CN110964562A

CN110964562A - Up-flow hydrogenation reaction process combination method for different hydrocarbon materials

Info

Publication number: CN110964562A
Application number: CN201811201910.0A
Authority: CN
Inventors: 何巨堂
Original assignee: Individual
Current assignee: Individual
Priority date: 2018-09-30
Filing date: 2018-09-30
Publication date: 2020-04-07

Abstract

A combination method of up-flow hydrogenation reaction processes of different hydrocarbon materials is provided, wherein a first up-flow hydrogenation reaction process 1R of the hydrocarbon material is combined with a second up-flow hydrogenation reaction process 2R of the hydrocarbon material, usually, hot high-molecular gas of 1RP of a 1R reaction product enters 2R, condensed oil of hot high-molecular gas of 2RP of the 2R reaction product can be circulated to 1R or 2R, hydrogen and hydrogen supply solvent components are repeatedly used, the hydrogen flow is simplified, the investment is reduced, the energy consumption is reduced, and the overall process economy is improved; according to different requirements, on the premise of ensuring that heavy components with the specified boiling points above in the target hydrogenation product are not mixed basically, the separation processes of the reaction products of 1R, 2R and the two products are combined or shared to the maximum extent; the different hydrocarbon materials are selected from coal, coal and oil mixture, heavy oil with high aromaticity, heavy oil with low aromaticity, etc.

Description

Up-flow hydrogenation reaction process combination method for different hydrocarbon materials

Technical Field

The invention relates to a method for combining the upflow hydrogenation processes of different hydrocarbon materials, wherein the upflow hydrogenation process 1R of the first hydrocarbon material is combined with the upflow hydrogenation process 2R of the second hydrocarbon material, usually the hot high-molecular gas of the 1RP of the 1R reaction product enters 2R, the condensed oil of the hot high-molecular gas of the 2RP of the 2R reaction product can be circulated to 1R or 2R, and the hydrogen and hydrogen supply solvent components are repeatedly used, so that the hydrogen flow is simplified, the investment is reduced, the energy consumption is reduced, and the overall process economy is improved; according to different requirements, on the premise of ensuring that heavy components with the specified boiling points above in the target hydrogenation product are not mixed basically, the separation processes of the reaction products of 1R, 2R and the two products are combined or shared to the maximum extent; the different hydrocarbon materials are selected from coal, coal and oil mixture, heavy oil with high aromaticity, heavy oil with low aromaticity, etc.

Background

The basic idea of the invention is: a combination method of up-flow hydrogenation reaction processes of different hydrocarbon materials is provided, wherein a first up-flow hydrogenation reaction process 1R of the hydrocarbon material is combined with a second up-flow hydrogenation reaction process 2R of the hydrocarbon material, usually, hot high-molecular gas of 1RP of a 1R reaction product enters 2R, condensed oil of hot high-molecular gas of 2RP of the 2R reaction product can be circulated to 1R or 2R, hydrogen and hydrogen supply solvent components are repeatedly used, the hydrogen flow is simplified, the investment is reduced, the energy consumption is reduced, and the overall process economy is improved; according to different requirements, on the premise of ensuring that heavy components with the specified boiling points above in the target hydrogenation product are not mixed basically, the separation processes of the reaction products of 1R, 2R and the two products are combined or shared to the maximum extent; the different hydrocarbon materials are selected from coal, coal and oil mixture, heavy oil with high aromaticity, heavy oil with low aromaticity, etc.

The invention relates to a method for combining upflow hydrogenation reaction processes of different hydrocarbon feeds, which can combine 2 or more upflow hydrogenation reaction processes of different hydrocarbon ratio raw materials (coal or oil), wherein the combined upflow hydrogenation reaction processes can be 2 or more, but the stream relationship among the combined upflow hydrogenation reaction processes can be described by the stream relationship among the combined 2 upflow hydrogenation reaction processes R10 and R20. The invention is applicable at least to the following situations:

① the first hydrocarbon is coal and the second hydrocarbon is coal and oil;

② the first hydrocarbon is coal and the second hydrocarbon is oil;

③ the first hydrocarbon is coal and oil and the second hydrocarbon is coal;

④ the first hydrocarbon is coal and oil and the second hydrocarbon is oil;

⑤ the first hydrocarbon is oil and the second hydrocarbon is coal;

⑥ the first hydrocarbon is oil and the second hydrocarbon is coal and oil;

⑦ the first hydrocarbon is a high aromatic oil and the second hydrocarbon is a low aromatic oil;

⑧ the first hydrocarbon is a low aromatic oil and the second hydrocarbon is a high aromatic oil.

As an application example of the invention, 2 paths of upflow hydrogenation reaction processes containing asphalt materials with different aromaticity can be combined, moderate hydrogenation modification reaction of high-temperature coal tar can be carried out at 1R to produce high-quality needle coke raw material, and deep hydrogenation thermal cracking of medium-low temperature coal tar can be carried out at 2R to produce light distillate oil.

The combined objectives of the invention are:

① hydrogen is reused, hydrogen flow is simplified, and 1R reaction product 1RP hot high-molecular gas enters 2R;

② the hydrogen solvent component is reused, the 1R reaction product 1RP carrying the hydrogen donor hot high-molecular gas enters 2R, or the 2R reaction product 2HRP hot high-molecular gas containing the hydrogen donor condensed oil can be recycled to 1R or 2R;

③ the separation and fractionation systems of the hydrogenation reaction products are combined, and on the premise of ensuring that heavy components with the boiling points above the specified boiling point in the target hydrogenation products are not mixed basically, the separation processes of the reaction products of 1R, 2R and the two products are combined or shared to the utmost extent, thereby simplifying the separation and fractionation process and reducing the investment.

The invention can be used in combination with other processes.

The invention has the distinct characteristic that the clear concept of '2 or more reaction sections connected in series' is used, the hydrogen is repeatedly used in series, the hydrogen flow is simplified, and the investment is reduced.

There are many methods related to the upflow hydrogenation process of coal, mixture of coal and oil, heavy oil with high aromaticity, heavy oil with low aromaticity, but none of them suggest the technical solution of the present invention, and the following are several related patent methods or patent application methods:

① Chinese patent ZL201110167634.2 relates to a process for preparing needle coke raw material by using medium and low temperature tar and high temperature asphalt, which relates to a process for producing needle coke raw material by hydrogenation of coal tar containing coal asphalt, but does not provide a technical scheme similar to the invention;

② Chinese patent ZL201110167667.7 relates to a process for preparing needle coke raw material by using medium and low temperature tar, which relates to the process of producing needle coke raw material by hydrogenation of coal tar containing coal pitch, but does not provide a technical scheme similar to the invention;

③ patent ZL201110171211.8 centrifugal method of coal tar purification and coal tar preparation needle coke raw material process, relate to the process of producing needle coke raw material of coal tar hydrogenation containing coal pitch, but does not propose the technical scheme similar to this invention;

④ Chinese patent ZL201110284485.8 relates to a process for preparing needle coke raw material by coal tar pitch, which relates to a process for producing needle coke raw material by coal tar containing coal pitch through hydrogenation, but does not provide a technical scheme similar to the invention;

⑤ Chinese patent ZL201110339693.3 is a process for preparing needle coke raw material by utilizing coal tar and heavy phase circulation, which relates to the process of producing needle coke raw material by hydrogenation of coal tar containing coal pitch, but does not provide a technical scheme similar to the invention;

⑥ Chinese patent ZL201110339694.8 is a process for preparing needle coke raw material by combining coal tar pitch with heavy phase circulation, which relates to the process of producing needle coke raw material by hydrogenation of coal tar containing coal pitch, but does not provide a technical scheme similar to the invention;

⑦ Chinese patent ZL201310231391.3 relates to a process for producing needle coke, which relates to a process for producing needle coke raw materials by hydrogenation of coal tar containing coal pitch, but does not provide a technical scheme similar to the invention;

⑧ Chinese patent ZL 201310231394.7 relates to a process for producing needle coke raw material by hydrogenation of coal tar containing coal pitch, but does not provide a technical scheme similar to the invention;

⑨ patent ZL201610902143.0 relates to a method and a system for producing light fuel and needle coke by coal tar, which relates to a process for producing needle coke raw materials by coal tar suspension bed hydrogenation containing coal pitch, but does not provide a technical scheme similar to the invention;

⑩ patent ZL 201610902769.1 discloses a method and a system for producing light fuel and needle coke in a maximized mode by coal tar, and relates to a process for producing needle coke raw materials by coal tar-containing pitch through coal tar suspension bed hydrogenation.

There are many methods for producing light fuel oil by hydrogenation of coal tar containing coal pitch, but none of them proposes the technical scheme of the invention, and the following are several related patent methods or patent application methods:

① A method for preparing heavy oil from coal tar by hydrogenation in suspension bed reactor, which comprises using all coal tar components (including fraction with conventional boiling point lower than 370 deg.C and fraction with conventional boiling point higher than 370 deg.C) as raw material for hydrocracking, and using suspension bed or bubbling bed reactor for hydrocracking, wherein the method introduces all hydrocarbon components with conventional boiling point lower than 330 deg.C into hydrocracking, which has harmful process defects;

② chinese patent ZL201010217358.1 describes a coal tar heavy oil hydrogenation lightening method using a suspension bed hydrogenation reactor, which comprises coal tar raw material pretreatment and distillation separation, coal tar heavy fraction suspension bed hydrocracking and light distillate conventional upgrading process.zl 201010217358.1 is characterized in that a suspension bed or bubbling bed hydrocracking reactor is used to perform a hydrocracking reaction process R1 by using a fraction of coal tar at a temperature higher than 370 ℃, most (about 80%) of heavy oil (pitch) with a reaction product higher than 370 ℃ is directly recycled back to be cracked, a small part (about 20%) of the heavy oil (pitch) is subjected to solid-liquid separation, and the solid catalyst is separated and then recycled back to be cracked, so that large molecular pitch in coal tar is cracked into small molecular light oil products as much as possible, the separated catalyst is thrown out, and the purpose of the external throwing is to remove a small amount of high molecular polymer and deactivated catalyst generated in the cracking process;

③ Chinese patent ZL201210022921.9 proposes a hydrogenation and lightening method of heavy oil with low hydrogen content using hydrogen supply hydrocarbon, for all fractions of coal tar higher than 330 ℃, the coal tar heavy oil with conventional boiling point higher than 450 ℃ directly enters an expansion bed (such as a suspension bed or a boiling bed) hydrogenation and thermal cracking reaction zone HPU21, and simultaneously the coal tar distillate with conventional boiling point of 330-450 ℃ is used as a precursor of hydrogen supply solvent oil to pass through a hydrogenation modification reaction zone HPU1 to produce hydrogen supply hydrocarbon material flow SHS mainly composed of hydrocarbon components with conventional boiling point of 330-450 ℃, and then the SHS is introduced into the hydrogenation and thermal cracking reaction zone HPU21, which obviously improves KL ratio of hydrogen supply hydrocarbon weight to fresh raw material F10X weight in the SHS flow reaction zone, thereby having obvious effects of reducing coking condensation speed, improving liquid product yield in coal tar heavy oil hydroconversion process, improving product quality, reducing reaction temperature rise, and enhancing device operation stability and safety;

④ Chinese patent ZL200810166719.7 describes a combined hydro-conversion method of coal tar fractions with different boiling ranges, a first hydrocarbon fraction containing a coal tar light fraction with a conventional boiling point lower than 390 ℃ is converted in a first hydro-refining reaction part, a first hydro-refining reaction effluent and a second hydrocarbon fraction containing a coal tar heavy fraction with a conventional boiling point higher than 370 ℃ are converted in a second hydro-refining reaction part, a second hydro-refining reaction effluent is separated and a product is recovered, Chinese patent ZL200810166719.7 performs combined hydro-conversion on the coal tar fractions with different boiling ranges to form more proper hydro-refining reaction temperature distribution, has the advantages of improving the product quality, stabilizing the operation, prolonging the operation period and the like, and is particularly suitable for small-scale classified combined hydro-conversion of medium-high temperature coal tar wide fractions;

⑤ CN105623724A discloses a hydro-thermal cracking method for producing low carbon number single six-membered ring hydrocarbon from high aromatic hydrocarbon, which can economically produce C in a large amount by using medium and low temperature coal tar₆～C₁₀The cyclohexane hydrocarbon or benzene hydrocarbon is catalytically reformed and aromatic extracted to obtain benzene hydrocarbon. For the coal tar hydrogenation thermal cracking reaction process R10, the process recovers a first hydrogenation reaction effluent R10P to obtain a material flow containing first hydrogenation product oil, the first hydrogenation product oil hydrocarbon with the normal boiling point higher than 350 ℃ is returned to the first hydrogenation reaction process R10 to be contacted with a first hydrogenation catalyst R10C, at least one part of the first hydrogenation product oil hydrocarbon with the normal boiling point of 260-350 ℃ is returned to the first hydrogenation reaction process R10 to be contacted with a first hydrogenation catalyst R10C, and the first hydrogenation reaction flowThe conventional, particularly high boiling, thermally liable condensed hydrocarbon components in the effluent R10P are no longer returned to the first hydrogenation process R10; the method does not propose a technical solution similar to the present invention.

