RU2757120C1 - Method and installation for producing gasoline from liquid hydrocarbon fractions, oxygenates and olefin-containing gases - Google Patents

Abstract

FIELD: oil and gas industry.

SUBSTANCE: invention relates to a method for producing gasoline, in which three streams are used as raw materials, one of which includes a hydrocarbon fraction, the second stream includes oxygenate, the third stream includes an olefin-containing fraction containing olefins C₂-C₄ in the total amount from 10 to 50 wt. %, and where three reaction zones filled with a zeolite catalyst are used, with the distribution of the hydrocarbon fraction into at least one reaction zone, with the distribution of the olefin-containing fraction into three reaction zones, with the distribution of oxygenate into three reaction zones, wherein the heating of a catalyst layer in each reaction zone is from 5 to 35°C, for each shelf of a reactor, the temperature of the catalyzate at an outlet of the catalyst layer exceeds the temperature of feeding raw materials to this shelf of the reactor. The invention also concerns an installation for producing gasoline.

EFFECT: smoothing of the temperature profile of the catalyst, the production of gasoline with an octane number above 90 units with the aromatic hydrocarbon content in the product of no more than 35 vol.%, the rejection of recycling of the gaseous product and the use of internal heat exchangers in the production of high-octane gasoline.

25 cl, 5 ex, 7 tbl, 3 dwg

Description

The invention relates to the field of oil refining and petrochemical industries. More specifically: to methods and installations for the production of gasoline by joint processing of hydrocarbon fractions, oxygenates and olefin-containing gases.

The following terms are used in the description:

Gasoline - commercial gasoline or the main component (base) for the production of gasoline. In particular, gasoline obtained by the proposed method can be used to obtain motor gasolines by compounding methods (mixing gasoline fractions obtained by various oil refining processes). For example, according to the proposed method, a basis for the production of motor gasoline of ecological class K5 of the AI-92 brand in accordance with GOST 32513-2013 can be produced. In some cases, the gasoline obtained by the proposed method may not meet all the requirements for commercial gasoline by a particular region or organization. For example, the benzene content of the produced gasoline can exceed 1.0 vol. %. As another example, the aromatic content of the gasoline produced can exceed 35 vol. %. As gasoline can be considered a liquid hydrocarbon product produced by the proposed method.

Hydrocarbon fraction - a fraction of the gasoline boiling range (the boiling point is not standardized, the boiling point is not more than 215 ° C). In particular, the end of the boiling point can be 200 ° C, 180 ° C, 160 ° C or 85 ° C. Preferably, the end boiling point is not higher than 180 ° C. The onset of boiling may be, for example, 62 ° C, 85 ° C, 140 ° C. Preferably, the initial boiling point is at least 62 ° C.

Oxygenate is an aliphatic alcohol or ether. It can be selected from the group including: methanol, raw methanol, technical methanol, ethanol, dimethyl ether, other aliphatic alcohols, ethers, as well as mixtures thereof, incl. with water. May contain impurities such as aldehydes, carboxylic acids, esters, aromatic alcohols. The method does not involve the use of unsaturated (unsaturated) alcohols, for example, allyl alcohol, but their presence is possible as impurities.

Olefin-containing gas - a fraction comprising 10-50 wt. % olefins C ₂ -C ₄ (ethylene, propylene, normal butylenes, isobutylene).

The olefin-containing gas may contain inert or weakly reactive components other than olefins, for example: methane, ethane, propane, butane, hydrogen, nitrogen. For example, the olefin-containing gas can contain from 0.5 to 8.0 wt. % hydrogen, preferably from 2 to 8 wt. % hydrogen. Preferably, the mass fraction of C ₅₊ hydrocarbons in the olefin-containing gas is not more than 5.0 wt. %. Preferably, the volume fraction of hydrogen sulfide in the olefin-containing gas is not more than 0.005%.

The reaction zone is a separate volume in the reactor containing the catalyst. Several consecutive reaction zones can be located in one reactor. For example, the reaction zones can be shelves in a shelf reactor. Also, each reaction zone can be located in a separate reactor. A separate reactor can act as the reaction zone.

C _n - hydrocarbons with the number of carbon atoms p.

C _{n +} - hydrocarbons with the number of carbon atoms equal to or more than n.

REE - rare earth elements.

DME stands for dimethyl ether.

SGKK dry gas of catalytic cracking.

FAU fraction of aromatic hydrocarbons.

LPG - liquefied petroleum gases.

Mass feed rate of raw materials, h ^-1 is the amount of raw materials passed per unit time through a catalyst mass unit. For example, the mass feed rate of the i-th component:

массовый поток i-ro компонента на входе, г/ч.where

mass flow of the i-ro component at the inlet, g / h.

m _{c at} - catalyst mass, g.

Conversion is the ratio of the amount of raw material that has reacted to the amount of raw material fed into the reaction. For example, the conversion of the i-th component:

- массовый поток i-го компонента на входе, г/ч.where

- mass flow of the i-th component at the inlet, g / h.

- массовый поток i-го компонента на выходе, г/ч.

- mass flow of the i-th component at the outlet, g / h.

Selectivity is the ratio of the amount of the target component to the total amount of hydrocarbons obtained in the given process. For example, the selectivity of the i-th component of the product:

- массовый поток i-ro компонента на выходе, г/ч,where

- mass flow of the i-ro component at the outlet, g / h,

- суммарный массовый поток всех произведенных углеводородов, г/ч.

- total mass flow of all produced hydrocarbons, g / h.

The percent substitution of the oxygenate for the olefin-containing gas (% substitution) is calculated as follows:

where n _{i oxygenate} ^_ molar flow of the i-th supplied oxygenate, mol / h,

k - coefficient advising the i-th oxygenate supplied.

The k _i factor is 0.5 for methanol, 1 for other oxygenates (eg ethanol, propanol, DME).

- мольный поток j-го подаваемого олефина, моль/ч.

- molar flow of the j-th fed olefin, mol / h.

For example, if it is necessary to replace methanol with dry catalytic cracking gas containing ethylene, the replacement percentage is calculated as follows:

where n _{methanol, mol / h} - molar flow of supplied methanol, mol / h,

n _{ethylene, mol / h} ^_ molar flow of supplied ethylene, mol / h,

0.5 is the coefficient corresponding to methanol.

Similarly, if ethanol is used as the oxygenate and the olefin-containing gas contains ethylene, propylene and butylenes, the percentage of substitution is calculated as follows.

where n _ethanol is the molar flow of supplied methanol, mol / h,

n _ethylene - molar flow of supplied ethylene, mol / h,

n _propylene - molar flow of supplied propylene, mol / h,

n _butylenes - total molar flow of supplied butylenes (including isobutylene), mol / h,

1 factor for ethanol.

The percent substitution of an oxygenate for an olefin-containing gas does not take into account the water supplied to the reaction zones.

LEVEL OF TECHNOLOGY

The invention according to RU 2671568 dated 09/27/2016 includes a method for producing high-octane gasoline from low-octane gasoline and / or olefin-containing fractions and methanol. Raw materials - raw hydrocarbon fraction C ₁ -C ₁₀ I and isobutane from the raw tank E-1 are fed to the intake of the feed pump N-2, where the raw oxygen-containing fraction C ₁ -C ₆ II is fed by the pump N-1 from the feed tank E-2.

Straight-run or gas-condensate gasoline fraction or olefin-containing fraction or pentane - hexane (heptane) fraction or propane - butane fraction or NGLs can be used as hydrocarbon fraction C ₁ -C _10. NGL can be preliminarily separated in a K-2 distillation column (not shown in the diagrams) into fractions C ₁ -C ₄ , C ₅ -C ₇ , C ₅ -C ₆ , C ₅₊ , C ₆₊ or C ₇₊ , and any of these fractions can be used as a raw material for the plant. If necessary, the scheme may include a reactor for desulfurization of the raw gasoline fraction, pentane - hexane (heptane), 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 heat exchangers T-3 and T-1, where it is heated by the gas-product flow coming from the reactor R-1 or R-2, operating in the synthesis mode. Further, the feed stream III is heated in the P-1 furnace to the required reaction start temperature (250-510 ° C, depending on the raw material used) and enters the P-1 or P-2 reactor operating in the synthesis mode.

