5.8 5.4 5.6 5.4 5.1 5.2 aromatic hydrocarbons 17.5 11.5 18.5 20.1 20.1 26.5 17.1 14.3 16.3 9.1 16.6 12.4 Octane number of debutanized gasoline, points: motor method 84.2 81.3 84.6 83.0 83.0 83.5 84.5 82.1 84.8 82.3 83.6 81.7 research method 97.9 95.5 98.2 95.9 95.1 95.0 97.7 95.8 97.7 96.0 97.3 95.4

- 14 039642

Table 3. Catalyst operation parameters and methods for their calculation

№№ p/n Indicator Calculation one 2 3 one Conversion of olefins, % wt. (wt olefins reacted/wt olefins supplied) x 100 The mass of unreacted olefins was determined by the sum of their content in the off-gases and debutanization gases 2 Isobutane conversion, % wt. (Mass of ISO-С4Н10 supplied - mass of ISO-С4Н10 in reaction products *) / mass of ISO-С4Н10 supplied) x 100 (*taking into account its content in debutanization gases) 3 Conversion of n-butane, % May. (Mass of supplied Н-С4Н10 - mass of НС4Н10 in reaction products*)/mass of supplied Н-С4Н10) x 100 (*taking into account its content in debutanization gases) 4 Yield of gasoline for raw olefins, % ma. (The mass of gasoline after debutanization/mass submitted olefins) x 100

CLAIM

Claims (12)

CLAIM 1. The method of catalytic oligomerization of individual C ₂ -C _{1o olefins} or mixtures thereof with each other in any ratio and/or with C1-C10 paraffins and/or with hydrogen in the presence of C1-C6 oxygen-containing compounds selected from C1-C6 alcohols or their ethers, into gasoline hydrocarbons at elevated temperature and pressure, selected from the temperature range of 250-450°C and pressure of 0.1-4.0 MPa, in which the raw material is contacted with a zeolite catalyst for oligomerization of the composition, wt.%: zeolite - 40 .0-80.0, promoter(s) - 1.0-4.0, binder component - the rest up to 100, containing pentasil group zeolite selected from ZSM-5, ZSM-11, with mole ratio SiO2/Al2O ₃ = 20-80 mol/mol and a residual sodium oxide content of not more than 0.2 wt.%, containing a binder selected from γ-alumina, silicon dioxide, synthetic aluminosilicate, zirconium dioxide, containing a zinc oxide promoter (s), or a mixture zinc oxide and one or more elements or and x oxides selected from iron (III), magnesium, calcium, gallium, cerium, lanthanum, neodymium, praseodymium, platinum, cobalt, chromium, copper, boron, fluorine, phosphorus, characterized in that isobutane is additionally introduced into the feed stream in an amount (0.5-5.0):1 in relation to the olefin-containing fraction (by volume) with the ratio of the olefin-containing fraction to the oxygen-containing fraction (by volume) 1.0:(0.05-0.5). 2. The method according to claim 1, characterized in that the process of oligomerization of olefin-containing raw materials is carried out at a temperature of 250-450°C, a pressure of 0.1-4.0 MPa, a volumetric feed rate of raw materials (by liquid) of 0.1-10.0 h ^-1 under adiabatic or isothermal conditions. 3. Integrated installation for the implementation of the methods according to claims 1, 2 for the production of high-octane gasolines, diesel fractions or aromatic hydrocarbons from a mixture of C1-C ₁₀ hydrocarbons of various compositions, selected from low-octane gasoline fractions n.k. - 180°C, 90-160°C or more narrow cuts, pentane-heptane or hexane cuts, propane-butane cuts, wide cuts of light hydrocarbons - a product of gas processing plants (NGLs), and/or C ₂ -C _{1o olefins} , and /or mixtures thereof with each other, and/or with C1- _C10 paraffins, and/or with hydrogen in the presence of oxygen-containing compounds selected from C1-C6 alcohols and/or their ethers, comprising one or more parallel sectioned adiabatic reactors , consisting of one or more stationary layers (sections) of a zeolite-containing catalyst with heat supply or removal between the layers (sections) of the catalyst, or one or more parallel isothermal reactors with heat pipes, and / or coils, and / or pipe heat exchangers, and / or panels with heat supply or removal with a zeolite-containing catalyst and the possibility of supplying it to the raw mix, as well as to the second and each subsequent layer (section) with zeolite-containing - 15 039642 using a catalyst in an adiabatic reactor of the part of the gas (H2 and hydrocarbons C1-C ₄ ) heated in a fire or electric heater, separated in a three-phase separator (C-1) from the flow of reaction products after their partial condensation, in order to circulate it through the catalyst for supply or removal of heat in the adiabatic reactor (R-1 or R-2), conversion of unsaturated hydrocarbons contained in it and increase in the inter-regeneration run of the catalyst, technologically