The method of the invention is not reported.

The invention aims to provide a method for combining upflow hydrogenation reaction processes of different hydrocarbon materials.

Disclosure of Invention

The invention discloses a combined method of upflow hydrogenation processes of different hydrocarbon materials, which is characterized by comprising the following steps:

the combination method of the upflow hydrogenation processes of different hydrocarbon materials comprises the steps that the upflow hydrogenation process 1R of a first hydrocarbon material 1RF is combined with the upflow hydrogenation process 2R of a second hydrocarbon material 2 RF;

the different hydrocarbon feeds 1RF, 2RF are selected from coal, mixtures of coal and oil, heavy oil with high aromaticity, heavy oil with low aromaticity; the heavy oil refers to hydrocarbon oil containing hydrocarbon components having a conventional boiling point of more than 450 ℃;

in the thermal high-pressure separation process 1RP-THPS, the reaction product 1RP in the upflow hydrogenation reaction process 1R is separated into thermal high-pressure gas 1RP-THPS-V and thermal high-pressure oil 1 RP-THPS-L;

a hydrogen-containing stream of hot high-gas 1RP-THPS-V based on the reaction product 1RP of the upflow hydrogenation process 1R is passed into an upflow hydrogenation process 2R for contact with a second hydrocarbon 2RF or its hydroconverter.

The invention comprises 3 upflow hydrogenation reaction processes, and is characterized in that:

the combination method of the up-flow hydrogenation reaction processes of different hydrocarbon materials comprises the steps that an up-flow hydrogenation reaction process 1R of a first hydrocarbon material 1RF, an up-flow hydrogenation reaction process 2R of a second hydrocarbon material 2RF and an up-flow hydrogenation reaction process 3R of a third hydrocarbon material 3RF are combined;

the different hydrocarbon feedstocks 1RF, 2RF, 3RF are selected from coal, mixtures of coal and oil, heavy oil with high aromaticity, heavy oil with low aromaticity; the heavy oil refers to hydrocarbon oil containing hydrocarbon components having a conventional boiling point of more than 450 ℃;

1RP-THPS is separated in the thermal high-pressure separation process, and 1RP of a reaction product in the upflow hydrogenation reaction process 1R is separated into thermal high-pressure gas 1RP-THPS-V and thermal high-pressure oil 1 RP-THPS-L;

feeding a hydrogen-containing stream of hot high-gas 1RP-THPS-V based on the reaction product 1RP of the upflow hydrogenation process 1R into an upflow hydrogenation process 2R for contact with a second hydrocarbon 2RF or its hydroconverter;

in the thermal high-pressure separation process of 2RP-THPS, the reaction product 2RP in the upflow hydrogenation reaction process of 2R is separated into thermal high-pressure gas 2RP-THPS-V and thermal high-pressure oil 2 RP-THPS-L;

a hydrogen-containing stream of hot high-molecular gas 2RP-THPS-V based on the reaction product 2RP of the upflow hydrogenation process 2R is passed into an upflow hydrogenation process 3R for contact with a third hydrocarbon feed 3RF or its hydroconverter.

The invention comprises 4 upflow hydrogenation reaction processes, and is characterized in that:

the combination method of the up-flow hydrogenation reaction processes of different hydrocarbon materials comprises the steps that the up-flow hydrogenation reaction process 1R of a first hydrocarbon material 1RF, the up-flow hydrogenation reaction process 2R of a second hydrocarbon material 2RF, the up-flow hydrogenation reaction process 3R of a third hydrocarbon material 3RF are combined, and the up-flow hydrogenation reaction process 4R of a fourth hydrocarbon material 4RF are combined;

the different hydrocarbon feedstocks 1RF, 2RF, 3RF, 4RF are selected from coal, mixtures of coal and oil, heavy oil with high aromaticity, heavy oil with low aromaticity; the heavy oil refers to hydrocarbon oil containing hydrocarbon components having a conventional boiling point of more than 450 ℃;

feeding a hydrogen-containing stream of hot high-molecular gas 2RP-THPS-V based on the reaction product 2RP of the upflow hydrogenation process 2R into an upflow hydrogenation process 3R for contact with a third hydrocarbon feed 3RF or its hydroconverter;

in the thermal high-pressure separation process of 3RP-THPS, the reaction product 3RP of the upflow hydrogenation reaction process 3R is separated into thermal high-pressure gas 3RP-THPS-V and thermal high-pressure oil 3 RP-THPS-L;

a hydrogen-containing stream of hot high-molecular gas 3RP-THPS-V based on the 3R reaction product 3RP is brought into contact with the fourth hydrocarbon feed 4RF or its hydroconverter.

In the invention, the upflow hydrogenation reaction process in which hydrogen is used in series takes the hydrogen flowing direction as the positive direction, and the working mode can be selected from 1 or the combination of several of the following:

① the hot high-molecular gas carrying hydrogen-donor of the reaction product in the upstream reaction process enters the downstream reaction process to contact with the raw oil or hydrogenated transformation product thereof in the downstream reaction process, and the hydrogen-donor solvent component is reused;

② reusing hydrogen solvent component to recycle the condensed oil containing hydrogen donor of hot high-molecular gas of reaction product in downstream reaction process to upstream reaction process, contacting with the raw oil or hydrogenated and converted product thereof in upstream reaction process, and reusing hydrogen solvent component;

③ the separation and fractionation system of the hydrogenation reaction products is combined, and the gas obtained in the gas-liquid separation process of the reaction products of 2 or more reaction processes or the gas and hydrocarbon oil steam are processed jointly on the premise of ensuring that the heavy components with the specified boiling point or above in the target hydrogenation products, namely the hydrocarbon components with the conventional boiling point higher than 330 ℃, are not mixed basically;

2 or more reaction products of the reaction process, or jointly processing the gas or the gas and hydrocarbon oil steam obtained in the gas-liquid separation process, wherein a fractionating tower can be commonly used;

④ the separation and fractionation system of the hydrogenation reaction products is combined, and the gas obtained in the gas-liquid separation process of the reaction products of 2 or more reaction processes or the gas and hydrocarbon oil steam are processed jointly on the premise of ensuring that the heavy components with the specified boiling point or above in the target hydrogenation products, namely the hydrocarbon components with the conventional boiling point higher than 380 ℃ are not mixed basically;

⑤ the separation and fractionation system of the hydrogenation reaction products is combined, and the gas obtained in the gas-liquid separation process of the reaction products of 2 or more reaction processes or the gas and hydrocarbon oil steam are processed jointly on the premise of ensuring that the heavy components with the specified boiling point or above in the target hydrogenation products, namely the hydrocarbon components with the conventional boiling point higher than 450 ℃, are not mixed basically;

⑥ the separation and fractionation system of the hydrogenation reaction products is combined, and the gas obtained in the gas-liquid separation process of the reaction products of 2 or more reaction processes or the gas and hydrocarbon oil steam are processed jointly on the premise of ensuring that the heavy components with the specified boiling point above in the target hydrogenation products, namely the hydrocarbon components with the conventional boiling point higher than 500 ℃, are not mixed basically;

⑦ the separation and fractionation system of the hydrogenation reaction products is combined, and the gas obtained in the gas-liquid separation process of the reaction products of 2 or more reaction processes or the gas and hydrocarbon oil steam are processed jointly on the premise of ensuring that the heavy components with the specified boiling point or above in the target hydrogenation products, namely the hydrocarbon components with the conventional boiling point higher than 550 ℃, are not mixed basically;

the gas obtained in the gas-liquid separation of the reaction products of 2 or more reaction processes or the combined processing with hydrocarbon oil vapor may be used together with a fractionating tower.

In the invention, at least 1 upflow hydrogenation reaction process can receive the material containing the hydrogen donor in the upflow hydrogenation reaction process in which the hydrogen is used in series.

In the upflow type hydrogenation reaction processes in which the hydrogen is used in series, at least 1 upflow type hydrogenation reaction process can receive the material containing the hydrogen donor with the conventional boiling point of 230-400 ℃.

In the present invention, the different hydrocarbon feedstock may be selected from coal, a mixture of coal and oil, a heavy oil with high aromaticity, a heavy oil with low aromaticity, which means a hydrocarbon oil comprising hydrocarbon components with a conventional boiling point above 530 ℃.

In the present invention, the different hydrocarbon feedstock may be selected from coal, a mixture of coal and oil, a high aromaticity heavy oil, a low aromaticity heavy oil, which refers to a hydrocarbon oil comprising hydrocarbon components having a conventional boiling point above 550 ℃.

In the present invention, the different hydrocarbon feedstock may be selected from coal, a mixture of coal and oil, a heavy oil with high aromaticity, a heavy oil with low aromaticity, which means a hydrocarbon oil comprising hydrocarbon components with a conventional boiling point above 450 ℃;

the at least one hydrocarbon material refers to a hydrocarbon oil comprising coal tar.

the at least one hydrocarbon material refers to 2 different heavy oils, one heavy oil being a hydrocarbon oil containing high temperature coal tar, and the other heavy oil being a hydrocarbon oil containing medium and low temperature coal tar.

In the invention, generally, the hydrogen is used in series in the upflow hydrogenation reaction process, the reaction product of the downstream-most upflow hydrogenation reaction process is separated by taking the hydrogen flowing direction as the forward direction to obtain the hydrogen-rich gas RHPV, and at least a part of the hydrogen-rich gas RHPV returns to the upstream upflow hydrogenation reaction process for recycling.

In the invention, generally, the hydrogen gas is used in series in the upflow hydrogenation reaction process, the hydrogen gas flow direction is used as the forward direction, the reaction product in the downstream-most upflow hydrogenation reaction process is separated to obtain the hydrogen-rich gas RHPV, and at least a part of the hydrogen-rich gas RHPV is returned to the upstream-most upflow hydrogenation reaction process for recycling.

According to the invention, in general, the upflow hydrogenation reaction process in which the hydrogen is used in series takes the hydrogen flowing direction as the forward direction, the reaction product of the downstream-most upflow hydrogenation reaction process is separated to obtain the hydrogen-rich gas RHPV, at least a part of the hydrogen-rich gas RHPV returns to the upstream upflow hydrogenation reaction process for recycling, and the hydrogen volume concentration is usually more than 70% and more than 85%.

In the invention, generally, the operating temperature of the thermal high-pressure separation process of the reaction product in the existing upflow hydrogenation reaction process is 200-480 ℃, and the operating pressure is 6-30 MPa.

In the invention, generally, the operating temperature of the thermal high-pressure separation process of the reaction product in the existing upflow hydrogenation reaction process is 300-460 ℃, and the operating pressure is 10-25 MPa.

In the invention, preferably, the operating temperature of the reaction product of the upflow hydrogenation reaction process in the thermal high-pressure separation process is 350-430 ℃ and the operating pressure is 13-20 MPa.

The upflow reactor used in the upflow hydrogenation reaction process can work in a mode selected from 1 or a combination of several of the following modes:

① suspension bed hydrogenation reactor;

② boiling bed hydrogenation reactor, discharging the catalyst with reduced activity from the bottom of the bed, and supplementing fresh catalyst from the upper part of the bed to maintain the catalyst inventory in the bed;

③ combined hydrogenation reactor of suspension bed and boiling bed;

④ micro-expanded bed.

Detailed Description

The present invention is described in detail below.

The pressure in the present invention refers to absolute pressure.

The conventional boiling point of the invention refers to the vapor-liquid equilibrium temperature of a substance at one atmospheric pressure.

The conventional boiling range as referred to herein refers to the conventional boiling range of the distillate fraction.

The specific gravity of the present invention refers to the ratio of the density of a liquid at ordinary pressure and 15.6 ℃ to the density of a liquid at ordinary pressure and 15.6 ℃ unless otherwise specified.

The compositions or concentrations or amounts or yield values of the components described herein are weight basis values unless otherwise specified.