Simultaneously, gas heated in the recuperative heat exchanger T-2 and furnace P-3 to 410-480 ° C is fed to the first or each subsequent catalyst layer or between these layers in the adiabatic reactor, as well as to the inlet to the isothermal reactor by the PK-1 circulation compressor, 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, the catalytic process of converting the raw material is carried out. Product stream IV is removed from the reactor, which sequentially passes through the shell space of heat exchangers T-1, T-2 and T-3, where it gives off heat to the feed stream and circulating gas, then heat exchanger T-4, where it gives off heat to the unstable product stream, then is cooled to air cooler XB-1 to a temperature of 55 ° C and then with circulating 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, a water 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 obtaining high-octane gasolines) and 4.0-6.0 (in the preferred case when 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.

Unstable product VII (gasoline, aromatic hydrocarbon concentrate or diesel fraction) is fed by pump N-3 to the coil of the heat exchanger T-4, then to the coil of the heat exchanger T-5, where it is heated by the heat of the stable product, and then goes for stabilization to the distillation column K- 1. Vapors are removed from the top of the column, which are cooled and condensed in the air cooler XB-2 and the water cooler X-2, and then in the separator C-2, uncondensed gas is emitted, which is sent to the fuel network, and liquefied gas, part of which is pumped N-4 sent to the column as reflux, 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 are then fed under the lower plate of the column, and the stable product IX - high-octane gasoline or olefin-containing gasoline or diesel fraction - is cooled in the shell space of the T-5 heat exchanger and fed to warehouse.

As a disadvantage of the invention, according to RU 2671568, the need to isolate, heat and circulate a part of the gaseous product (H ₂ and C ₁ -C ₄ hydrocarbons) through the catalyst is considered. The patent shows that the recycle of a part of the product is necessary for the conversion of the unsaturated hydrocarbons contained therein, and for the control of the temperature of the reaction zones. This complicates the equipment used. In particular, due to the return of gas, the dimensions of the reactor, pipes, fittings increase. This problem is caused by an increase in the amount of circulating ballast, which is not processed, but is only needed to maintain the temperature in the reactor.

The proposed invention makes it possible to refuse such a recycle. Due to this, it is possible to abandon the equipment for the return, compression and heating of gas, in particular from the equipment PK-1, T-2, P-3, and also to reduce the required dimensions of the reactor, pipes, fittings.

As a disadvantage of the invention according to RU 2671568, one can also consider the fact that the entire oxygenate is mixed with the hydrocarbon feed and goes in one stream to the top of the reactor. This approach can lead to a sharp overheating of the top of the reactor. This will cause overheating of the narrow frontal catalyst bed, shortening the unit runtime between regenerations and reducing the gasoline yield. In the described installation, under these conditions, a temperature drop will be observed on the second and subsequent reactor shelves. At the same time, reactions of the formation of high-octane components, in particular alkyl-aromatic hydrocarbons, are suppressed on the second and subsequent reactor shelves.

The proposed invention makes it possible to equalize the temperature of the catalyst layer in each reaction zone, in particular, due to the refusal to supply a mixture of all hydrocarbon feedstock and oxygenate in one stream to the reactor inlet.

As a disadvantage of the invention, according to RU 2671568, the need to add significant amounts of isobutane to the feed is also considered in order to control the temperature of the reaction zones. Isobutane is a highly demanded refinery product with a high cost. Its consumption for the processing of the hydrocarbon fraction leads to an increase in the cost of gasoline production. The theoretically possible recycle is associated with a complex process of separating isobutane from the produced wide fraction of light hydrocarbons.

The proposed invention makes it possible to refuse the addition of isobutane to the feed in order to control the temperature of the reaction zones, in particular due to the possibility of distributed supply of both oxygenate and olefin-containing gas to each of the reaction zones in conjunction with a system of recuperative heat exchangers.

Application for invention WO 2017155431 dated 03/07/2017 describes a method for producing gasolines from hydrocarbon fractions, fractions of gaseous olefins and oxygenates. This method is the closest to the present invention and is taken as a prototype.

As a reactor, a reactor is used containing at least two reaction zones with a zeolite-containing catalyst, between which there is additionally located means for mixing the reaction products of the previous reaction zone and supplied methanol and / or other oxygenates and olefin-containing raw materials, and by means of a unit for feeding streams:

• to the first reaction zone of the reactor: a stream of methanol and / or other oxygenates and olefin-containing raw materials and a stream of raw hydrocarbon fractions,

• to the second reaction zone of the reactor - a stream of methanol and / or other oxygenates and olefin-containing feedstock.

The described method allows you to reduce the amount of benzene and refractory components (durene, methylnaphthalenes) in the resulting product. Temperature control over the entire catalyst bed without the use of complex heat exchange equipment is achieved by controlling:

• supply of oxygenates and olefin-containing raw materials to the first and subsequent reaction zones;

• temperature of the mixture supplied to the first reaction zone.

Unfortunately, the document does not take into account the problems that arise when replacing olefins with cheap sources of C ₂ -C ₄ olefins. In particular, the regularities of the influence of the degree of dilution of olefins on the temperature profile and quality of the product are not described. Also, the possible chemical activity of diluents (methane, ethane, impurities of heavier hydrocarbons, hydrogen, nitrogen) in the conversion process is not taken into account.

As a disadvantage of the method, it can also be considered that when obtaining gasoline under optimal conditions:

• heating of the catalyst bed does not exceed 420 ° С,

• however, the temperature of the end bed of the catalyst is 40-70 ° C lower than the specified maximum temperature.

This decrease in temperature reduces the activity of the end catalyst bed, effectively leaving it to operate under sub-optimal conditions. Uneven heating of the catalyst bed can lead to uneven coking of the catalyst bed. It is also possible to activate side reactions, for example, oligomerization of olefins instead of involving them in the processes of aromatic alkylation.

As a disadvantage of the method, it can also be considered that at low consumption of methanol, less than 20% of the mass of the converted raw material, the authors urge

• additionally overheat the feed stream supplied to the last and / or penultimate reaction zones (up to a maximum of 500 ° C);

• or use isothermal reaction zones as one or two last reaction zones of the reactor.

This patent application did not elaborate on the layout of the installation.

DISCLOSURE OF THE INVENTION

The proposed method and installation make it possible to smooth the temperature profile of the frontal catalyst layer in each of the reaction zones. In known methods, including the processing of oxygenates into gasolines, when the oxygenate flow is fed to the reaction zone, a sharp jump in temperature occurs at about 60-120 ° C. This heating is localized in the first quarter of the catalyst layer. This uneven heating leads to accelerated coking of the catalyst in the frontal bed, and also increases cracking, reducing the yield of liquid product. At the same time, after a sharp rise in temperature in the frontal catalyst layer, the temperature falls. As a result, from half to 0.75 vol. % of the catalyst used operate at temperatures less than 360 ° C. This cooling of the end catalyst layer leads to a drop in the RON of the gasoline produced. The proposed method and installation make it possible to simultaneously solve the problem of overheating of the frontal catalyst layer and cooling of the end catalyst layer.

By smoothing the temperature profile in each reaction zone, the catalyst is heated to no more than 60 ° C, preferably no more than 35 ° C. At the same time, during tests for more than 100 h, the heating peak was not localized in the first quarter of the catalyst layer, but was distributed over the first 0.75 vol. % catalyst. In this case, no noticeable shifts are observed in the position of the maximum of the heating peak.

It was found that due to the use of the proposed method and installation, the temperature of the catalyzate at the outlet of the catalyst layer is higher than the temperature of the feedstock supply to this catalyst zone. This result allows the use of only one furnace in the proposed installation scheme. The increased temperature of the catalyzate from the previous reaction zone eliminates the need for intensive heating of the catalyzate between the reaction zones using furnaces and / or electric heating. Due to the reduced load, recuperative heat exchangers can cope with the task of heating new feed streams supplied to the second and subsequent reaction zones, which are significantly cheaper in maintenance and repair compared to furnaces and / or electric heaters.