connected with the reactor (reactors) by pipelines, furnaces for heating raw materials, fire or electric heater for heating circulating reaction products , heat exchangers for heating raw materials and products, refrigerators for cooling products, separators, including three-phase (C-1), for partial condensation and separation of products, tanks, pumps, compressors, distillation columns for distillation of reaction products, characterized in that it includes unit, consisting of at least one reactor (R-3), to remove sulfur compounds using hydrogen, including its own, formed in the process of hydrogen-containing gas, from at least part of the C1-C ₁₀ hydrocarbon feedstock, into which the heated feedstock is sent, a stream of raw material purified from sulfur is removed from the reactor (R-3) and sent into the feed heater of an isothermal reactor (R-1 or R-2) or the first layer (section) of an adiabatic reactor (R-1 or R-2), includes a distillation column (K-2) for separating C ₅₊ components from hydrocarbon feedstock into in the case of aromatization of this raw material, and into the heater of the raw material of the isothermal reactor or the first layer (section) of the adiabatic reactor (R-1 or R-2), a vapor-phase stream containing hydrocarbons C ₁ -C ₄ is sent from the top of the column (K-2) after reflux separation, with In this case, fractions C ₃ , C ₄ , C ₅ -C ₆ or C5-C7 are separated from NGL by rectification, and these fractions are sent for aromatization separately from each other in different isothermal reactors or on different shelves (sections) with a zeolite-containing catalyst in adia batic reactor under different temperature conditions, includes at least one adsorber (A-1) filled with an adsorbent and selectively adsorbing C ₅₊ hydrocarbons, including aromatic hydrocarbons, into which the vapor-phase stream separated in a three-phase separator (C-1) from product stream, and a vapor phase stream depleted in C ₅₊ -hydrocarbons containing hydrogen and hydrocarbons Ci-C ₄ is removed, while isobutane is also used as a feed stream, which is preheated alone and / or in a mixture with oxygen-containing compounds C1-C6 and / or in a mixture with olefin-containing feedstock containing C ₂ -C ₁₀ olefins, both at the inlet to the reaction part of the isothermal reactor or into the first layer (section) of the adiabatic reactor, and into the second and / or each subsequent layer (section) adiabatic reactor, while the sequence of connecting technological apparatuses by pipelines in the installation is such that the raw material is the raw hydrocarbon fraction C1-C ₁₀ and isobutane - from the feed tank (E-1) goes to the intake of the feed pump (H-1), where the other pump (H-2) supplies the feed oxygen-containing fraction C1-C6 from the second feed tank (E-2), while NGL is preliminarily is separated in a distillation column for the separation of NGL (K-2) into fractions C1-C4, C5-C7, C ₅ -C ₆ , C ₅₊ , C ₆₊ or C7+, and any of these fractions can be used as a raw material for the plant, the unit includes a reactor (reactor) for desulphurization (R-3) of the raw hydrocarbon fraction C1-C ₁₀ , located upstream of the NGL separation distillation column (K-2), using the hydrogen-containing gas obtained in the process, then the raw mixture passes through the pipe space of two heat exchangers (T-1, T-3), where it is heated by the gas product stream coming from the catalytic reactor of the plant (R-1 or R-2) operating in the synthesis mode, then the feed stream is heated in the furnace (P-1) to the required start temperature reaction (250-510°C) and enters the catalytic reaction a reactor (R-1 or R-2) operating in the synthesis mode, at the same time, circulating gas is supplied to the first or each subsequent catalyst layer between these layers in the adiabatic reactor, as well as to the inlet to the isothermal reactor by a circulation compressor (PC-1) , heated in the recuperation heat exchanger (T-2) and furnace (P-3) to 410-480°C, as well as hot recycle flow from the distillation column, while in the reactor (R-1 or R-2) at a temperature of 250- 510°C and a pressure of 0.1-10.0 MPa, the process of transformation of raw materials is carried out, a product stream is removed from the reactor, which sequentially passes through the annular space of three successively located heat exchangers (T-1), (T-2) and (T-3) , where it gives off heat to the feed stream and circulating gas, then another heat exchanger (T-4), where it gives off heat to the unstable product flow, then it is cooled in an air cooler (XB-1) to a temperature of 55 ° C and then