The conventional gaseous hydrocarbon refers to hydrocarbon which is gaseous under conventional conditions, and comprises methane, ethane, propane and butane.

The conventional liquid hydrocarbon refers to hydrocarbon which is liquid under conventional conditions, and includes pentane and hydrocarbon with higher boiling point.

The impurity elements in the invention refer to non-hydrogen, non-carbon and non-metal components in the raw oil, such as oxygen, sulfur, nitrogen, chlorine and the like.

The impurity component in the invention refers to the hydrogenation conversion product of non-hydrocarbon component in the raw oil, such as water, ammonia, hydrogen sulfide, hydrogen chloride and the like.

The naphtha component of the present invention refers to conventional liquid hydrocarbons having a conventional boiling point of less than 200 ℃.

The conventional boiling point of the hydrocarbon contained in the diesel component is usually 155-375 ℃, and the conventional boiling point is usually 200-350 ℃.

The normal boiling point of the hydrocarbon contained in the wax oil component is generally 350-575 ℃ and generally 370-530 ℃.

The heavy oil component of the present invention contains hydrocarbons having a conventional boiling point generally greater than 350 c, generally greater than 450 c, specifically greater than 530 c, and more specifically greater than 575 c.

The atmospheric resid component of the present invention, typically an atmospheric fractionation tower bottoms, contains hydrocarbons having a conventional boiling point typically greater than 330 c, typically greater than 350 c, and particularly greater than 370 c.

The vacuum residue component of the present invention, typically a vacuum fractionation tower bottoms, typically contains hydrocarbons having a conventional boiling point generally greater than 450 c, typically greater than 530 c, and particularly greater than 575 c.

The medium hydrocarbon refers to hydrocarbon with a conventional boiling point of 230-400 ℃.

The heavy hydrocarbon refers to hydrocarbon with a conventional boiling point higher than 350 ℃.

The gas-liquid volume ratio or the hydrogen-oil volume ratio in the hydrogenation reaction process refers to the ratio of the volume flow of the hydrogen in the standard state to the volume flow of the specified oil material flow at normal pressure and 20 ℃.

The said low carbon number single six-membered ring hydrocarbon in the present invention refers to C₆～C₉The benzene-series hydrocarbon or cyclohexane-series hydrocarbon has a normal boiling point of usually 70 to 180 ℃, and is suitable for being used as raw material naphtha for preparing aromatic hydrocarbon by catalytic reforming.

The aromatic hydrocarbon with a double ring structure in the invention refers to hydrocarbons containing two ring structures, wherein at least one ring belongs to aromatic rings, such as naphthalene, tetrahydronaphthalene and hydrocarbons with side chains.

The tricyclic aromatic hydrocarbon refers to a hydrocarbon containing three ring structures, at least one of which belongs to an aromatic ring, such as fluorene, dibenzofuran, dibenzothiophene, carbazole, dibenzopyridine, anthracene, phenanthrene, and side chain hydrocarbons thereof or partially hydrogenated saturated products thereof.

The polycyclic aromatic hydrocarbon of the present invention is a hydrocarbon having four or more ring structures, at least one of which belongs to an aromatic ring.

The high aromatic hydrocarbon refers to a hydrocarbon material with high aromatic carbon rate, generally refers to a hydrocarbon material with the aromatic carbon rate higher than 40%, and particularly refers to an oil product with high aromatic concentration containing tricyclic aromatic hydrocarbon or polycyclic aromatic hydrocarbon, such as coal tar distillate, coal hydrogenation direct liquefaction oil distillate or hydrogenation modified oil based on the coal tar distillate and the coal hydrogenation direct liquefaction oil distillate, wherein the main product of such a high aromatic hydrocarbon hydrogenation thermal cracking process can be low-carbon single six-membered ring hydrocarbon.

The aromatic ring number of the polycyclic aromatic hydrocarbon is more than or equal to 3.

The hydrogenation reaction space, which refers to a process fluid flow space where the hydrogenation reaction takes place, may be a reaction inner space such as a hollow cylinder reactor zone, a gas stripping hydrogen mixing zone, a liquid collecting cup upper space region, etc., and may be a reactor outer space such as a pipe inner space, a valve inner space, a mixer inner space, a pump inner space, etc.

According to the upflow hydrogenation reactor, the macroscopic flow leading direction of a process medium in a reaction space or a hydrogenation catalyst bed layer is from bottom to top.

The upflow type expanded bed reactor is a vertical upflow type reactor, and belongs to an expanded bed reactor when a catalyst is used; the vertical type means that the central axis of the reactor is vertical to the ground in a working state after installation; the upflow means that the material main body flows in the reaction process from bottom to top to pass through the reaction space or the catalyst bed layer or flow in the same direction with the upward catalyst; the expanded bed means that a catalyst bed layer is in an expanded state in a working state, the expansion ratio of the catalyst bed layer is defined as the ratio KBED of the maximum height CWH of the working state when a reaction material passes through the catalyst bed layer and the height CUH of an empty bed standing state of the catalyst bed layer, generally, when the KBED is lower than 1.10, the bed is called a micro-expanded bed, when the KBED is between 1.25 and 1.55, the bed is called an ebullated bed, and a suspended bed is considered as the most extreme form of the expanded bed.

The back-mixed flow expanded bed reactor refers to an operation mode of using a reaction zone or a main reaction zone of the expanded bed reactor, wherein liquid flow back-mixing or circulating liquid exists; the return flow or the circulating liquid refers to at least one part of liquid phase XK-L in the intermediate product XK or the final product XK at the flow point K as a circulating liquid flow XK-LR to return to an upstream reaction zone of the flow point K, and the reaction product of the circulating liquid flow XK-LR flows through the point K and exists in XK. The mode of forming the back flow can be any suitable mode, such as arranging a built-in inner circulation tube, a built-in outer circulation tube, a built-in liquid collecting cup, a flow guide tube, a circulating pump, an external circulating tube and the like.

The invention discloses a liquid product circulating upflow type expanded bed hydrogenation reactor system, which is characterized in that a liquid product returns to an upstream reaction space for circular processing or liquid product circulation exists in an operation mode of a reaction zone or a main reaction zone of an expanded bed reactor; the liquid product circulation in the hydrogenation reactor refers to that at least a part of the liquid phase XK-L in the intermediate product XK or the final product XK at the flow point K is used as a circulating liquid flow XK-LR to return to a reaction area upstream of the flow XK, and the circulating liquid flow XK-LR passes through the point K and exists in XK. The way of forming the circulation of the liquid product can be any suitable way, but a gas-liquid separation zone must be arranged in the head space in the reactor to obtain the circulating liquid and other products, namely a built-in liquid collecting cup, a diversion pipe and a circulating booster, wherein the circulating booster is usually a circulating pump and can be arranged inside or outside the reactor.

The liquid collecting cup or the liquid collector arranged in the reactor refers to a container which is arranged in the reactor and is used for collecting liquid, the upper part or the upper part of the container is usually provided with an opening on the side surface, and a guide pipe is arranged on the bottom part or the lower part of the container for conveying or discharging the collected liquid; the top liquid collector of the expanded bed reactor is usually arranged in a liquid removal area of gas-liquid materials to obtain liquid and gas-liquid mixed phase material flow containing a small amount of bubbles or obtain liquid and gas, and at least part of liquid phase products are pressurized by a circulating pump and then return to a reaction space for circular processing. Typical examples are the heavy OIL ebullated bed hydrogenation reactor, the HTI coal hydrogenation direct liquefaction reactor, the Shenhua coal hydrogenation direct liquefaction reactor used in the H-OIL process.

The suspended bed reactor of the invention can be in any suitable structural form, can be an empty cylinder suspended bed reactor to form piston flow or back mixing flow with internal circulation, can be an internal circulation guide cylinder to form internal circulation flow or internal external circulation flow, can be a back mixing flow pattern using an external circulation pipe to make liquid in an upper reaction space flow into external circulation flow of a lower reaction space former, and can be a back mixing flow pattern using a top product separation system, a liquid collection system and a guide system to form forced internal circulation flow through a circulation pressurization system.

The thermal high separator is a gas-liquid separation device for separating intermediate products or final products of hydrogenation reaction, and can be provided with a hydrogen gas stripping function to reduce the content of low-boiling-point components in separated liquid.

The two-stage or multi-stage hydrogenation method of the invention refers to a hydrogenation method comprising two reaction stages or a plurality of reaction stages.

The hydrogenation reaction stage refers to a flow path section from the beginning of a hydrogenation reaction process of a hydrocarbon raw material to the gas-liquid separation of a hydrogenation product of the hydrocarbon raw material to obtain at least one liquid-phase product consisting of at least one part of generated oil, and comprises the hydrogenation reaction process of the hydrogenation reaction stage and the gas-liquid separation process of at least one part of the hydrogenation reaction product of the hydrogenation reaction stage. Therefore, the first-stage hydrogenation method refers to a flow mode that the processing process of the initial hydrocarbon raw material only comprises one hydrogenation reaction step and a gas-liquid separation process of a product of the hydrogenation reaction step, wherein 1 or 2 or more hydrogenation reactors which are operated in series can be used according to the requirement of the hydrogenation reaction step, so that the number and the form of the reactors are not the basis for determining the reaction level, and the reaction step consisting of one or a plurality of series reactors and the product separator are combined together to form the hydrogenation reaction level in the sense of completion.

The secondary hydrogenation method of the invention refers to a flow mode that the processing process of the initial hydrocarbon raw material comprises a liquid material processing flow which is operated in series and is formed by two different hydrogenation reaction steps and a gas-liquid separation process of products of the corresponding hydrogenation reaction steps, wherein at least a part of a flow formed by the oil generated by the primary hydrogenation enters the secondary hydrogenation reaction process.

The three-stage hydrogenation method refers to a flow mode that the processing process of an initial hydrocarbon raw material comprises a liquid material processing flow which is operated in series and is formed by three different hydrogenation reaction steps and a gas-liquid separation process of products of the corresponding hydrogenation reaction steps, wherein at least one part of a material flow formed by the oil generated by the first-stage hydrogenation enters a second-stage hydrogenation reaction process, and at least one part of a material flow formed by the oil generated by the second-stage hydrogenation enters a third-stage hydrogenation reaction process. The flow structure of the hydrogenation method with more stages can be analogized according to the principle. The multistage hydrogenation method refers to a flow mode that the processing process of the initial hydrocarbon raw material comprises a liquid material processing flow which is operated in series and consists of three or more different hydrogenation reaction processes and hydrogenation product gas-liquid separation processes.

The three-stage hydrogenation method refers to a flow mode that the processing process of the initial hydrocarbon raw material comprises a liquid material processing flow which is operated in series and comprises three different hydrogenation reaction steps and a gas-liquid separation process of products of the corresponding hydrogenation reaction steps.

The invention relates to a method similar to a two-stage hydrogenation method, which is a method similar to the two-stage hydrogenation method, and is regarded as the two-stage hydrogenation method when the ratio of the flow of a back-mixing liquid phase of a rear-stage upper feeding back-mixing flow expansion bed reactor to the flow of a liquid phase in an upper feeding tends to be infinite.

The reaction product BASE-R10P of the upflow hydrogenation modification reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F is at least a gas-liquid two-phase material flow, and in most cases, the material flow belongs to a gas-liquid-solid three-phase material flow. The hydrogenation reaction effluent R10P is used for discharging a hydrogenation reaction product BASE-R10P, appears in the form of 1-path or 2-path or multi-path materials, and is a gas phase or liquid phase or gas-liquid mixed phase or gas-liquid-solid three-phase material flow.

The solvent hydrocarbon ADSC refers to hydrogen donor hydrocarbon or hydrogen donor hydrocarbon precursor, and serves as hydrogen donor hydrocarbon or hydrogen transfer hydrocarbon or viscosity reduction hydrocarbon or diluent in the upflow hydrogenation process (hydrogenation modification reaction process and hydrogenation thermal cracking reaction process) of heavy oil.

The hydrogen donor refers to hydrocarbon components with hydrogen donor function in the coal hydrogenation direct liquefaction reaction process, the heavy oil hydrogenation reaction process and the kerosene co-refining hydrogenation reaction process, and the hydrogen donor comprises partially saturated bicyclic aromatic hydrocarbon and partially saturated polycyclic aromatic hydrocarbon. The hydrogen supply hydrocarbon releases active hydrogen to stabilize the hydrogenation of thermal cracking free radicals, and reduces the concentration of the thermal cracking free radicals in the reaction space, thereby having the function of inhibiting thermal cracking and reducing the thermal cracking rate of heavy hydrocarbons, for example, in the front reaction section R10A of the heavy oil hydrocracking reaction process R10 where a large amount of thermal cracking reactions occur, the hydrogen supply hydrocarbon with sufficient amount has the function of inhibiting thermal condensation coking, and has positive influence on the production process; in the rear reaction section R10B in which the number of thermal cracking reactions is greatly reduced, the same amount of hydrogen-supplying hydrocarbons contains a part of excess hydrogen-supplying hydrocarbons, which have negative effects of inhibiting thermal cracking of heavy hydrocarbons and reducing the thermal cracking rate of heavy hydrocarbons.