This result is achieved without the use of reactors with built-in heat exchangers, in particular, isothermal reactors, heat pipe reactors. In this case, the proposed method and installation allow the use of reactors without built-in heat exchangers, in particular adiabatic reactors. Reactors without built-in heat exchangers are cheaper to repair and maintain. They are scalable, allowing the size of the reactors used to be changed tens of times, with a predictable change in heat exchange.

The technical results of the present invention are also considered:

• obtaining gasoline with OCHI above 90 units. when the content of aromatic hydrocarbons in the product is not more than 35 vol. %;

• an increase in the product yield up to 70 wt. % on the supplied hydrocarbon fraction;

• the ability to use as a raw material olefin-containing gases with an olefin content of less than 50 wt. % without preliminary concentration of olefins;

• the ability to use olefin-containing gases, including hydrogen, as a raw material, without additional separation;

• the ability to exclude the recycling of gaseous products;

• reduction of oxygenate consumption due to partial replacement of oxygenate with olefin-containing gases.

The technical results indicated above are achieved due to the proposed method for producing gasolines, which includes the following sequential operations:

a. a hydrocarbon fraction and oxygenate are fed to the inlet of the heat exchanger T-1, from the outlet of T-1 a mixture of oxygenate and a hydrocarbon fraction is mixed with an olefin-containing gas, and enters the P-1 furnace for heating, from P-1 the reaction mixture is fed to the first shelf of the reactor P-1/1 or P-1/2, including a catalyst,

b. oxygenate and olefin-containing gas are fed to the inlet of the heat exchanger T-2, from the outlet of the heat exchanger T-2 a mixture of olefin-containing gas and oxygenate is fed to the second shelf of the reactor R-1/1 or R-1/2, including the catalyst,

c. oxygenate and olefin-containing gas are fed to the inlet of the T-3 heat exchanger, from the outlet of the T-3 heat exchanger a mixture of olefin-containing gas and oxygenate is fed to the third shelf of the R-1/1 or R-1/2 reactor, including the catalyst,

d. Catalyst from the outlet of the reactor R-1/1 or R-1/2 is distributed into three streams supplied to the heat exchangers T-1, T-2 and T-3, catalyzate from the outlet of the heat exchangers T-1, T-2, T-3 enters the inlet of the air condenser VX-1 and then through the water condenser T-4 and the chiller T-5 enters the three-phase separator BR-1, where the waste gases of the catalysis product are separated from the aqueous and hydrocarbon phases,

e. the aqueous phase is discharged into the atmospheric tank E-5, where sediment and degassing of the aqueous phase occurs, the degassing gases are discharged for afterburning into the P-1 furnace, the hydrocarbon condensate of the catalyzate is fed through the T-6 recuperator to the K-1 debutanizer column,

f. heat is supplied to the bottom of the column K-1 through the RB-1 reboiler, stabilized gasoline is removed from the reboiler of the column RB-1, and then, passing through the recuperator T-6, it cools down in the VX-3 air cooler and is taken to the finished product warehouse,

g. off-gases from column K-1 are fed to the air condenser VX-2 and then enter the reflux tank BR-2,

h. water from the BR-2 tank is discharged into the E-5 atmospheric tank, the reflux from the BR-2 tank is fed to the irrigation of the K-1 column, the excess reflux is discharged, the exhaust gas from the BR-2 tank is fed to the T-8 condenser, where propane condensation ,

i. propane is separated from non-condensed gases in the BR-3 tank and is discharged for mixing with excess reflux from the BR-2 tank, and then the mixture of reflux and condensed propane is discharged,

j. water from the tank BR-3 is discharged into the atmospheric tank E-5, the exhaust gases from the tanks BR-2, BR-3 are discharged through the steam heater T-9.

An embodiment of the invention is possible, in which a hydrocarbon fraction is supplied to the input of the heat exchanger T-1 from the supply tank E-1 by the pump N-1/1 or N-1/2, additionally from the supply tank E-3 by the pump N-3/1 or N- 3/2 oxygenate is fed to the input of the heat exchanger T-1, from the output of T-1 the mixture of oxygenate and hydrocarbon fraction is mixed with the olefin-containing gas supplied by the KP-1 compressor and fed to the P-1 furnace for heating, from the P-1 reactionary the mixture is fed to the first shelf of the reactor R-1/1 or R-1/2, the oxygenate from the tank E-3 and the olefin-containing gas from the compressor are fed to the inlet of the heat exchanger T-2 by the pump N-3/1 or N-3/2 KP-1, from the outlet of the heat exchanger T-2 a mixture of olefin-containing gas and oxygenate is fed to the second shelf of the reactor R-1/1 or R-1/2, to the inlet of the heat exchanger T-3 by pump N-3/1 or N-3 / 2 oxygenate is fed from tank E-3 and olefin-containing gas from compressor KP-1, from the outlet of heat exchanger T-3 a mixture of olefin-containing gas and oxygenate is fed to the third shelf of the reactor R-1/1 or R-1/2.

It is possible to implement the invention, in which the aqueous phase is discharged into the atmospheric container E-5, which, if necessary, operates with heating.

Perhaps the implementation of the invention, in which the reflux from the tank BR-2 pump N-4/1 or N-4/2 is fed to the irrigation of the column K-1, excess reflux from the pump N-4/1 or N-4/2 is taken to the warehouse LPG, off-gas from the BR-2 tank is fed to the T-8 condenser, where propane condensation, propane is separated from the non-condensed gases in the BR-3 tank and pump N-5/1 or N-5/2 is discharged for mixing with excess reflux from the tank BR-2 and then the mixture of reflux and condensed propane is fed to the LPG warehouse, and the exhaust gases from the tanks BR-2, BR-3 are discharged into the fuel gas network through the steam heater T-9.

It is possible to implement the invention, in which the catalyst regeneration is carried out in a stream of nitrogen-air mixture along the circuit, and the nitrogen o-air mixture is supplied from the outlet of the reactor P 1 / 1.2 to the recuperator T-10, then from the outlet of the T-10 it is fed to the inlet of the air the regeneration gas cooler VX-4, then from the VX-4 outlet it is fed to the S-1 regeneration gas separator, then from the C-1 outlet it is fed to the KP-2 regeneration gas compressor, then from the KP-2 outlet it is fed to the T-10 recuperator, then from the T-10 outlet it is fed to the P-2 regeneration gas heating furnace, then from the P-2 outlet it is fed to the P1 / 1,2 reactor inlet.

An embodiment of the invention is possible, in which the frequency of circulation for feeding with nitrogen and air through the lines of supply of nitrogen and air is from 10 to 50, and the regeneration is carried out at a pressure of 6-10 bar.

An embodiment of the invention is possible, in which the zeolite catalyst comprises:

a. zeolite of the ZSM-5 type with a modulus of SiO ₂ / Al ₂ Cl ₃ from 43 to 95, in an amount from 65 to 80 wt. %,

b. sodium oxide in an amount from 0.04 to 0.15 wt. %,

c. zinc oxide in the amount of 1.0-5.5 wt. %,

d. oxides of rare earth elements in a total amount of 0.5 to 5.0 wt. %,

e. a binder comprising silicon dioxide, aluminum oxide, or mixtures thereof. An embodiment of the invention is possible, in which the heating of the catalyst bed in each

the reaction zone is from 5 to 35 ° C.

An embodiment of the invention is possible, in which for each reactor shelf the temperature of the catalyzate at the outlet of the catalyst bed exceeds the temperature of the feedstock supply to this reactor shelf.

An embodiment of the invention is possible in which the process pressure is from 1.5 to 4.0 MPa, preferably from 2.2 to 2.7 MPa.

Possible execution of the invention, in which the temperature of the flow at the entrance to the first / second / third reaction zone is 340-370 ° C / 341-370 ° C / 350-376 ° C.

Possible implementation of the invention, in which the distribution of oxygenate between the three reaction zones is 44-62 wt. % / 18-35 wt. % / 14-21 wt. % of the total amount of oxygenate.