cooled with circulating water in a water cooler (X-1) up to 35°C, then the gas-liquid mixture of products enters the three-phase separator (C-1), where it is separated into a gas stream, water condensate and an unstable liquid product, then the main part of the gas stream separated in the three-phase separator enters the circulation compressor unit (PC-1) and at a pressure of 1, 9-2.0 MPa in the production of high-octane gasolines and 4.0-7.0 MPa in the production of diesel fractions from olefin-containing raw materials enters the heat exchanger coil (T-2), then to the circulating gas heating furnace (P-3) and then to a reactor (R-1 or R-2) operating in the synthesis mode, water condensate containing acidic components, and unreacted alcohol or ether are sent to a water treatment system or an oil desalination unit, and an unstable product, which is a gasoline or diesel fraction, or concentrate of aromatic hydrocarbons, the pump (H-3) is fed into the heat exchanger coil (T-4), then into the heat exchanger coil (T-5), where it is heated by the heat of a stable product and further along steps to stabilize in the main river - 16 039642 fication column of the process (K-1), from the top of which vapors are removed, which are cooled and condensed in an air cooler (XB-2) and a water cooler (X-2), and then uncondensed gas, which is sent to the fuel network, and liquefied gas, part of which is sent by a pump (H-4) to the main distillation column (K-1) as irrigation, and the balance part is removed from the installation, the bottom product of this column enters the reboiler (RB ), where light fractions are stripped from it, which then go under the lower plate of the column, and a stable product, which is a high-octane gasoline or diesel fraction, is cooled in the annulus of the heat exchanger (T-5) and then goes to the warehouse, if received in as the target stable product of the distillation column (K-1) of the diesel fraction obtained by oligomerization of olefin-containing raw materials, the olefin-containing gasoline fraction formed in the process is removed sideways distillation column (K-1) and is recycled to the synthesis reactors (R-1 or R-2), in the case of obtaining a concentrate of aromatic hydrocarbons as a stable product of the column (K-1) in the process of aromatization of hydrocarbon raw materials aromatic and non-aromatic hydrocarbons are separated by absorption and/or adsorption and/or extractive distillation (distillation) in the respective block (A-2) to which the product fraction is sent, and the content of non-aromatic hydrocarbons in the composition of aromatic hydrocarbons after the separation block does not exceed 1000 ppm, after separation unit, aromatic hydrocarbons can be separated in another distillation column (K-3) into individual hydrocarbons, such as benzene, toluene, xylenes, ethylbenzene, C ₉₊ aromatic hydrocarbons, naphthalene and its derivatives, and liquid non-aromatic hydrocarbons are removed from the unit as goods or can be recycled to the aromatization process. 4. The installation according to claim 3, characterized in that as the catalyst coking in the synthesis reactors, the temperature of the incoming feed stream rises, after reaching the maximum temperature for processing each type of feedstock, the reactor with the catalyst that has lost activity is disconnected from the feed stream and switched to the oxidative regeneration of nitrogen- air mixture through the regeneration circuit, the scheme for supplying nitrogen and air to the first separator of the regeneration unit (C-4) and the line of nitrogen-air mixture from this separator to the intake of the circulation compressor (PK-2) are started, at the beginning of regeneration, the reactor is purged from combustible compounds pure nitrogen, which after the circulating compressor, the pipe space of two heat exchangers of the regeneration unit (T-6, T-7) and through the furnace of the regeneration unit (P-2) enters the reactor with coked catalyst, after leaving the reactor, nitrogen, and subsequently nitrogen - air mixture passes through the annulus of two heat exchangers of the regenerator unit tations (T-7, T-6), where heat is given off to the fresh nitrogen-air mixture, then nitrogen or nitrogen-air mixture through another cooler of the regeneration unit (XV-3) enters the second separator of the regeneration unit (S-3), where part of these gases is blown off to the candle, and the other part through the first separator of the regeneration unit (C-4) enters the regeneration circuit again. 