The hydrogen donor precursor herein refers to a hydrocarbon component which can be converted into a hydrogen donor after hydrogenation or a converted product after hydrogen donor hydrocarbons lose part of hydrogen.

The hydrogen transfer hydrocarbon refers to hydrocarbon components with hydrogen transfer function in a coal hydrogenation direct liquefaction reaction process, a heavy oil hydrogenation reaction process and a kerosene co-refining hydrogenation reaction process, such as polycyclic aromatic hydrocarbon.

The parts of the present invention are described in detail below.

The following describes a hydrogen donating hydrocarbon (or hydrogen donating hydrocarbon component) DS, a hydrogen donating hydrocarbon precursor DS-BF, a hydrogen donating solvent SHS, a hydrogen losing and donating solvent (or hydrogen donating hydrocarbon precursor, or hydrogen donating hydrocarbon to be reactivated) MFS, a hydrogenation stabilization reaction process MR for conducting a reactivation process of the hydrogen losing and donating solvent MFS or the hydrogen donating hydrocarbon precursor DS-BF.

The hydrogen-supplying hydrocarbon component DS herein refers to a hydrocarbon component having a hydrogen-supplying function in a heavy oil thermal cracking reaction process (including a heavy oil hydrocracking reaction process), a coal hydrogenation direct liquefaction reaction process, and a kerosene co-refining hydrogenation reaction process, and the hydrogen-supplying hydrocarbon includes a partially saturated bicyclic aromatic hydrocarbon and a partially saturated polycyclic aromatic hydrocarbon (generally, a tricyclic hydrocarbon and a tetracyclic hydrocarbon are preferable). In the hydrogen supply hydrocarbon, the hydrogen supply speed of a dihydro body is higher than that of a tetrahydro body, and the hydrogen supply speed of the dihydro body of tricyclic aromatic hydrocarbon is higher or lower than that of the dihydro body of bicyclic aromatic hydrocarbon; tests have demonstrated that polycyclic aromatic hydrocarbons, although not having a hydrogen donating ability, have the ability to transfer hydrogen. The relative hydrogen supply rates at 400 ℃ for the following components were as follows:

for the hydrogen donor solvents SHS used in industry, which are usually mixed hydrocarbons containing the hydrogen donor hydrocarbon component DS or and the hydrogen donor hydrocarbon precursor hydrocarbon component DS-BF, common sources of hydrogen donor solvents SHS are:

① hydrocarbon fraction with the conventional boiling point of 220-480 ℃ of low-temperature coal tar;

② hydrocarbon fraction with the conventional boiling point of 220-480 ℃ of medium-temperature coal tar;

③ hydrocarbon fraction with the conventional boiling point of 220-480 ℃ of the high-temperature coal tar;

④ hydrocarbon fraction with the conventional boiling point of 220-480 ℃ of pulverized coal pyrolysis tar;

⑤ hydrocarbon fraction of ethylene tar at 220-480 ℃;

⑥ heavy oil is used as basic hydrocarbon fraction of 220-480 ℃ obtained in the heavy oil thermal processing process, wherein the thermal processing process is a heavy oil catalytic cracking process or a heavy oil catalytic cracking process;

⑦ hydrocarbon fractions with the temperature of 220-480 ℃ obtained in the process of direct liquefaction reaction by coal hydrogenation;

⑧ hydrocarbon fraction with normal boiling point of 450-570 ℃;

⑨ other hydrocarbons rich in the hydrogen-donating hydrocarbon component DS or mixed hydrocarbons with the hydrogen-donating hydrocarbon precursor hydrocarbon component DS-BF.

Taking the hydrocracking reaction process of heavy oil as an example, in the hydrocracking reaction process of hydrocarbons, the hydro-stabilization process of obtaining active hydrogen from hydrocarbon thermal cracking radicals is carried out, the hydrocarbon thermal cracking radicals belong to hydrogen-capturing agents, and meanwhile, the hydrocarbon components with excellent hydrogen-donating capability release active hydrogen atoms (called hydrogen loss) to become hydrocarbons with higher aromatic carbon rate and poorer hydrogen-donating capability; because the hydrogen supply hydrocarbon has special composition and higher price, for reducing the cost, for the occasion that a large amount of hydrogen supply hydrocarbon needs to exist, in order to reduce the consumption of the externally supplied hydrogen supply hydrocarbon, the DS-BF of the hydrogen loss and supply hydrocarbon (or a hydrogen supply hydrocarbon precursor or the hydrogen supply hydrocarbon to be reactivated) is generally required to be recovered in a certain way to obtain the MFS of the hydrogen loss and supply solvent, and the hydrogen supply capacity of the MFS of the hydrogen loss and supply solvent is recovered through the MR in the hydrogenation stable reaction process and then recycled; it is apparent that the hydrogen-losing hydrogen-donating solvent MFS is also a mixed hydrocarbon in general and is usually mixed with the product having the same boiling point in the heavy oil hydrogenation process, so that if the product having the same boiling point in the heavy oil hydrogenation process belongs to the hydrogen-donating hydrocarbon component DS or the hydrogen-donating hydrocarbon precursor component DS-BF, the amount of the hydrogen-donating solvent may be increased, and if the product having the same boiling point in the heavy oil hydrogenation process does not belong to the hydrogen-donating hydrocarbon component DS or the hydrogen-donating hydrocarbon precursor component DS-BF, the concentration of the hydrogen-donating hydrocarbon in the hydrogen-donating solvent may be decreased, and for a stable production system in which the hydrogen-donating solvent circulates, a recycled material in which the hydrocarbon component is substantially stable is formed.

Because the hydrogen supply solvent can rapidly provide active hydrogen and rapidly transfer the active hydrogen (for example, the active hydrogen on the surface of the catalyst is rapidly transferred so as to improve the efficiency of the catalyst for generating the active hydrogen and improve the utilization rate of the active hydrogen) in the hydrogenation and thermal cracking reaction process of the heavy oil, the utilization efficiency of the active hydrogen can be improved if the hydrogen supply hydrocarbon component DS can transfer more active hydrogen in a reasonable flow manner (for example, through more hydrocarbon hydrogenation reaction processes) in the circulation process of the hydrogen supply solvent, thereby forming the efficient use method of the active hydrogen.

The beneficial effect of the hydrogen donor hydrocarbon component DS in the hydro-thermal cracking reaction process of the hydrocarbons is mainly shown as follows:

① in the process of converting into hydrogen loss solvent, the molecular is uniformly dispersed in the whole reaction space, and provides active hydrogen for the free radical in the liquid phase reaction space, which has hydrogen supply ability, hydrogen supply agent and coking inhibitor function, and the distribution uniformity can not be realized by the present nanometer catalyst with the smallest granularity;

② the whole process of providing active hydrogen for hydrocarbon belongs to hydrogen transfer between hydrocarbon molecules, basically does not generate reaction heat, and has the function of reducing the reaction heat in the hydrogenation process of the target hydrocarbon oil;

③ can reduce the temperature of the hydrocarbon thermal cracking reaction, and has the function of a dynamic coking inhibitor;

④ has molecule inducing function, and can reduce the cleavage energy of molecular hydrogen and accelerate the dissociation speed of molecular hydrogen;

⑤ rapidly transferring active hydrogen (such as rapidly transferring active hydrogen out of the surface of the catalyst to improve the efficiency of the catalyst in generating active hydrogen and the utilization rate of active hydrogen);

⑥ under proper conditions and under the action of hydrogenation catalyst, it can convert the state of hydrogen-supplying hydrocarbon and its precursor for several times to act as active hydrogen transfer agent for several times.

① can induce thermal cracking reaction, reduce thermal cracking reaction temperature, and reduce thermal condensation reaction amount, thereby improving operation stability and prolonging operation period;

② can shorten the reaction time, reduce the amount of thermal condensation reaction, thereby improving the operation stability and prolonging the operation period;

③ can reduce the total temperature rise of the reaction;

④ can increase the retention rate of pyrolysis molecules, reduce the yield of thermal condensation compounds such as coke, and reduce the yield of gas, i.e. increase the yield of light oil products and save the energy consumption of solid-liquid separation;

⑤ can improve the operation stability, prolong the operation period and improve the catalyst efficiency;

⑥ can increase the overall thermal cracking conversion of heavy oil.

The hydrogenation reaction zone MR targeted for the production of hydrogen-donating hydrocarbons is described in detail below.

According to the invention, the stream SHS containing the hydrogen-donating hydrocarbon DS which is recycled is a stream or a separated stream of a hydrogen-donating hydrocarbon precursor stream SHSBF which is rich in bicyclic aromatic hydrocarbons and/or polycyclic aromatic hydrocarbons and is obtained by converting MRP (MRP) obtained as a hydrogenation reaction effluent in a hydrogenation reaction zone MR which aims at preparing the hydrogen-donating hydrocarbons; the hydrogenation reaction zone MR, which is targeted for the production of hydrogen-donating hydrocarbons, can be operated under any suitable conditions.

The hydrogenation stabilizing reaction process MR can adopt a granular catalyst bed layer (a down-flow fixed bed, an up-flow fixed bed and an up-flow micro-expansion bed) reaction mode, and the temperature is generally 280-440 ℃, the pressure is generally 6.0-20.0 MPa, and the volume space velocity of the hydrogenation catalyst MR-CAT is generally 0.05-10.0 hr^-1And the volume ratio of the hydrogen to the raw oil is 30: 1-3000: 1.

The hydrogenation stabilizing reaction process MR can adopt a moving bed or fluidized bed hydrogenation reaction mode using a particle catalyst, and the temperature is generally 280-440 ℃, the pressure is 6.0-20.0 MPa, and the volume space velocity of the hydrogenation catalyst MR-CAT is 0.05-10.0 hr^-1And the volume ratio of the hydrogen to the raw oil is 100: 1-1200: 1.

The hydrogenation stabilizing reaction process MR can even adopt a suspension bed hydrogenation reaction mode, and generally operates under the reaction conditions that the temperature is 280-440 ℃, the pressure is 6.0-20.0 MPa, the added hydrogenation catalyst is preferably an oil-soluble catalyst or a water-soluble catalyst with high dispersity, and the volume ratio of hydrogen to raw oil is 100: 1-1200: 1.

The aromatic hydrogenation partial saturation reaction in the hydrogenation reaction zone MR aimed at hydrogen supply hydrocarbon preparation of the present invention refers to a hydrogen-consuming reaction process in the presence of hydrogen and a suitable hydrogenation catalyst MR-CAT (catalyst having aromatic hydrogenation partial saturation function) for the occurrence of a hydrocarbon material SHSBF rich in bicyclic aromatic hydrocarbons and/or polycyclic aromatic hydrocarbons, wherein the minimum reaction depth has the minimum industrial significance: the hydrogenation reaction depth is determined according to the aromatic hydrocarbon component structure in the SHSBF and the expected aromatic hydrocarbon partial saturation degree, the higher the hydrogen supply hydrocarbon weight concentration value SHN in the hydrocarbon fraction with the conventional boiling point of 250-530 ℃ in the effluent MRP of the hydrogenation reaction is, the better the SHN is, the SHN is usually more than 6 wt%, and generally more than 10 wt%.

The hydrogenation reaction zone MR targeted for hydrogen supply hydrocarbon preparation has wide variation range of operation conditions due to different properties of raw materials (metal content, oxygen content, olefin content, sulfur content, nitrogen content, aromatic hydrocarbon content, distillation range and specific gravity) and different hydrogenation reaction (hydrodeoxygenation, hydrodesulfurization, hydrodenitrogenation and hydrogenation partial saturation) depths, and is determined according to specific process conditions.