Possible implementation of the invention, in which the distribution of the olefin-containing gas between the three reaction zones is 15-30 wt. % / 20-40 wt. % / 30-60 wt. % of the total amount of olefin-containing gas.

Possible execution of the invention, in which the hydrocarbon fraction contains normal paraffins in an amount of 15-24 wt. %, isoparaffins in the amount of 28-56 wt. %, naphthenes in the amount of 22-40 wt. %, the rest is aromatic hydrocarbons and olefins.

It is possible to implement the invention, in which the hydrocarbon fraction can be selected from the group including straight-run gasoline, stable gas gasoline, light gas condensate, gasoline fraction with boiling points of about 62 ° - 85 ° C, raffinate, and mixtures thereof.

It is possible to carry out the invention in which the oxygenate is selected from the group consisting of aliphatic alcohols, for example, methanol, ethanol, crude methanol, technical methanol, ethanol; ethers, for example, dimethyl ether, as well as mixtures thereof, including with water.

Possible execution of the invention, in which the olefin-containing fraction includes 2.3-8.0 wt. % hydrogen.

An embodiment of the invention is possible, in which the olefin-containing fraction is selected from the group including dry catalytic cracking gas, catalytic cracking wet gas, other catalytic cracking gases and their fractionation products, off-gas from a coking unit, Fischer-Tropsch synthesis gases, and mixtures thereof ...

An embodiment of the invention is possible, in which the reaction is carried out in the gas phase in a fixed catalyst bed.

An embodiment of the invention is possible in which built-in heat exchangers are not used.

An embodiment of the invention is possible in which each reactor is an adiabatic reactor.

The above technical results are also achieved due to the proposed installation for the production of gasoline, including:

a. heat exchanger T-1 for heating the hydrocarbon fraction and oxygenate, heat exchangers T-2 and T-3 for heating oxygenate and olefin-containing gas,

b. furnace P-1 for heating a mixture of hydrocarbon fraction, oxygenate and olefin-containing gas,

c. reactor R-1/1 or R-1/2,

d. air condenser VKh-1 for feeding the catalyst from the outlet of heat exchangers T-1, T-2, T-3, water condenser T-4 for additional cooling of the catalyst after the air condenser VKh-1, chiller T-5 for cooling down the catalyst after condenser T-4 ,

e. three-phase separator BR-1 for feeding the catalyzate from the chiller T-5, where the off-gases of the catalysis product are separated from the aqueous and hydrocarbon phases,

f. atmospheric tank E-5 for settling and degassing the water phase with BR-1, BR-2, BR-3,

g. line for the return of degassing gases from E-5 to afterburning to the P-1 furnace, a T-6 recuperator for stabilized gasoline from BR-1 and an air cooler VX-3 for additional cooling of gasoline,

h. debutanizer column K-1 for stabilizing the catalyst from T-6, reboiler RB-1 of column K-1 for supplying heat and removing stabilized gasoline,

i. air cooler VH-3 for cooling stabilized gasoline, air condenser VH-2 for partial condensation of the upper product from the column K-1,

j. reflux tank BR-2 for separation of non-condensed gases, overhead product and water condensate,

k. condenser T-8 for off-gas from the tank BR-2, where propane condensation takes place,

l. tank BR-3 for separating propane from non-condensed gases, steam heater T-9 for waste gases from tanks BR-2, BR-3. It is possible to implement the invention, which additionally includes a feed tank for the hydrocarbon fraction E-1, a pump N-1/1 or N-1/2 for supplying a hydrocarbon fraction, a feed tank E-3 for oxygenate, a pump N-3/1 or N-3 / 2 for oxygenate supply, KP-1 compressor for olefin-containing gas.

Possible execution of the invention, additionally including pump N-4/1 or N-4/2 for reflux from the tank BR-2 for irrigation of column K-1, LPG storage for excess reflux from pump N-4/1 or N-4/2 and for a mixture of reflux and condensed propane, pump N-5/1 or N-5/2 for removing propane for mixing with excess reflux from the container BR-2.

It is possible to implement the invention, which additionally includes equipment for carrying out regeneration: recuperator T-10, air cooler of regeneration gases VX-4, separator of regeneration gas S-1, compressor of regeneration gas KP-2, heating furnace of regeneration gas P-2.

An embodiment of the invention is possible in which built-in heat exchangers are not used.

An embodiment of the invention is possible in which each reactor is an adiabatic reactor.

BRIEF DESCRIPTION OF DRAWINGS

Figure 1a. Schematic flow diagram of the installation, schematic representation of the process (general view of the diagram).

Figure 1b. Schematic flow diagram of the installation, schematic representation of the process (the first part of the diagram).

Figure 1c. Schematic flow diagram of the installation, schematic representation of the process (the second part of the diagram).

CARRYING OUT THE INVENTION

Getting gasoline from mixtures of hydrocarbon fractions with oxygenates and olefin-containing gases is carried out as follows.

According to Figures 1a, 1b and 1c:

There is a variant of the invention intended for the joint processing of hydrocarbon fractions, oxygenates and olefin-containing gases into gasolines, where:

• stable gasoline (BGS) is used as a hydrocarbon fraction, which is a mixture of raffinate gasoline and a 62-85 ° C fraction;

• technical methanol is used as oxygenate, then methanol;

• dry desulfurized catalytic cracking gas (SGCC) is used as the olefin-containing gas.

When describing the installation, the following designations are used:

The setup in the "synthesis" mode works as follows:

The unit operates in a semi-regenerative mode. When operating one of the reactors pos. Р-1/1 or Р-1/2 in the "synthesis" mode, another reactor can be in the modes of catalyst drying / activation, coke burning, oxidative catalyst regeneration, reactor idle, idle with pressure build-up, catalyst desorption by depressurization and nitrogen purging , catalyst desorption with circulating nitrogen.

At the inlet of the recuperative heat exchanger pos. T-1 from the supply container pos. E-1 pump pos. Н-1 / 1,2 is supplied by BGS. Additionally, from the supply container pos. E-3 with the pump pos. Н-3 / 1,2 methanol is fed to the input of the recuperative heat exchanger Т-1. From the outlet of the recuperator pos. T-1 a mixture of methanol and BGS is mixed with the SGCC supplied by the compressor pos. KP-1 and enters the heating coil of the furnace pos. P-1. From the oven coil pos. P-1 the reaction mixture is fed to the first shelf of the corresponding reactor P-1/1 (P1 / 2). For the first shelf, the reaction mixture is a heated gaseous mixture of a hydrocarbon fraction, an oxygenate, and an olefin-containing gas. The mixture temperature is 350-370 ° C.

At the inlet of the recuperative heat exchanger pos. T-2 pump pos. Н-3 / 1,2 methanol is fed from the supply tank pos. E-3 and SGKK from the compressor SGKK pos. KP-1. From the outlet of the heat exchanger pos. T-2 a mixture of SGCC and methanol with a temperature of 150-250 ° C is fed to the second shelf of the corresponding reactor, pos. P-1/1 (P-1/2).

At the inlet of the recuperative heat exchanger pos. T-3 pump pos. Н-3 / 1,2 methanol is fed from the supply tank pos. E-3 and SGKK from the compressor SGKK pos. KP-1. From the outlet of the heat exchanger pos. T-3 mixture of SGKK and methanol with a temperature of 150-250 ° C is fed to the second shelf of the corresponding reactor, pos. P-1/1 (P-1/2).

Catalyst from the outlet of the reactors pos. P-1 / 1.2 is distributed into 3 streams supplied to the heat exchangers pos. T-1, T-2 and T-3. Catalyst consumption at the inlets of heat exchangers pos. T-2, T-3 is set by the temperature of the raw mixtures supplied to the second and third shelf of the reactor pos. P-1 / 1.2. The temperatures of the reaction mixtures fed to the catalyst beds on the second and third shelf of the reactors are controlled by the temperatures of the feed mixtures supplied to the second and third shelf.