5. Installation according to claims 3, 4, characterized in that the process of processing raw materials containing hydrocarbons Ci-Cio, oxygen-containing compounds C1-C6 and isobutane, is carried out in the presence of catalysts based on zeolites of the pentasil group in hydrogen form, selected from ZSM-5 , ZSM-11 with a molar ratio of SiO2/Al2O ₃ =20-80 mol/mol and a Na ₂ O content of not more than 0.2 wt.% at a temperature of 250-600°C, a pressure of 0.1-10.0 MPa, volumetric feed rate of raw materials (liquid) 0.1-10 h ^-1 . 6. Installation according to claims 3-5, characterized in that the partitioned adiabatic reactor, consisting of two or more layers (sections) of the catalyst, has an increasing volume distribution of the catalyst in layers (sections) from the first to the last, namely 1:2: 3:4:6 and so on. 7. Installation according to claims 3-6, characterized in that it has a separate independent circulation circuit for regeneration of a zeolite-containing catalyst with a circulating compressor, heat exchange of reverse flows, separation equipment, a fire or electric heater for circulating a nitrogen-air mixture through the reactors with nitrogen and air replenishment with a flow rate of a mixture of 60-1800 nm ³ /m ^{3 of} the catalyst and blowing regeneration gases. 8. Installation according to claims 3-7, characterized in that to increase the selectivity of the process for liquid products selected from aromatic hydrocarbons or diesel fraction, in the feed mixture after the desulphurization reactor of the adiabatic or isothermal reactors, as well as in the second and / or each subsequent a layer (section) of a zeolite-containing catalyst in an adiabatic reactor is fed for the purpose of their circulation through the catalyst, unreacted components of the liquid feedstock heated in a fire or electric heater, which are non-aromatic hydrocarbons in the case of a hydrocarbon aromatization process, or non-target liquid products, which are an olefin-containing gasoline fraction in the case of obtaining a diesel fraction from olefin-containing raw materials, separated from the product stream by absorption and/or adsorption and/or extractive distillation (distillation) or rectification methods. 9. Installation according to claims 3-8, characterized in that to control the thermal regime of the synthesis process, a separate supply of hydrocarbon raw materials is provided, namely: isobutane, C ₁ -C ₁₀ paraffins, and / or C ₂ -C ₁₀ olefins and oxygen-containing compounds both directly to the beginning of the layers (sections) of the adiabatic and isothermal reactors, and along the height of each layer (section) of the adiabatic - 17 039642 reactor through special switchgears. 10. Installation according to claims 3, 4, 6, 7, 9 for processing an olefin-containing raw mixture containing in its composition individual olefins C ₂ -C ₁₀ , or mixtures thereof with each other in any ratio, isobutane, oxygen-containing compounds selected from alcohols C1-C6 or their ethers, and/or hydrogen, and/or C1-C10 paraffins, into high-octane gasoline containing not more than 18 vol.% olefinic hydrocarbons and not more than 35 vol.% aromatic hydrocarbons, using one or more parallel sectioned adiabatic conversion reactors, consisting of two or more stationary layers (sections) with a zeolite-containing catalyst with heat removal between the layers (sections) of the catalyst, characterized in that in the first section (layer, shelf) in the direction of the feed mixture, or in the first and second, or in the first, second and third, a zeolite-containing catalyst for the oligomerization of olefin-containing components of the raw material is loaded, and in the last (last) in the direction of movement of the raw material In a section (layer, shelf) a zeolite-containing catalyst for aromatization of aliphatic hydrocarbons is loaded, and sections (layers, shelves) with zeolite-containing oligomerization and aromatization catalysts can alternate with each other, starting with a section with an oligomerization catalyst. 11. Installation according to claim 10, characterized in that the layers (sections) with the aromatization catalyst are equipped with a system for supplying heat with recycle flows to compensate for the amount of heat missing for the aromatization reactions of the raw material. 12. Installation according to claims 10, 11, characterized in that the temperature of the raw mixture at the entrance to each section (layer, shelf) with a zeolite-containing aromatization catalyst is at least 15 ° C higher than at the entrance to the previous section (layer, shelf ) with a zeolite-containing oligomerization catalyst.