For the reaction mode of the granular catalyst bed layer (downflow fixed bed, upflow micro-expansion bed), the hydrogenation reaction zone MR targeted for preparing the hydrogen-supplied hydrocarbon, the hydrogenation catalyst MR-CAT used can be one or the combination and the mixed loading of two or more kinds of hydrogenation refining catalysts, can be a special catalyst for specific raw materials, and can also be a hydrogenation refining catalyst which is used in the proper petroleum refining heavy diesel oil type or wax oil type hydrogenation refining process and has the functions of hydrogenation demetallization, hydrogenation deoxidation, hydrogenation desulfurization, hydrogenation denitrification, hydrogenation saturation and the like, and the combination thereof. The catalyst for the aromatic hydrocarbon hydrogenation partial saturation reaction process of producing the coal liquefaction solvent oil by using the coal liquefaction crude oil and the deep hydrofining catalyst of the coal tar light fraction can be generally used.

The hydrogenation reaction zone MR which aims at preparing hydrogen-supplying hydrocarbon uses a hydrogenation catalyst MR-CAT which at least comprises an aromatic hydrocarbon hydrogenation saturation catalyst, usually also comprises a hydrogenation demetalization catalyst and an olefin hydrogenation saturation catalyst (the position of the process is usually positioned before the bed layer of the aromatic hydrocarbon hydrogenation saturation catalyst).

Any make-up sulphur may be added to the hydrogenation reaction zone MR targeted for hydrogen-donating hydrocarbon production, as required, to ensure the minimum hydrogen sulphide concentration necessary in the reaction section, such as 500ppm (v) or 1000ppm (v), to ensure that the hydrogen sulphide partial pressure necessary for the catalyst does not fall below the minimum necessary value. The supplementary sulfur may be hydrogen sulfide or a material which can be converted into hydrogen sulfide and has no adverse effect on the hydroconversion process, such as hydrogen sulfide-containing gas or oil, or carbon disulfide or dimethyl disulfide which generates hydrogen sulfide after contacting with high-temperature hydrogen.

The hydrogen supply solvent is used in the upflow hydrogenation reaction process of the material containing the coal tar pitch, so that free radicals can be rapidly eliminated, the hydrogen content of a thermal cracking product can be improved, and the thermal cracking reaction can be inhibited, namely the thermal cracking conversion rate is reduced; while the enhanced residuum quality of the upflow hydroprocessing reaction process of the coal tar pitch-containing material allows for further hydropyrothermal cracking (such as cyclic hydropyrocracking) to increase the overall thermal cracking conversion. As for the overall effect of the primary thermal cracking of the material containing the coal tar pitch and the secondary thermal cracking of the primary thermal cracking tail oil of the material containing the coal tar pitch, the hydrogen supply solvent can be used for effectively improving the overall hydrogenation thermal cracking conversion rate and effectively reducing the yield of the externally thrown solid tail oil.

For the invention, the main purpose of the upflow hydrogenation process of the material containing the coal tar pitch is to perform thermal cracking desulfurization or hydrodesulfurization and hydrodemetallation to a proper depth, and simultaneously perform hydrodemetallation to a certain degree, hydrogenation aromatic hydrocarbon partial saturation reaction of high boiling point hydrocarbon components to a certain degree, hydrogenation thermal cracking reaction, thermal cracking reaction and hydrogenation stabilization reaction of thermal cracking free radicals to produce a suspension bed hydrogenation product with low sulfur content, low metal content and more needle coke suitable components, therefore, the suspension bed hydrogenation product may need to be fractionated first, then the obtained hydrogenated coal tar pitch is separated into primary hydrogenation refined pitch and primary hydrogenation heavy pitch, and then the primary hydrogenation heavy pitch is returned to the suspension bed hydrogenation modification process for secondary processing, thereby improving the beneficial overall thermal cracking conversion rate. As for the overall effects of the primary hydrocracking of the coal tar pitch and the secondary hydrocracking of the tail oil of the primary hydrocracking of the coal tar pitch, the hydrogen supply solvent can effectively improve the overall hydrocracking conversion rate, effectively reduce the yield of the externally thrown solid tail oil and further improve the economical efficiency of the process.

The direct coal liquefaction process, which includes the coal hydrogenation direct liquefaction process and other direct coal liquefaction processes, is described in detail below.

The direct coal liquefaction process of the invention refers to a method for directly obtaining hydrocarbon liquid by treating coal, and can be divided into the following processes according to the difference of solvent naphtha and catalyst, the difference of pyrolysis mode and hydrogenation mode and the difference of process conditions:

① pyrolysis liquefaction method, wherein the coal is pyrolyzed and extracted by heavy solvent to obtain low ash extract (called as bentonite), and the heavy oil is extracted by light solvent under supercritical condition to obtain heavy oil as main oil, the method does not use hydrogen, the yield of the former process is high, but the product is solid, and the extraction rate of the latter process such as supercritical extraction (SCE) is not too high;

② solvent hydrogenation extraction liquefaction method, such as solvent refining coal method 1 and II (SRC-1 and SRC-II), hydrogen supply solvent method EDS, Japan New energy development organization liquefaction method (NEDOL), etc., using hydrogen, but the pressure is not too high, the solvent naphtha has obvious effect;

③ high pressure catalytic hydrogenation, such as the German old and New liquefaction process (IG and NewLG) and the United states hydrogen Coal process (H-Coal) belong to this category;

④ Combined processing method of coal and residual oil (C0. processing), which comprises passing residual oil as solvent oil through a reactor together with coal in one step without circulating oil, and performing hydrocracking on the residual oil to obtain light oil;

⑤ underground liquefaction is carried out by injecting solvent into underground coal bed to depolymerize and dissolve coal, adding fluid impact force to disintegrate coal, suspending incompletely dissolved coal in solvent, pumping out solution, and separating;

⑥ dry distillation liquefaction method, which comprises pyrolyzing coal to obtain tar, hydrocracking tar, and upgrading.

The coal hydrogenation direct liquefaction process is described below.

The direct coal hydrogenation liquefaction process is a method for coal hydrogenation liquefaction in the presence of solvent oil, wherein the solvent oil can be hydrogen supply solvent oil with improved hydrogen supply capacity in a hydrogenation stabilization process or solvent oil without being modified in the hydrogenation stabilization process, and various processes such as the following processes are available according to the difference of the solvent oil and the catalyst and the difference of hydrogenation process conditions:

① solvent hydrogenation extraction liquefaction method, such as solvent refining coal method I and II (SRC-I and SRC-II), hydrogen supply solvent method EDS, Japan New energy development organization liquefaction method (NEDOL), etc., using hydrogen, but the pressure is not too high, the solvent naphtha has obvious effect;

② high pressure catalytic hydrogenation, such as the German old and New liquefaction process (IG and NewLG) and the United states hydrogen Coal process (H-Coal) belong to this category;

③ Combined processing method of coal and residual oil (C0. processing) comprises passing residual oil as solvent oil through a reactor together with coal in one step without circulating oil, hydrocracking residual oil to obtain light oil, and performing various processes in America, Canada, Germany and former Soviet Union;

④ direct liquefaction of Shenhua group coal;

⑤ patent CN 100547055C discloses a hot-melt catalysis method for preparing liquid fuel from lignite, which belongs to the process of medium-pressure hydrogenation direct liquefaction of lignite, and comprises two processes of a coal liquefaction reaction process and a coal liquefaction oil hydrogenation modification process.

In the direct coal hydrogenation liquefaction process, no matter what kind of direct coal hydrogenation liquefaction process, the objective is to obtain an oil product, the sought function is coal-to-oil, the necessary chemical change is coal hydrogenation, the common characteristic of the prior art is to use solvent oil and a catalyst, the conventional boiling range of the solvent oil is generally 200-530 ℃, most of the solvent oil is 200-450 ℃, the best solvent oil is 265-430 ℃, most of the solvent oil is distilled oil or hydrogenation modified oil thereof, and most of the contained aromatic hydrocarbon is aromatic hydrocarbon with 2-4 ring structures. Therefore, no matter what kind of coal hydrogen direct liquefaction process, the produced external oil discharge or coal liquefaction oil (usually coal liquefaction light oil) or coal liquefaction oil modified oil can be processed in the high aromatic hydrocarbon hydrogenation thermal cracking reaction process by using the method of the present invention as long as the composition of the oil has the characteristics of the raw material composition of the present invention.

The direct coal hydrogenation liquefaction process is a hydrogenation liquefaction reaction process in which coal and molecular hydrogen which may exist are used as raw materials, a specific oil product (usually, hydrogenation modified oil of coal liquefaction oil) is used as hydrogen supply solvent oil, and under certain operation conditions (such as operation temperature, operation pressure, solvent oil/coal weight ratio, hydrogen/solvent oil volume ratio and a proper hydrogenation catalyst), the coal directly undergoes carbon-carbon bond thermal cracking, free radical hydrogen addition stabilization and the like.

The direct coal hydrogenation liquefaction oil refers to an oil product produced in the coal hydrogenation liquefaction reaction process, exists in the effluent of the coal hydrogenation liquefaction reaction, and is a comprehensive reaction product based on hydrogen supply solvent oil, reaction consumed coal and reaction transferred hydrogen.

After the coal hydrogenation direct liquefaction reaction process is normally operated, the hydrogen-supplying solvent oil is generally hydrogenated modified oil of coal liquefied oil (usually distillate oil with a conventional boiling range higher than 165 ℃) produced in the coal hydrogenation liquefaction reaction process, and the main goal of the coal liquefied oil hydrogenation modification process is to produce the solvent oil for the coal hydrogenation direct liquefaction reaction process, specifically, to improve the content of components with good hydrogen supply function in oil products, such as naphthenic benzenes and dicycloalkylbenzenes, and the coal liquefied oil hydrogenation modification process is a hydrogenation process with moderate aromatic hydrocarbon saturation based on the fact that the coal liquefied oil contains a large amount of bicyclic aromatic hydrocarbons and a large amount of tricyclic aromatic hydrocarbons.

The final goal of the coal liquefaction reaction process is to produce oil products for external supply, and generally, the hydrogenated modified oil produced in the coal liquefied oil hydrogenation modification process is divided into two parts: one part is used as hydrogen supply solvent oil for the coal liquefaction reaction process, and the other part is used as external oil discharge in the coal liquefaction oil preparation process. Usually, at least a part of coal liquefaction light oil generated in the coal liquefaction reaction process is used as external oil discharge A in the coal oil preparation process, the rest of the coal liquefaction oil is used as raw oil in the coal liquefaction oil hydrogenation modification process to produce hydrogen supply solvent oil and external oil discharge B for the coal liquefaction reaction process, at this time, two paths of external oil discharge A and B exist, and the final outward oil discharge directions of the two paths of external oil discharge A and B are both generally used for producing high-quality oil products such as diesel oil fractions and naphtha fractions through a deep hydrogenation upgrading process.

In the direct coal hydrogenation liquefaction reaction process, a hydrogen supply solvent is essentially the most main foreground catalyst for the positive and negative reactions of coal liquefaction, rapidly provides most of active hydrogen in the coal liquefaction process, and directly determines the rapid hydrogenation stable speed of pyrolysis free radical fragments, so that the thermal condensation reaction is inhibited; in the direct coal hydrogenation liquefaction reaction process, solid catalysts such as pyrite, molybdenum sulfide and the like are more similar to a retarder of coal liquefaction negative reaction in nature, and solid catalyst particles adsorb colloid and asphaltene molecules MK with high viscosity, and the MK is contacted with active hydrogen on the surface of the solid catalyst, so that the thermal shrinkage of the MK is inhibited; in the direct coal hydrogenation liquefaction reaction process, solid catalysts such as pyrite, molybdenum sulfide and the like are essentially simultaneously used as a recovery catalyst of a hydrogen donor dehydrogenation product SH-Z, solid catalyst particles adsorb SH-Z and enable SH-Z to be in contact with active hydrogen on the surface of the solid catalyst, so that hydrogenation is recovered into hydrogen-supplying hydrocarbon with hydrogen supply capacity, and the recovery speed of the hydrogen donor dehydrogenation product SH-Z is directly determined; in the direct coal hydrogenation liquefaction reaction process, solid catalysts such as pyrite and the like are basically and simultaneously weak catalysts for target hydrocracking reactions such as the hydrocracking of asphaltene and preasphaltene. Therefore, in the direct coal hydrogenation liquefaction reaction process, solid catalysts such as pyrite and molybdenum sulfide are more like a catalyst operating in a background in a certain sense, and play a supporting and promoting role in coal liquefaction target product distillate oil. In the reaction process of preparing the oil by coal hydrogenation, the function of the hydrogen donor solvent DS is very important, so the operation condition and the effect of the hydrogenation stabilization reaction process of the solvent oil are naturally very important.