Due to the conversion of additional feedstock (methanol and SGCC) supplied to the reactor shelves, the catalyst bed is adiabatically heated on shelves 1, 2 and 3 of the reactor, pos. P-1 / 1.2. The adiabatic heating of the catalyst bed on each shelf should not exceed 60 ° C. The amount of adiabatic heating is controlled by the mass ratio between the additional components of the reaction mixture (methanol + SGCA) supplied to the first, second and third shelves, and BGS. The initial temperature of the reaction mixture supplied to the frontal catalyst layer for the first shelf is determined by the temperature of the reaction mixture from the furnace coil, pos. P-1, and for the second and third shelves, the temperature of the reaction mixture at the outlet from the previous shelf, as well as the mass and temperature of the additional raw material supplied to the second and third shelves.

Catalyst from the outlet of the heat exchangers pos. T-1, T-2, T-3 enters the inlet of the air condenser pos. VX-1 and further through a water condenser along. T-4 and chiller pos. T-5 with a temperature of + 5-10 ° C enters the three-phase separator pos. BR-1. In a three-phase separator, the off-gases of the catalysis product are separated from the aqueous and hydrocarbon phases of the catalysis. The water phase is discharged into the atmospheric tank pos. E-5. In the container pos. E-5 sediment and degassing of the aqueous phase is carried out. Degassing gases are directed to a low pressure flare header. To destroy possible emulsions of the aqueous phase with catalyst particles, heating of the tank pos. E-5.

Hydrocarbon condensate catalyzed through the recuperator pos. T-6 is fed as a vapor-liquid feed to the debutanizer column pos. K-1. Heat is supplied to the bottom of the column through a medium-pressure steam-heated reboiler pos. RB-1. The temperature of the bottom of the column is 160-170 ° C, the pressure of the bottom is 0.6 MPa. The column is equipped with 25-30 once-through valve trays.

Stabilized gasoline with a saturated vapor pressure of 40-60 kPa is removed from the column reboiler, pos. RB-1, cools down to 110-120 ° C in the air cooler pos. VX-3 and further through the recuperator pos. T-6 with a temperature of 25-35 ° C is removed by the pump pos. Н-8 / 1,2 for use as a component for the preparation of commercial gasolines.

Exhaust gases from the helmet pipe of the column pos. K-1 is fed to the air condenser pos. VX-2 and then through a water cooler, chilling to 30-45 ° C, enter the reflux tank pos. BR-2 equipped with a drainage device. Water from the container pos. BR-2 is positionally diverted into the atmospheric container pos. E-5.

Reflux (hydrocarbon condensate) from the tank pos. BR-2 pump pos. Н-4 / 1,2 is fed to the column irrigation pos. K-1. Reflux ratio (as the ratio of the volumetric flow rate for refluxing the column at pos. K-1 to the flow rate of unstable gasoline (hydrocarbon condensate of catalyzate)) from the outlet of pos. BR-1 is 0.8 -0.9). Excessive reflux from the pump pos. Н-4 / 1,2 is taken to the LPG warehouse.

Exhaust gas from the tank pos. BR-2 is fed to the condenser pos. T-8. In the condenser pos. T-8 at a temperature of + 5-10 ° C, propane condensation occurs. Propane is separated from non-condensed gases in the tank pos. BR-3 and pump pos. Н5 / 1,2 is diverted for mixing with excess reflux from the tank pos. BR-2 and then a mixture of reflux and condensed propane is fed to the LPG warehouse. Capacity pos. BR-3 is equipped with a drainage device. Water from the container pos. BR-3 is positionally diverted into the atmospheric container pos. E-5. The saturated vapor pressure at a temperature of 45 ° C of the mixed LPG from the installation is controlled by the temperature at the outlet of the condenser pos. T-8. The content of liquid hydrocarbons C ₅₊ in LPG is controlled by the temperature of the top of the column pos. K-1 and reflux ratio. Mixed LPG (reflux + propane condensate) in composition and physical. chem. indicators correspond to the requirements for commercial LPG of the SPBT type in accordance with GOST 20448-90. Waste gases from containers pos. BR-2 and BR-3 through a steam heater pos. Т-9 with a temperature of 35-45 ° С are diverted to the fuel gas network.

Examples of

The results achieved are illustrated in Examples 1-5 below. For examples 1-5, the yield and other process indicators are shown for stable gasoline (hereinafter liquid hydrocarbon product), which does not contain dissolved gases C ₁ -C ₄ (fraction C _{5+ of the} hydrocarbon fraction of the product). This is due to the fact that liquid hydrocarbon products obtained and stored under different conditions may contain different amounts of dissolved gases. In this case, the content of dissolved gases in can vary unevenly over time, changing the chemical composition. This can lead to inadequate comparison of the yield and quality of products from different experiments, especially when comparing results from different enterprises. Therefore, it is more preferable to compare the parameters of a product that does not contain dissolved gases. However, we note that, depending on the goals of a particular production, a liquid hydrocarbon product may contain not only C ₅₊ hydrocarbons, but also a different amount of dissolved gases C ₁ -C ₄ .

Since C ₅₊ hydrocarbons (hydrocarbons with five or more carbon atoms) are the main component of a liquid hydrocarbon product, the yield of a liquid hydrocarbon product increases simultaneously with the yield of C ₅₊ hydrocarbons.

For experiments 2-5, a catalytic unit was used, including three reactors connected in series, with a total catalyst load of up to 9 liters. The reactors are designated as first, second and third reaction zones, R ₁₀₁ , R ₂₀₁ , R _301, respectively.

The reactors are structurally as close as possible to the adiabatic type, the heat exchange between the catalyst bed and the vessel is minimized. The catalyst baskets are placed in the reactor vessel so that there is a gap (approximately 2 mm) between the wall of the basket and the strong body. Each reactor is installed in a thermostat with three heating zones. Three thermocouples are placed between the surfaces of the heating elements and the outer surface of the reactor vessel. On the contrary, thermocouples are also located on the inner wall of the reactor basket casing. There is also an air gap not exceeding 3-4 mm between the inner surface of the thermostat and the outer surface of the reactor. The control loops maintain a constant temperature difference between the thermocouples at the outer wall of the reactor and the thermocouple opposite at the inner surface of the reactor basket.

Liquid and gaseous products for analysis began to be taken 4 hours after the start of the feed.

Example 1 was carried out as follows. The material balance of example 1 is shown in Table 1. To carry out the reaction, a P-1/1 reactor was used. If necessary, instead of the R-1/1 reactor, the R-1/2 reactor can be used.

A hydrocarbon fraction is fed to the inlet of the heat exchanger T-1, which is a mixture of a 62-85 ° C fraction and a raffinate. The mass consumption of the 62-85 ° C fraction was 2493.5 g / h, the mass consumption of the 62-85 ° C fraction was 2493.5 g / h.

Oxygenate, which is methanol, is fed to the inlet of the T-1 heat exchanger with a mass flow rate of 110.0 g / h. From the outlet of T-1, a mixture of oxygenate and a hydrocarbon fraction is mixed with olefin-containing gas of the SGCC, the mass flow rate of which was 558.7 g / h, and is fed to the P-1 furnace for heating.

From P-1, the reaction mixture is fed to the first shelf of the R-1/1 reactor, at a feed temperature of 360 ° C. For the first shelf, the reaction mixture is a heated gaseous mixture of BGS, methanol and SGCC. Heating over the catalyst bed on the first shelf was 24 ° C. The temperature at the outlet of the catalyst bed was 371 ° C, which is 11 ° C higher than the temperature of the feed to the first shelf. This effect makes it possible, in particular, to direct the product leaving the first shelf to the second shelf without additional heating of the intermediate conversion product. In particular, it is possible to abandon the use of internal heat exchangers in the used reactor. Also, for example, it is possible to avoid heating the product stream leaving the first shelf in the oven before being directed to the third shelf.

Oxygenate, which is methanol, is fed to the inlet of the T-2 heat exchanger, where the mass flow rate of methanol was 87.5 g / h. The inlet of the heat exchanger T-2 is also fed with olefin-containing gas from the SGCC, the mass flow rate of which was 558.7 g / h.

From the outlet of the heat exchanger T-2, a mixture of olefin-containing gas and oxygenate is fed to the second shelf of the R-1/1 reactor. The feed temperature to the second shelf was 345 ° C. Due to the fact that the intermediate product stream from the first shelf is not heated on the T-2 heat exchanger, the T-2 heat exchanger is able to provide a given feed temperature without using a furnace or an internal heat exchanger. Heating over the catalyst bed on the first shelf was 32 ° C. The temperature at the outlet of the catalyst bed was 370 ° C, which is 25 ° C higher than the temperature of the feed to the second shelf. A similar result is achieved using an adiabatic reactor.

Oxygenate, which is methanol, with a flow rate of 52.5 g / h and an olefin-containing gas, which is SGCC, with a flow rate of 744.9 g / h, is fed to the inlet of the heat exchanger T-3. From the outlet of the heat exchanger T-3, a mixture of olefin-containing gas and oxygenate is fed to the third shelf of the R-1/1 reactor. The feed temperature to the second shelf was 375 ° C. Heating up along the catalyst bed on the first shelf was 21 ° C. The temperature at the outlet of the catalyst bed was 389 ° C, which is 14 ° C higher than the temperature of the feed to the second shelf. Thus, in each of the three reaction zones, heating over the catalyst bed does not exceed 35 ° C. In this case, cooling of the end layer of the catalyst is not allowed, as evidenced by the fact that the temperature of the stream at the outlet from each reaction zone is higher than the temperature of the feed to this reaction zone.

Thanks to the proposed method and installation, it is possible to maintain uniform heating of the catalyst layer in each reaction zone. At the same time, it is possible to abandon the use of reactors with internal heat exchangers, in particular, from tubular reactors. It is also possible to maintain relatively high temperatures of the end catalyst bed of each reaction zone without overheating the front catalyst bed (the maximum temperature of the catalyst bed is no more than 35 ° C higher than the feed temperature in each reaction zone).

Catalyst from the outlet of the R-1/1 reactor is distributed into three streams fed to heat exchangers T-1, T-2 and T-3. Catalyst from the outlet of heat exchangers T-1, T-2, T-3 enters the inlet of the air condenser VX-1 and then through the water condenser T-4 and chiller T-5 enters the three-phase separator BR-1, where the product off-gases are separated catalysis from aqueous and hydrocarbon phases.

The water phase is discharged into the atmospheric tank E-5, where the sediment and degassing of the aqueous phase occurs, the degassing gases are discharged for afterburning into the P-1 furnace, the hydrocarbon condensate of the catalyzate is fed through the T-6 recuperator to the K-1 debutanizer column. Heat is supplied to the bottom of the column through the RB-1 reboiler.

Stabilized gasoline is removed from the RB-1 column reboiler, and then, passing through the T-6 recuperator, is chilled in the VX-3 air cooler, and delivered to the finished product warehouse. Under the conditions of the achieved uniform heating of each reaction zone, the RON of stabilized gasoline of 92.2 units is achieved. In this case, the yield of C ₅₊ hydrocarbons per supplied hydrocarbon fraction exceeds 70 wt. %. The mass flow of stable gasoline was 3654.4 g / h (yield 73.3 wt% for the supplied hydrocarbon fraction).

Waste gases from the column K-1 are fed to the air condenser VX-2 and then enter the reflux tank BR-2.

The water from the BR-2 tank is discharged into the E-5 atmospheric tank, the reflux from the BR-2 tank is fed to the irrigation of the K-1 column, the excess reflux is discharged, the exhaust gas from the BR-2 tank is fed to the T-8 condenser, where propane condensation ...

Propane is separated from non-condensed gases in the BR-3 tank and is discharged for mixing with excess reflux from the BR-2 tank, and then a mixture of reflux and condensed propane is removed.

Water from the BR-3 tank is discharged into the E-5 atmospheric tank, the exhaust gases from the BR-2, BR-3 tanks are discharged through the T-9 steam heater.

Table 2 shows the chemical compositions of the hydrocarbon fractions used in Examples 1-5. In particular, the 62-85 ° C fraction (fr. 62-85 ° C) is the benzene-forming part of the catalytic reforming feedstock (approximate boiling range 62-85 ° C). The raffinate can be a mixture of predominantly gasoline-range hydrocarbons that have not undergone conversion during the catalytic reforming process. The raffinate can be described as a by-product gasoline cut from an aromatic hydrocarbon extractive distillation unit. For example, the raffinate can be a by-product of the extractive distillation of the benzene-toluene fraction. Also, the raffinate can be a by-product of the extractive distillation of the toluene-xylene fraction.

Table 3 shows the compositions of the olefin-containing fractions used in Examples 1-5. The olefin-containing fractions used can be considered, for example, as a model of a dry gas of catalytic cracking (the compositions of the SGCC were obtained by averaging the data of the refinery over several months of the catalytic cracking unit operation). However, we note that the name and the process of origin of the olefin-containing fractions may vary depending on the enterprise and the region. Attention should be paid to the chemical composition of the fraction used, in particular, the olefin-containing fraction should include C ₂ -C ₄ olefins in a total amount of 10 to 50 wt. %. Preferably, the mass fraction of C ₅₊ hydrocarbons in the olefin-containing fraction is not more than 5.0 wt. %. Preferably, the volume fraction of hydrogen sulfide in the olefin-containing fraction is at most 0.005%. In particular cases, the olefin-containing fraction may contain from 0.5 to 8 wt. % hydrogen, preferably from 2.3 to 8 wt. % hydrogen.

Methanol technical grade "A" GOST 2222-95 was used as oxygenate in examples 1-3. Example 4 used dimethyl ether (DME), 99%. Example 5 uses 95% ethanol.

Table 4 shows the compositions of the zeolite catalysts used in Examples 1-5.

Table 5 shows the conditions and basic parameters of examples 1-5. The hydrocarbon fraction in examples 1-5 is fed to the first reaction zone. The experiments according to the invention were carried out at a pressure of 15-40 bar (1.5-4.0 MPa), preferably 22-27 bar (2.2-2.7 MPa).

Table 6 shows the composition of the liquid hydrocarbon product free of dissolved C ₁ -C ₄ gases (C ₅₊ fraction of the hydrocarbon product fraction) in Examples 1-5. The composition of the C ₅₊ fraction is shown for ease of comparison of experimental results under different conditions. Depending on the requirements of a particular enterprise, the liquid hydrocarbon product may contain different amounts of dissolved gases C ₁ -C ₄ .

Table 7 shows the temperatures of the catalyst bed in the first, second and third reaction zones (R101 / R201 / R301), ° C.

Observations

The heating of the catalyst in each reaction zone does not exceed 35 ° C. The temperature of the catalyzate at the outlet of the reaction zone exceeds the temperature of the feedstock supply.

Moreover, in examples 1-5, a similar result was obtained using adiabatic reactors without the use of internal heat exchangers (for example, tubular reactors). The use of isothermal reactors in laboratory experiments often makes it possible to keep the temperature of the catalyst bed in a narrow range. However, the use of isothermal reactors in industry is difficult to justify from an economic point of view. Built-in heat exchangers add to the cost of manufacturing and maintaining reactor equipment. Therefore, it is important that the proposed method and installation make it possible to abandon the use of internal heat exchangers.

It was found that without the use of distributed supply of both methanol and SGCC according to the proposed invention, heating along the catalyst bed increases to 58-105 ° C, which in particular leads to a drop in the yield of stable gasoline by 7-16 wt. % on the supplied hydrocarbon phase. When heated above 100 ° C, traces of CO ₂ and CO appear in the gaseous product, which indicates the decomposition of methanol.

In the absence of an external heat exchanger system according to the present invention, the temperature of the end catalyst bed falls below the feed temperature for at least one reaction zone, in particular for the third reaction zone. This leads, in particular, to a drop in the RHI of the product by 3-5 units. Obtaining gasoline with OCHI above 90 units. when the content of aromatic hydrocarbons in the product is not more than 35 vol. % and an output of more than 70 wt. %

Examples 1-5 show RON (Research Octane) of a liquid hydrocarbon product of more than 90 units, with a total aromatics content of not more than 37.9 wt. % and benzene content not more than 1.2 wt. %. This provides a product yield of more than 70 wt. % on the supplied hydrocarbon fraction.