The coal hydrogenation direct liquefaction reaction process generally uses an upflow reactor, and the working mode can be selected as follows:

① suspension bed hydrogenation reactor;

② ebullated bed hydrogenation reactor, which discharges the catalyst with reduced activity from the bottom of the bed in a batch mode, and replenishes fresh catalyst from the upper part of the bed in a batch mode to maintain the catalyst inventory in the bed;

③ micro-expanded bed.

The coal hydrogenation direct liquefaction reaction process generally refers to a coal hydrogenation liquefaction method under the condition of solvent oil, wherein the solvent oil can be hydrogen supply solvent oil with improved hydrogen supply capacity in a hydrogenation stabilization process or solvent oil without being modified in the hydrogenation stabilization process, and various processes are available according to the difference of the solvent oil and the catalyst and the difference of hydrogenation process conditions.

The oil product obtained by directly liquefying coal through hydrogenation in the combined process comprises naphtha (the fraction with the conventional boiling range of 60-180 ℃), first light diesel oil (the fraction with the conventional boiling range of 180-220 ℃), second light diesel oil (the fraction with the conventional boiling range of 220-265 ℃), heavy diesel oil (the fraction with the conventional boiling range of 265-350 ℃), light wax oil (the fraction with the conventional boiling range of 350-480 ℃), heavy wax oil (the fraction with the conventional boiling range of 480-530 ℃), and liquefied residual oil (hydrocarbons with the conventional boiling point higher than 530 ℃).

According to the combined process, naphtha (the fraction with the conventional boiling range of 60-180 ℃) in the product obtained by directly liquefying coal through hydrogenation is a target product fraction, and can be subjected to deep hydrofining such as desulfurization and denitrification according to needs, and the benzene ring hydrogenation saturation reaction is expected to occur as little as possible.

The first light diesel oil (the fraction with the conventional boiling range of 180-220 ℃) in the product obtained by the coal hydrogenation direct liquefaction combined process is usually not suitable for entering the coal hydrogenation direct liquefaction reaction process, and the boiling point is too low and is easy to vaporize, so that the first light diesel oil is difficult to serve as a liquid phase solvent component; if the coal is subjected to the direct coal hydrogenation liquefaction reaction process, the products of the further thermal cracking reaction generate a large amount of gas and are not economical; therefore, unless the value of the gaseous hydrocarbon is huge, the first light diesel oil is generally not suitable for being processed in a coal hydrogenation direct liquefaction reaction process or a special hydrocracking process or a hydrocracking process or other thermal cracking processes, and can be generally subjected to a hydrofining reaction process for desulfurization and denitrification to produce clean light diesel oil.

In the combined process, the second light diesel oil (the fraction with the conventional boiling range of 220-265 ℃) in the product obtained by the direct coal hydrogenation liquefaction is a hydrogenation stable oil product which is hydrogen supply solvent oil with proper boiling point and excellent hydrogen supply capability and is required in the direct coal hydrogenation liquefaction reaction process, and in addition, in the direct coal hydrogenation liquefaction process, the second light diesel oil or the hydrogenation stable oil thereof plays a role of a liquid phase basic solvent component in the front reaction process of the direct coal hydrogenation liquefaction reaction process, but most of the second light diesel oil or the hydrogenation stable oil thereof is vaporized in the rear reaction process of the direct coal hydrogenation liquefaction reaction process, and the second light diesel oil or the hydrogenation stable oil thereof usually partially serves as the light hydrogen supply solvent oil to be used in the direct coal hydrogenation liquefaction reaction process, and part of the raw materials is used as hydrogenation upgrading raw materials for producing final products in the hydrogenation upgrading reaction process.

In the combined process, the heavy diesel oil (fraction with the conventional boiling range of 265-350 ℃) in the product obtained by the direct coal hydrogenation liquefaction is a hydrogenation-stabilized oil product which is the most needed hydrogen-supplying solvent oil with proper boiling point and excellent hydrogen-supplying capability in the direct coal hydrogenation liquefaction reaction process, and in addition, the heavy diesel oil or the hydrogenation-stabilized oil thereof plays a role of a liquid-phase basic solvent component in the whole flow of the direct coal hydrogenation liquefaction reaction process in the direct coal hydrogenation liquefaction reaction process, and usually, residual resources exist in the direct coal hydrogenation liquefaction reaction process, so the heavy diesel oil or the hydrogenation-stabilized oil thereof belongs to a main product in the direct coal hydrogenation liquefaction reaction process, therefore, part of the heavy diesel oil or the hydrogenation-stabilized oil thereof in the coal hydrogenation direct liquefaction reaction process is usually used as the heavy hydrogen-supplying solvent oil in the direct coal hydrogenation liquefaction reaction process, and part of the heavy diesel oil is used as a.

In the product obtained by the direct coal hydrogenation liquefaction combined process of the invention (the fraction with the conventional boiling range of 350-480 ℃), the hydrogenated stable oil product is the hydrogen-supplying solvent oil with proper boiling point and excellent hydrogen-supplying capability which is most needed in the direct coal hydrogenation liquefaction reaction process, and in addition, in the direct coal hydrogenation liquefaction process, the light wax oil or the hydrogenated stable oil thereof plays a role of a liquid-phase basic solvent component in the final high-temperature stage of the direct coal hydrogenation liquefaction reaction process, and is usually a scarce resource which is difficult to balance by itself in the direct coal hydrogenation liquefaction reaction process, so that the light wax oil or the hydrogenated stable oil thereof which is the coal liquefaction product is usually completely used as the heavy hydrogen-supplying solvent oil in the direct coal hydrogenation liquefaction reaction process, and simultaneously, the hydrogenation thermal cracking reaction which is needed in the light coal liquefaction process is carried out.

The heavy wax oil (fraction with the conventional boiling range of 480-530 ℃) in the product obtained by the combined process of the invention through direct coal hydrogenation liquefaction needs to be carried out under the liquid phase condition which is rich in hydrogen-supplying hydrocarbon and can provide a large amount of active hydrogen atoms so as not to be rapidly coked to maintain the long-period operation of the device, the dispersion of the coal liquefied heavy wax oil in the liquid phase in a reactor also needs to be carried out by means of the dispersion and dissolution of a large amount of hydrogen-supplying hydrocarbon, the thermal condensation compound or the coking compound of the coal liquefied heavy wax oil also needs to be dispersed and carried out of a reaction space by relying on liquefied semicoke as an aggregation carrier, therefore, the deep conversion of the coal liquefied heavy wax oil in the direct coal hydrogenation liquefaction reaction process is a reasonable inevitable choice, or the hydrogenation stable oil obtained by the coal liquefaction heavy wax oil through the hydrogenation stable reaction process enters the coal hydrogenation direct liquefaction reaction process for deep conversion, which is a reasonable inevitable choice; in addition, for the direct coal hydrogenation liquefaction process, the heavy wax oil or the hydrogenation stabilized oil thereof plays a role of a liquid phase basic solvent component at the last high-temperature stage of the direct coal hydrogenation liquefaction reaction process, and is usually a scarce resource which is difficult to balance in the direct coal hydrogenation liquefaction reaction process, so that the heavy wax oil or the hydrogenation stabilized oil thereof which is a coal liquefaction product is usually completely used as the heavy hydrogen supply solvent oil to be used in the direct coal hydrogenation liquefaction reaction process, and meanwhile, the hydrogenation thermal cracking reaction required in the light coal liquefaction process is carried out.

As the hydrocarbons with the conventional boiling point higher than 530 ℃ in the product obtained by the direct coal hydrogenation liquefaction of the combined process of the invention, namely the liquefied residual oil, exist in the coal liquefaction residue stream at the bottom of the vacuum tower, the liquefied residual oil is usually discharged from a system for reprocessing and is not recycled for processing, and of course, part of the liquefied residual oil can be recycled as required.

The kerosene co-refining integrated process is particularly suitable for optimizing the hydrogenation thermal cracking process of inferior heavy oil with high aromatic carbon rate, is beneficial to reducing the coke yield and prolonging the operation period, has high conversion rate of the inferior heavy oil and high yield of liquid products, and can mainly produce high aromatic latent naphtha and use aromatic hydrocarbon as downstream products.

The fraction section of the oil generated in the coal hydrogenation direct liquefaction reaction process used as the circulating cracked oil is preferably a fraction section with low yield of cracked gas, such as heavy diesel oil (fraction with the conventional boiling range of 265-350 ℃) and wax oil (fraction with the conventional boiling range of 350-530 ℃), and can be used for producing naphtha in high yield.

The heavy oils R10F according to the invention generally have a conventional boiling point for hydrocarbons of > 450 ℃, generally > 500 ℃, in particular > 530 ℃ and more particularly > 570 ℃.

The upflow type hydrocracking reaction process R10 of the heavy oil R10F refers to an upflow type expanded bed hydrocracking reaction process of the heavy oil, such as a suspension bed hydrocracking reaction process, an ebullated bed hydrocracking reaction process and the like.

In the upflow expanded bed hydrogenation thermal cracking reaction process R10 of the heavy oil R10F, at least part of thermal cracking reaction and thermal cracking free radical hydrogenation stable reaction of the heavy oil R10F are carried out, and at least part of hydrocarbon products with lower boiling points are generated; the heavy oil upflow hydrocracking reaction process R10 generally cannot achieve total light ends in a single pass reaction, and generally has a reasonably high single pass conversion rate of 70-85%, so that a certain amount of tail oil (hydrocracking product residue) THC-VR exists in the hydrocracking reaction product R10P, for example, 15-30%.

If, from the viewpoint of the component structure, the THC-VR as a hydrocracking product residue is itself a residue of heavy oil with large molecules not being lightened or a converted product or a concentrate of large molecules of a thermal condensate, the colloid content, asphaltene content and carbon residue content are significantly increased, for example, by 30% to 100%, as compared with the same boiling range fraction of the heavy oil R10F as a hydrocracking precursor raw material.

In order to increase the processing efficiency of the plant, it is desirable to increase the per pass conversion of heavy oil R10F; the increase of the conversion per pass of the heavy oil R10F inevitably reduces the amount of distillable distillate oil in the reaction product, and increases the amount of the thermal condensation product colloid, asphaltene and liquid-phase coke, and the increase of the amount of the colloid, the asphaltene and the liquid-phase coke and the decrease of the amount of the solvent oil of the colloid, the asphaltene and the liquid-phase coke develop to a super-saturation degree or a critical saturation degree, which can cause the colloid, the asphaltene and the liquid-phase coke to be separated out from a stable colloid solution system to become a super-saturated asphalt phase, namely a second liquid phase, cause the rapid operation in containers such as a reactor and the like, and force the shutdown of a coking device.

Under the premise of the fact that the method for avoiding the precipitation of the colloid, the asphaltene and the liquid-phase coke is inevitably reduced or the foreign solvent oil is introduced under the premise of the fact that the one-way conversion rate of the heavy oil R10F is increased, the one-way conversion rate of the heavy oil R10F is inevitably reduced or the one-way conversion rate of the heavy oil R10F is reduced, the processing efficiency of the device is reduced, therefore, the solvent oil (or hydrogen supply agent) for introducing the foreign proper colloid, the asphaltene and the liquid-phase coke and the using method thereof become an important technical problem, the economical efficiency of the upflow expanded bed hydrocracking reaction process R10 of the oil R10F is emphasized, and the core aim of the invention is to improve the existing solvent oil using method to reduce the using amount of the solvent oil (or the hydrogen supply agent) under the premise of admitting the foreign proper colloid, the asphaltene and the solvent oil (or the hydrogen supply agent) for introducing the liquid-phase coke, so as to reasonably reduce the scale of R10 and a product separation system, namely, obviously save investment, reduce hydrogen consumption, reduce the amount of KVGO condensed into residual oil and improve the economy of R10.

The conventional boiling range of main hydrocarbons of the aromatic hydrocarbon-rich wax oil KVGO is usually 370-570 ℃, generally 400-570 ℃ and particularly 450-550 ℃.

The conventional boiling range of main hydrocarbons of the aromatic hydrocarbon-rich medium-quality wax oil KVGO is usually 400-520 ℃, generally 425-510 ℃ and particularly 450-510 ℃.

The upflow hydrocracking reaction process R10 of the heavy oil R10F of the present invention is described in detail below.

The upflow hydrocracking reaction process of heavy oil R10F according to the present invention, the hydrocracking reaction of heavy oil R10 is described below.