At the same time, the content of olefins in the liquid hydrocarbon product during the implementation of the proposed method does not exceed 2.0 wt. %. This is an order of magnitude lower than the maximum olefin content allowed for motor fuels (typically no more than 18 vol.%).

Such a composition makes it possible to use the produced product not only as the main component for the production of high-octane gasolines, but also to sell it independently as commercial gasoline without the need for compounding with other refined products. The ability to use olefin-containing gases with an olefin content of less than 50 wt. % without preliminary concentration of olefins

As a technical result, the possibility of using little-demanded olefin-containing fractions as feedstock for the production of gasoline is also considered. Refineries produce olefin-containing fractions that are used as fuel. These are gases from catalytic cracking, gases from a delayed coking unit, olefin-containing fuel gases of various origins, etc. The content and composition of olefins in such streams is too low for commercially viable recovery. At the same time, the cost of streams burned as fuel is minimal. The involvement of such olefin-containing fractions in the production of gasoline significantly increases the value of the stream for the enterprise.

Examples 1-5 demonstrate the possibility of using olefin-containing fractions with an olefin content of not more than 50 wt. %. In particular, it is possible to use gaseous sources of olefins with an olefin content of 10 wt. % (and more). This result allows for significant cost savings in the production of a unit of product compared to methods that use highly concentrated sources of olefins or chemically pure olefins.

The proposed method allows the use of dilute olefins instead of highly concentrated sources of olefins (eg, pure ethylene). This creates an opportunity as a source of olefins, intermediates and by-products of already existing petrochemical plants. Among them are dry gases of catalytic cracking, various fuel gases with an olefin content of 10 to 50 wt. %.

Possibility to use olefin-containing gases including hydrogen as raw material without additional separation

It was found that the proposed method allows the use of olefin-containing fractions with a high hydrogen content as raw materials. Moreover, the proposed method does not require additional separation of hydrogen from the olefin-containing fraction. In particular, examples 1-5 use olefin-containing fractions with a hydrogen content of 0.5 to 8 wt. %. In this case, the olefin-containing fractions were fed to the reaction zones without preliminary separation of hydrogen from them. This result is important because fuel gases often contain olefins along with significant amounts of hydrogen. However, the presence of hydrogen in the olefin source can lead to side reactions.

Ability to eliminate the recycling of gaseous products

Found that the proposed method eliminates the recycle of gaseous products. All examples according to the proposed method show the conversion of olefins C ₂ -C ₄ feed above 98 wt. %. Such a high degree of conversion in one pass of the feedstock through the reactor makes it possible to eliminate the use of the recycle of gaseous products for the purpose of a more complete processing of the olefins of the feedstock.

Reducing oxygenate consumption by partially replacing oxygenate with olefin-containing gases

It was found that the problem of reducing the consumption of oxygenates for the production of gasoline can be solved by partially replacing oxygenates with low-demand olefin-containing fractions. This approach allows you to reduce the consumption of oxygenates while maintaining the yield and quality of the product. Co-processing of hydrocarbon fractions and oxygenates (without the involvement of olefin-containing fractions) often makes it possible to obtain gasolines with high RON and high product yields. However, oxygenates such as methanol, ethanol, dimethyl ether are rarely available in refineries as cheap feedstocks. When the source of oxygenate cannot be found in the by-product or intermediate product of the enterprise itself, it has to be purchased from outside at the prices of the commercial product. This increases the cost of producing a unit of commercial gasoline, and complicates the production logistics.

At the same time, the proposed method makes it possible to partially replace oxygenates with a source of dilute olefins (olefin-containing fractions). In examples 1-5, the olefin-containing fractions act as a partial replacement for the oxygenates of the feed.

The calculation of the percentage of substitution of oxygenates for olefin-containing fractions is carried out according to formulas (4) - (6) on page 3 of this Description. Examples 1-5 show the possibility of replacing from 43 to 85% oxygenate with olefin-containing fractions while maintaining the yield of the C ₅₊ fraction of the hydrocarbon product of more than 70% for the supplied hydrocarbon fraction, and the RON of the liquid product above 90 units.

The provided formulas (4) - (6) can be applied to the already known methods of joint processing of hydrocarbon fractions and oxygenates (without involving olefin-containing fractions) into gasolines. In this case, formulas (4) - (6) make it possible to calculate the amount (molar flow, mol / h) of oxygenate in a known method, which can be replaced by available olefin-containing fractions without loss of quality and product yield.

Claims (57)

1. A method for producing gasoline, including the following sequential operations: a) a hydrocarbon fraction and an oxygenate are fed to the inlet of the heat exchanger T-1, from the outlet of T-1 a mixture of oxygenate and a hydrocarbon fraction is mixed with an olefin-containing gas and fed to the P-1 furnace for heating, from P-1 the reaction mixture is fed to the first shelf reactor R-1/1 or R-1/2, including a catalyst, b) oxygenate and olefin-containing gas are fed to the inlet of heat exchanger T-2, from the outlet of heat exchanger T-2 a mixture of olefin-containing gas and oxygenate is fed to the second shelf of the reactor R-1/1 or R-1/2, including the catalyst, c) oxygenate and olefin-containing gas are fed to the inlet of the heat exchanger T-3, from the outlet of the heat exchanger T-3 a mixture of olefin-containing gas and oxygenate is fed to the third shelf of the reactor R-1/1 or R-1/2, including the catalyst, d) catalysate from the outlet of the reactor R-1/1 or R-1/2 is distributed into three streams supplied to heat exchangers T-1, T-2 and T-3, catalyzate from the outlet of heat exchangers T-1, T-2, T -3 enters the inlet of the air condenser VX-1 and then through the water condenser T-4 and the chiller T-5 enters the three-phase separator BR-1, where the waste gases of the catalysis product are separated from the aqueous and hydrocarbon phases, e) the aqueous phase is discharged into the atmospheric tank E-5, where sediment and degassing of the aqueous phase occurs, the degassing gases are discharged for afterburning into the P-1 furnace, the hydrocarbon condensate of the catalyzate is fed through the T-6 recuperator to the K-1 debutanizer column, f) heat supply to the bottom of column K-1 is carried out through the RB-1 reboiler, stabilized gasoline is removed from the reboiler of the column RB-1 and then, passing through the recuperator T-6, is cooled down in the air cooler VX-3, is taken to the finished product warehouse, g) off-gases from column K-1 are fed to the air condenser VX-2 and then enter the reflux tank BR-2, h) the water from the BR-2 tank is discharged into the E-5 atmospheric tank, the reflux from the BR-2 tank is fed to the irrigation of the K-1 column, the excess reflux is discharged, the exhaust gas from the BR-2 tank is fed to the T-8 condenser, where condensation of propane, i) propane is separated from non-condensed gases in the BR-3 tank and is discharged for mixing with excess reflux from the BR-2 tank, and then the mixture of reflux and condensed propane is removed, j) water from the BR-3 tank is discharged into the E-5 atmospheric tank, the exhaust gases from the BR-2, BR-3 tanks are discharged through the T-9 steam heater, moreover, the heating of the catalyst bed in each reaction zone is from 5 to 35 ° C, for each shelf of the reactor, the temperature of the catalyzate at the outlet of the catalyst bed exceeds the temperature of the feedstock supply to this shelf of the reactor. 2. The method according to claim 1, in which a hydrocarbon fraction is supplied to the input of the heat exchanger T-1 from the supply tank E-1 by the pump N-1/1 or N-1/2, additionally from the supply tank E-3 by the pump N-3 / 1 or H-3/2 oxygenate is fed to the input of the heat exchanger T-1, from the outlet of T-1 the mixture of oxygenate and hydrocarbon fraction is mixed with the olefin-containing gas supplied by the KP-1 compressor and fed to the P-1 furnace for heating, from P-1 the reaction mixture is fed to the first shelf of the reactor R-1/1 or R-1/2, the oxygenate from the tank E-3 and olefin containing gas from the KP-1 compressor, from the outlet of the T-2 heat exchanger, a mixture of olefin-containing gas and oxygenate is fed to the second shelf of the R-1/1 or R-1/2 reactor, to the inlet of the T-3 heat exchanger by the N-3/1 pump or H-3/2 oxygenate is fed from the E-3 tank and olefin-containing gas from the KP-1 compressor, from the outlet of the T-3 heat exchanger a mixture of olefin-containing gas and oxygenate is fed to the third shelf of the R-1/1 or R- reactor 1/2. 3. The method according to claim. 1, in which the aqueous phase is discharged into the atmospheric container E-5, operating if necessary with heating. 4. The method according to claim 1, in which the reflux from the BR-2 tank is fed by the pump N-4/1 or N-4/2 to the irrigation of the column K-1, the excess reflux from the pump N-4/1 or N-4 / 2 is discharged to the LPG warehouse, the off-gas from the BR-2 tank is fed to the T-8 condenser, where propane condensation occurs, propane is separated from the non-condensed gases in the BR-3 tank and pump N-5/1 or N-5/2 is discharged to mixing with excess reflux from the tank BR-2 and then the mixture of reflux and condensed propane is fed to the LPG warehouse, and the exhaust gases from the tanks BR-2, BR-3 are discharged into the fuel gas network through the steam heater T-9. 5. The method according to claim 1, in which the catalyst regeneration is carried out in a stream of nitrogen-air mixture along the circuit, and the nitrogen-air mixture is supplied from the outlet of the reactor P1 / 1.2 to the recuperator T-10, then from the outlet of the T-10 it is fed to the inlet of the air cooler of the regeneration gases VX-4, then from the outlet VX-4 it is fed to the regeneration gas separator S-1, then from the outlet C-1 it is fed to the regeneration gas compressor KP-2, then from the outlet KP-2 it is fed to the recuperator T- 10, then from the T-10 outlet it is fed to the P-2 regeneration gas heating furnace, then from the P-2 outlet it is fed to the P1 / 1.2 reactor inlet. 6. The method according to claim 5, in which the frequency of circulation to feed with nitrogen and air through the supply lines of nitrogen and air is from 10 to 50 and the regeneration is carried out at a pressure of 6-10 bar. 7. A method for producing gasolines according to claim 1, wherein the zeolite catalyst comprises: a) a zeolite of the ZSM-5 type with a SiO ₂ / Al ₂ O ₃ modulus from 43 to 95 in an amount from 65 to 80 wt. %, b) sodium oxide in an amount from 0.04 to 0.15 wt. %, c) zinc oxide in the amount of 1.0-5.5 wt. %, d) oxides of rare earth elements in a total amount of 0.5-5.0 wt. %, e) a binder comprising silicon dioxide, alumina, or mixtures thereof. 8. A method according to claim 1, wherein the process pressure is between 1.5 and 4.0 MPa, preferably between 2.2 and 2.7 MPa. 9. The method according to claim 1, wherein the temperature of the stream at the inlet to the first / second / third reaction zones is 340-370 ° C / 341-370 ° C / 350-376 ° C. 10. The method according to p. 1, in which the distribution of oxygenate between the three reaction zones is 44-62 wt. % / 18-35 wt. % / 14-21 wt. % of the total amount of oxygenate. 11. The method according to p. 1, in which the distribution of the olefin-containing gas between the three reaction zones is 15-30 wt. % / 20-40 wt. % / 30-60 wt. % of the total amount of olefin-containing gas. 12. The method according to p. 1, in which the hydrocarbon fraction contains normal paraffins in an amount of 15-24 wt. %, isoparaffins in the amount of 28-56 wt. %, naphthenes in the amount of 22-40 wt. %, the rest is aromatic hydrocarbons and olefins. 13. The method according to claim 1, in which the hydrocarbon fraction can be selected from the group consisting of straight-run gasoline, stable gas gasoline, light gas condensate, gasoline fraction with a boiling range of about 62-85 ° C, raffinate, and mixtures thereof. 14. The method according to p. 1, in which the oxygenate is selected from the group consisting of aliphatic alcohols, for example, methanol, ethanol, crude methanol, technical methanol, ethanol; ethers, for example dimethyl ether, as well as mixtures thereof, including with water. 15. The method according to p. 1, in which the olefin-containing fraction includes 2.3-8.0 wt. % hydrogen. 16. The method of claim 1, wherein the olefin-containing fraction is selected from the group consisting of dry catalytic cracking gas, catalytic cracking wet gas, other catalytic cracking gases and their fractionation products, coking unit off-gas, Fischer-Tropsch synthesis gases, as well as mixtures thereof. 17. The method of claim 1, wherein the reaction is carried out in the gas phase in a fixed bed of catalyst. 18. The method of claim 1, wherein built-in heat exchangers are not used. 19. The method of claim 1, wherein each reactor is an adiabatic reactor. 20. Installation for implementing the method for producing gasoline according to claim 1, including: a) heat exchanger T-1 for heating the hydrocarbon fraction and oxygenate, heat exchangers T-2 and T-3 for heating oxygenate and olefin-containing gas, b) furnace P-1 for heating a mixture of hydrocarbon fraction, oxygenate and olefin-containing gas, the inlet of which is connected to the first outlet of the heat exchanger T-1, c) reactor R-1/1 or R-1/2, and the first inlet of the reactor is connected to the outlet of the P-1 furnace, the second inlet of the reactor is connected to the first outlet of the heat exchanger T-2, the third inlet of the reactor is connected to the first outlet of the heat exchanger T-3, the reactor outlet is connected to the inlets of the heat exchangers T-1, T-2 and T-3, d) air condenser VX-1 for receiving the catalyst from the second outlet of the heat exchanger T-1, from the second outlet of the heat exchanger T-2 and from the second outlet of the heat exchanger T-3, water condenser T-4 for additional cooling of the catalyst after the air condenser VX-1, chiller Т-5 for cooling the catalyzate after the condenser Т-4, e) three-phase separator BR-1 for feeding the catalyzate from the chiller T-5, where the off-gases of the catalysis product are separated from the aqueous and hydrocarbon phases, f) atmospheric tank E-5 for sediment and degassing of the aqueous phase with BR-1, BR-2, BR-3, g) line for return of degassing gases from E-5 for afterburning to furnace P-1, recuperator T-6 for stabilized gasoline from BR-1 and air cooler VX-3 for additional cooling of gasoline, h) debutanizer column K-1 for stabilizing the catalytic converter from T-6, reboiler RB-1 of column K-1 for supplying heat and removing stabilized gasoline, i) air cooler ВХ-3 for cooling stabilized gasoline, air condenser ВХ-2 for partial condensation of the upper product from the column К-1, j) reflux tank BR-2 for separation of non-condensed gases, overhead product and water condensate, k) condenser T-8 for off-gas from tank BR-2, where propane condensation takes place, l) tank BR-3 for separating propane from non-condensed gases, steam heater T-9 for waste gases from tanks BR-2, BR-3. 21. An installation according to claim 20, additionally including a supply container for the hydrocarbon fraction E-1, a pump N-1/1 or N-1/2 for supplying a hydrocarbon fraction, a supply container E-3 for oxygenate, a pump N-3/1 or Н-3/2 for oxygenate supply, compressor KP-1 for olefin-containing gas. 22. Installation according to claim 20, additionally including pump Н-4/1 or Н-4/2 for reflux from the tank BR-2 for irrigation of column К-1, LPG storage for excess reflux from pump Н-4/1 or Н -4/2 and for a mixture of reflux and condensed propane, pump H-5/1 or H-5/2 for removing propane for mixing with excess reflux from the BR-2 tank. 23. Installation according to claim 20, additionally including equipment for regeneration: recuperator T-10, air cooler for regeneration gases VKh-4, separator for regeneration gas S-1, compressor for regeneration gas KP-2, furnace for heating regeneration gas P-2. 24. Installation according to claim 20, which does not use built-in heat exchangers. 25. An installation according to claim 20, wherein each reactor is an adiabatic reactor.

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