In the upflow type hydrogenation thermal cracking reaction process R10 of the heavy oil R10F, at least part of thermal cracking reaction and thermal cracking free radical hydrogenation stable reaction of the heavy oil R10FL are carried out, and at least part of hydrocarbon products with lower boiling points are generated; generally, the heavy oil upflow hydrocracking reaction process R10 cannot achieve total light-ends in a single-pass reaction, that is, generally, a reasonably high thermal cracking depth is generally 70% to 85% of a single-pass conversion rate, so that a certain amount of tail oil, for example, 15% to 30% of tail oil, exists in the hydrocracking reaction product R10P, and in order to reduce the amount of discharged tail oil, a hydrocracking reaction process of the tail oil must be generally set, and in order to simplify the overall flow and reduce the investment and energy consumption, the hydrocracking reaction process of the tail oil and the upflow hydrocracking reaction process R10 of the heavy oil R10F are generally combined, that is, all or part of the tail oil and the upflow hydrocracking reaction process R10 are combined.

Although the upflow hydrocracking reaction process R10 of the heavy oil R10F targets thermal cracking reaction and thermal cracking radical hydrogenation stabilization reaction of macromolecular hydrocarbons, since the hydroprocessing catalyst generally used in the upflow hydrocracking reaction process R10 of the heavy oil R10F has a hydrofining function itself and active hydrogen is present to induce the hydrofining reaction of hydrocarbon molecules, some hydrofining reaction must occur in the upflow hydrocracking reaction process R10 of the heavy oil R10F.

In the upflow hydrocracking process R10 of heavy oil R10F, when the supply of active hydrogen is not timely, thermal cracking radicals of colloid and asphaltene undergo condensation reaction to produce molecules or structural groups with higher molecular weight, and the final result of repeated dehydrogenation and condensation reaction is the generation of coke, which is, of course, a negative reaction to be suppressed or reduced.

The main application object of the invention is an upflow hydrocracking reaction process R10 of heavy oil R10F, the number of used reactors can be 1 or 2 or more, and the number of commonly used reactors is 2-4; the reactor operation mode of the upflow type hydrocracking reaction process R10 of the heavy oil R10F can be any suitable mode, and is generally an upflow type expanded bed reactor or an upflow type expanded bed reactor with liquid product circulation, and the whole reaction zone of a single upflow type expanded bed reactor can be artificially divided into 2 or more reaction zones. The control mode of the inlet temperature of any reaction zone of the upflow type expanded bed reactor can be adjusting the temperature or the flow rate of hydrogen, adjusting the temperature or the flow rate of oil products, and certainly, the introduction of a heat exchanger can also be used for heat exchange and cooling.

The upflow hydrocracking reaction process R10 of heavy oil R10F, using a reactor whose volume ratio of liquid phase to gas phase (or vapor phase) in the reaction space may be the case of liquid phase being dominant, defines "actual volume of liquid phase/(actual volume of liquid phase + actual volume of gas phase)" in the reaction space as the liquid phase fraction KL of the reaction space, which is usually greater than 0.45, typically greater than 0.55, and even greater than 0.70, forming a virtual intensified liquid phase hydrogenation mode, may require 2 or more additions of hydrogen gas at different heights of the reactor in order to keep the hydrogen partial pressure of the reaction space sufficiently high.

When the latter half reaction process R10B of the upflow hydrocracking reaction process R10 of the heavy oil R10F is combined with the processing of the heavy oil component crpr in the reaction product CRP of the upflow hydro-modification reaction process CR of the inferior hydrocarbon CRF, the residence time of the latter half reaction process R10B usually satisfies the requirement of controlling the hydrocracking rate of the heavy oil component crpr, and the upper limit of the hydrocracking rate of the heavy oil component crpr is usually set to prevent the single-pass conversion rate from being excessively high.

The reactor form of the heavy oil upflow type hydrogenation thermal cracking reaction process R10 can be any suitable form, and has various known forms, such as an upflow fixed bed reactor, an upflow micro-expansion bed reactor, an upflow moving bed reactor, an upflow online replacement bed reactor, a boiling bed reactor, a suspension bed reactor, a combined bed reactor of the boiling bed and the suspension bed and the combination of the specific forms thereof, and most of the reactors have industrial application cases, and form fixed technical characteristics.

The colloidal asphalt-like component contained in the petroleum-based residual oil is a dispersed phase generally existing in a supermolecular structure, analysis data shows that the colloidal asphalt-like component dispersed phase is a stable structure group with the molecular weight as high as thousands to tens of thousands or even hundreds of thousands, and the group contains a large number of polycyclic aromatic hydrocarbon units and contains elements such as metal, sulfur, nitrogen and the like, the main task of the lightening process is to dissociate, hydrogenate and saturate the macromolecules into small molecules which are ten times, hundred times or even thousands times less than the original carbon, obviously, the thermal cracking task of the process is dominant, which cannot be realized by only depending on hydrodemetallization, hydrodesulfurization, hydrodenitrogenation and hydroaromatic saturation, and the pre-hydrogenation processes such as hydrodemetallization, hydrodesulfurization, hydrodenitrogenation, hydroaromatic saturation and the like of the petroleum-based residual oil are substantially the hydrofining process prepared for the subsequent hydrocracking reaction, otherwise, the hydrocracking catalyst active center in the conventional downflow fixed bed reactor is quickly covered due to metal deposition and quick coking, so that the operation period is too short, and the lowest economic operation period required by the industrial process cannot be maintained; even if the conventional processes of hydrodemetallization, hydrodesulfurization, hydrodenitrogenation, hydroaromatic saturation and the like in the downflow fixed bed reactor are used, the higher conversion rate cannot be achieved, because the problem of rapid and large-amount coking inevitably generated in the high-temperature thermal cracking process is difficult to overcome, which is determined by the thermodynamic property of the process; in order to overcome the defects of the fixed bed reactor system, the reactor forms are various forms such as an upflow fixed bed, an upflow micro-expansion bed, an upflow on-line replacement bed, an upflow strong expansion bed, namely a boiling bed, an upflow limit expansion bed, namely a suspension bed boiling and a downflow on-line replacement bed.

In the hydrocracking process of petroleum-based residual oil, the conversion rate of cracking the fraction with the conventional boiling point of more than 530 ℃ into the fraction with the conventional boiling point of less than 530 ℃ is usually 40-80 percent or even higher, in order to achieve such high cracking rate and improve the reaction speed, the high-temperature condition necessary for thermal cracking with stronger degree is inevitably used, the rapid coking of the catalyst active center is inevitable, in order to remove and replace the catalyst with the rapidly reduced activity caused by metal deposition and coking in the reactor bed layer, the technical personnel develop an up-flow type expansion bed reactor with the larger expansion ratio of the boiling bed layer and the suspension bed layer, and combine the subsequent hot high-pressure separator and the reactor into a combined device, greatly simplify the transfer system of the residual oil with high viscosity, easy foaming and easy solidification between the devices (between the reactor and the hot high fraction), improve the reliability of the system, and reduce the production cost of the petroleum-based residual oil, The safety and the heat insulation performance improve the uniformity of the temperature of the materials in the reactor and save the occupied area; the method has the advantages that the high conversion rate of residue oil boiling bed hydrocracking and suspension bed hydrocracking can not be achieved by a fixed bed reactor, endothermic cracking reaction and exothermic hydrogenation reaction are mixed to be beneficial to the utilization of reaction heat and the reduction of reaction temperature rise, and a large amount of thermal state reaction generated oil or intermediate reaction generated oil is recycled to directly heat raw oil so as to reduce the preheating temperature of the raw oil; in the unfavorable aspect, the expansion ratio of the catalyst bed layer is larger, compared with a fixed bed reactor, the complexity of the system is increased, the stability of the operation is reduced, and the engineering investment is greatly increased; the loss of the catalyst due to reasons other than coking is increased due to increased abrasion and collision of the catalyst; the quality of the product containing a portion of the fresh feed low conversion product is necessarily poor because of the severe back mixing of the catalyst and liquid phases present in the bed.

The residue OIL boiling bed hydrocracking industrialization technology comprises an H-OIL technology and an LC-FINING technology, in order to optimize and stably control the boiling state of a catalyst, a circulating OIL circulating pump system is arranged, a collector of circulating OIL is arranged above a catalyst bed layer in a reactor, namely, a high-temperature high-pressure separator which provides circulating OIL for a circulating pump and needs to be arranged at a high-elevation position is combined with a boiling bed reactor, the structure of the high-temperature high-pressure separator is simplified, but in order not to influence the fluidization state of the boiling bed, the arrangement position, the size and the form of the collector of the circulating OIL need to be carefully designed; usually, a collector of circulating oil is arranged right below a spherical seal head at the upper part of a reactor, a collector liquid guide pipe of the circulating oil is arranged in the reactor, and the liquid guide pipe has a certain rectification effect on the gas, liquid and solid multi-phase flow of a suspension bed or a boiling bed layer, so that the heat preservation and heat tracing problem of the liquid guide pipe is solved, and the adverse effect of the fluid flow in a flow guide pipe on the equipment stability of the reactor is weakened or eliminated; a hydrogenation reaction system of residue oil boiling bed, a catalyst intermittent discharge system and a catalyst intermittent feeding system are required to be arranged, and the system is complex, large in investment and complex in operation; another disadvantage of the residue ebullated-bed hydrogenation system is that part of the product is highly hydrogenated and saturated hydrocarbons, so the liquid phase of the product has poor ability to dissolve residual colloids and asphaltenes, and therefore, the conversion rate is low and the yield of tail oil is high, which limits the economy of the process to a certain extent; another disadvantage of the residue ebullated-bed hydrogenation system is that it is not possible to process inferior heavy residues with too high a content of carbon residue and too high a content of metals, because too high a content of metals makes the consumption of demetallization catalysts too large and makes the catalyst cost too large, and too high a content of carbon residue makes the conversion rate of the reaction process too low or rapid coking causes rapid shutdown, which limits the scope of application of the process.

The upflow fluidized bed hydrogenation technology has the technical key points that a catalyst bed layer is violently expanded by the upward flow of reaction materials (mainly liquid phase), the expansion rate of the catalyst bed layer is generally between 25 and 45 percent, and the catalyst bed layer can form the capability of damaging catalyst agglomeration and a wide area channel for freely discharging small particle impurities at the cost of losing the advantages of high activity, high interception rate and uniform material hydrogenation conversion depth of part of fixed bed hydrogenation catalysts, so that heavy oil with higher metal content and higher residual carbon content can be processed, the product quality of the heavy oil is reduced too much compared with the fixed bed technology, but the product quality of the heavy oil is better than that of a suspended bed; because the expansion power of the fluidized bed is mainly derived from carrying of liquid phase materials, a large amount of hydrogen is not suitable to be used in the process so as to prevent the volume efficiency of the liquid phase of the reactor from being too low, therefore, the exothermic effect in the reaction process cannot be too high, the fluidized bed hydrogenation technology is more suitable for processing paraffin-based or paraffin intermediate-base petroleum-based heavy oil, the macroscopic heat effect after the heat absorption amount of the thermal cracking reaction and the exothermic amount of free radical hydrogenation are offset is smaller, the total temperature rise of the reactor is lower, and the hydrogen consumption of the raw oil in unit weight is usually 1.4% -2.3%. However, even so, the deactivation rate of the catalyst is still too high, for which reason the average activity of the catalyst is maintained by periodically withdrawing part of the old catalyst with low activity and then supplementing part of the new catalyst with high activity, thus resulting in the high cost of consumption of the highly active hydrogenation catalyst, which is expensive, and in fact, it is not economical to process petroleum-based low-quality residues. Meanwhile, due to the characteristics of thermal reaction, the quality of hydrogenation tail oil is poor at high conversion rate, and only the hydrogenation tail oil can be used as fuel oil to vaporize the raw material, so that the conversion rate of the light weight of the raw material hydrogenated by the fluidized bed is usually 60-75 percent, namely the conversion rate is low. The granular catalyst used in the boiling bed hydrogenation technology is basically the same as the conventional fixed bed granular (preferably spherical) hydrofining catalyst, still belongs to a high-activity granular catalyst rich in a large number of internal pore channels and high internal surface area, and cannot meet the requirements of diffusion and hydro-conversion of low-quality residual oil macromolecules, the conventional boiling point is higher than 530 ℃ and has a huge molecular size and strong polarity, or the pore channels of the catalyst are blocked to lose activity, or the catalyst is adsorbed on the active center of the inner wall for a long time to generate a shielding effect, and under the condition of lacking active hydrogen, because the hydrogenation solid is difficult to desorb and desorb, a thermal condensation dominant reaction is generated, and the pore channels are blocked. The excessive catalyst deactivation speed results in unacceptable hydrogenation catalyst consumption cost, and more importantly, the great amount of reaction heat released by the great amount of saturated aromatic hydrogen consumption makes the boiling bed hydrogenation technology have no safety, the high temperature induced fast coking of colloid asphaltene also forms great amount of coking in the bottom distribution disc and central liquid circulation pipe of the reactor, and the equipment is forced to stop fast. If the upflow boiling bed hydrogenation technology is selected to process the inferior residual oil with high metal content and high carbon residue content, the results are necessarily that a large amount of coke is generated in the reactor, the operation period is too short, the reaction temperature cannot be controlled, namely unsafe, and the catalyst deactivation cost is surprisingly high, and the effects are proved by the industrial operation results of the trial-produced inferior heavy oil boiling bed hydrogenation device.

The development of residual oil suspension bed hydrocracking technology is based on the coal hydrogenation direct liquefaction technology of 20 th century 40 s, and is a process of residual oil thermal cracking reaction and thermal cracking free radical hydrogenation stable reaction which are caused under high temperature and high pressure by leading reaction under the condition of coexistence of hydrogen and fully dispersed catalyst or additive. In the hydrocracking reaction process of the suspension bed, the dispersed catalyst or additive is fine-particle powder which is suspended in reactants and can effectively inhibit the generation of coke. The residual oil suspension bed hydrogenation technology has almost no limit to the content of mechanical impurities of the raw materials, and can process asphalt and oil sand.

Typical residual oil suspension bed hydrocracking technologies with industrial operation performance include CANMET residual oil suspension bed hydrocracking process in Canada and EST residual oil suspension bed hydrocracking process in Eini, Italy. Other residual oil suspension bed hydrocracking technologies include BPVCC technology route from British oil company, BPVCC technology from British oil company, HDHPLUS technology from Venezuela national oil company (PDVSA), Uniflex technology from UOP in the United states, VRSH technology from Chevron in the United states, and the like.

In order to overcome the defects of the particle catalyst hydrogenation technology, the suspension bed hydrogenation technology thoroughly abandons the mode of using a huge amount of inner surfaces of particle catalysts as hydrogenation reaction sites, and the technical key points are that the outer surfaces of high-dispersity particle catalysts are used as the hydrogenation reaction sites, so that the problem of a diffusion path for colloid asphaltene to reach the hydrogenation reaction sites is thoroughly solved, the colloid asphaltene can be used for treating inferior heavy oil with higher metal content and higher carbon residue content, and certainly, the inferior heavy oil with extremely high metal content and extremely high carbon residue content is preferably treated by a coking process such as a delayed coking process; the bed expansion rate of the reaction space of the suspension bed hydrogenation reactor reaches the maximum value, and the addition amount of the solid catalyst is usually lower than 10 percent (based on the weight of the raw oil), thereby forming the advantages of 'having coke carrier capacity' and 'discharging free channel of suspended particle impurities'. However, in fact, the suspension bed hydrogenation reactor does not have the bed concept, the reaction space completely loses the advantages of high activity, high interception rate and uniform material hydrogenation conversion depth of the fixed bed hydrogenation catalyst, and the fixed bed hydrogenation catalyst has the dual characteristics of high liquid phase back mixing and high liquid phase short circuit, so that the product quality is greatly reduced compared with the fixed bed technology, and the suspension bed hydrogenation technology can only be used as a pretreatment process of poor oil, but cannot produce high-quality products.

The reaction efficiency of the catalyst surface of the suspension bed hydrogenation reactor strongly depends on the renewal frequency of the catalyst surface and the stable replacement rate of the reaction space, so the renewal means and the replacement means of the catalyst surface are important technical means which can not be lost and can improve the catalyst efficiency, and the existing reactor of the industrial heavy oil suspension bed hydrogenation device adopts a bubbling bed without a circulating pump, which is a huge technical defect, and the result is that: the internal back-mixing liquid phase quantity is uncontrollable, the internal back-mixing catalyst quantity (catalyst deposition quantity) is uncontrollable, the suitable particle size range of the catalyst is too narrow to be controlled, the liquid phase retention time is uncontrollable, the uncontrollable performance is stronger along with the enlargement of the diameter of the reactor, and the effects are proved by the industrial operation result of the trial production poor-quality heavy oil suspended bed hydrogenation device. The present invention recommends the use of a suspended bed reactor with liquid product circulation in order to achieve the desired renewal frequency of the catalyst surface and a stable rate of replacement of the reaction space.

The reaction efficiency of the catalyst surface of the suspension bed hydrogenation reactor is also influenced by the adsorption and occupation of polar impurities in the gas phase in the reactor, and a large amount of polar impurities such as H are generated in the coal tar hydrogenation process and the tar and coal co-refining process₂O、NH₃、CO、CO₂The catalyst can be strongly adsorbed on the surface of the catalyst to form a shielding effect, so that the CHEVRON company of the international well-known oil product technology supplier provides a scheme for arranging a gas-liquid separator in the middle of a reactor to discharge impurity gas in time and introduces high-purity hydrogen at the lower part of a subsequent suspension bed hydrogenation reactor, but the arrangement of the independent gas-liquid separator has large investment, difficult liquid level control and large operation risk; therefore, the project recommends that a 'gas short-flow' technology can be adopted, gas-liquid mixed phase materials containing gas are introduced into the space at the top of the suspension bed reactor for gas-liquid separation under the condition of not adding a gas-liquid separator, and gas phase is directly dischargedMost of the liquid phase enters a liquid phase reaction space through a circulating pipe, and a high-purity hydrogen material flow is introduced into the lower part of a subsequent suspension bed hydrogenation reactor to form a gas phase environment with extremely low impurities and a condition with high hydrogen volume concentration, so that a condition is created for fully exerting the activity of the catalyst, the total pressure of the device is reduced, the one-way conversion rate is improved, the thermal cracking gas-making reaction is reduced, and the thermal condensation reaction is reduced; the scheme for timely discharging the impurity gas also has the advantages of timely discharging the low-boiling-point hydrocarbon components and reducing the thermal cracking rate, and is favorable for improving the liquid yield and reducing the hydrogen consumption.

A typical heavy oil lightening reaction which occurs inside a suspension bed hydrogenation reactor for poor-quality heavy oil is essentially a series process of performing double bond hydrogenation of liquid-phase macromolecules into single bonds, cracking of the single bonds into free radicals and stable free radical hydrogenation in a liquid phase, a large number of free radicals are generated in the whole aggregation-state liquid phase at a high thermal cracking temperature (400-480 ℃) and are relatively uniformly distributed in the whole liquid phase space, the free radical hydrogenation is stabilized at the fastest speed for preventing thermal condensation, obviously, the purpose cannot be achieved by virtue of active hydrogen on the surface of a catalyst (because the probability of liquid-phase hydrocarbon molecules contacted by the catalyst is too low, the moving process of the active hydrogen can also be combined into inactive hydrogen molecules), preferably, the active hydrogen and the free radicals uniformly exist adjacently, and are synchronously released when the free radicals are generated, so as to realize high-efficiency active hydrogen supply. The timely addition of the hydrogen donor with proper boiling point can just over-meet the requirement, prevent thermal condensation and improve the retention rate of light products, and the effects are proved by the successful long-term operation results of the Shenhua coal hydrogenation direct liquefaction device which runs for 8 years and uses the hydrogen donor. For the heavy fraction with huge molecular size and strong polarity, which has the conventional boiling point higher than 530 ℃, if active hydrogen can not be provided timely, a large amount of thermal cracking free radicals of colloid and asphaltene can condense condensates larger than the cracking precursors thereof, so that the yield of hydrogenated thermal cracking distillate oil (hydrocarbons with the conventional boiling point lower than 530 ℃) is reduced, and even thermal condensates such as coke or coke precursors which are dissolved and carried by the liquid phase in the reaction process are generated to cause rapid shutdown of the device, and the effects are proved by a large number of experimental results. The invention uses the operation mode of sufficient hydrogen donor, aims to provide the raw material residual oil with more rigorous thermal cracking conversion rate or processing property by timely providing sufficient active hydrogen to inhibit coking, enlarges the application range of the process and improves the operation stability and the economical efficiency of the process.

Possible uses of the hot high pressure separation process or the warm high pressure separation process of the present invention are described in detail below with respect to the XHBM process.

In the gas stripping process XHBM, the countercurrent contact separation times of the liquid hydrocarbon W material and the stripping hydrogen XBH are as follows: generally 1 to 8 times, usually 2 to 4 times; the quantity of the stripping hydrogen XBH is determined according to the requirement of the separation target of the XHBM component in the stripping process; the operating pressure of the XHBM of the stripping process, typically slightly below that of its feed; the operation temperature of the gas stripping process XHBM is determined according to the requirement of the gas stripping process XHBM component separation target, and is usually 180-480 ℃, and is usually 250-440 ℃.

The working mode of the upflow reactor can be selected as follows:

① suspension bed hydrogenation reactor;

③ combined hydrogenation reactor of suspension bed and boiling bed;

④ micro-expanded bed.

The invention relates to a hydrocarbon catalytic thermal cracking reaction process, which refers to a thermal cracking reaction process taking carbon-carbon bond breakage as a main purpose under the condition of a thermal cracking catalyst, and comprises a conventional catalytic cracking reaction process or a conventional catalytic cracking reaction process, wherein the main raw material oil which is usually processed is wax oil and atmospheric residue, a document T001 recording the technology is ① publication name, catalytic cracking process and engineering, ② retrieval is encoded by books, ISBN encoding is 7-80043 537-7, Chinese edition library CIP data core word (2004) No. 131193, ③ main encoding is Chenjun Jun, ④ publication, China petrochemical press, documents T001 pages 459 to 488 of catalytic cracking process and engineering, record physical property data of typical catalytic cracking light diesel oil (light cycle oil), catalytic cracking return oil refining (heavy cycle oil) and catalytic cracking clarified oil, according to different operation conditions and product separation schemes of specific devices, the catalytic cracking light diesel oil (light cycle oil), catalytic return oil (heavy cycle oil) and catalytic cracking return oil (heavy cycle oil) can be processed in a certain range of hydrogenation reaction and the thermal cracking reaction yield can be improved based on the catalytic cracking reaction process.

The heavy oil coking reaction process refers to a thermal processing process of carrying out deep thermal cracking and condensation reaction on heavy oil (such as vacuum residuum, cracked residue oil and the like) which is poor in hydrogen as a raw material under the conditions of high temperature and long reaction time, wherein the raw material is converted into gas, naphtha, gasoline, diesel oil, heavy distillate oil (coked light wax oil, coked heavy wax oil) and coke, the process types of the coking process comprise kettle coking, open hearth coking, delayed coking, contact coking, fluid coking and flexible coking, the modern heavy oil coking process comprises processes of delayed coking, contact coking, fluid coking, flexible coking and the like, T002, which records the technology, is the name of ① publication, namely delayed coking process and engineering, ②, retrieval book code of ISBN code 978-7-80229-456-1, Chinese edition library CIP data keyword (2007) No. 168082, ③, main code of Dianthus superbus, ④ publication, T002, page 188 of the publication of China press and the delayed coking process, and page 188 of the engineering, the delayed coking process, the modified wax oil can be subjected to a specific coking reaction process based on the modified wax oil, the characteristic of the coking process, the yield of the coking process is improved, and the coking of the coking process is improved in a specific coking process, and the coking process, the coking process is improved in a specific coking process, and the coking process of the coking.

The quality of the catalytic cracking cycle oil (heavy cycle oil), catalytic cracking clarified oil and coking wax oil is poor, so that direct catalytic cracking or hydrocracking is difficult, and the optimized processing is realized by adopting the invention to carry out the hydrocracking.

The flow form or type of the process for the hydro-thermal cracking of petroleum heavy oil such as residual oil is described in detail below.

The reaction separation section of the present invention refers to a process comprising a raw heavy oil hydrocracking reaction process (or referred to as a reaction section) and a separation process (or referred to as a separation section) of heavy oil hydrocarbon components and lower boiling point hydrocarbon components in a hydrocracking reaction product; the process for separating the heavy oil hydrocarbon component from the lower boiling point hydrocarbon component may be a process for separating the residual oil from the wax oil component (usually including a vacuum fractionation process), a process for separating the heavy wax oil component from the light wax oil component (usually including a vacuum fractionation process), or a process for separating the diesel oil from the wax oil component (which may or may not include a vacuum fractionation process).

The reaction separation process of the present invention comprises a first hydrocracking reaction process of raw heavy oil and a first separation process of heavy oil hydrocarbon components and lower boiling point hydrocarbon components of a first hydrocracking reaction product, and the process can comprise a recycling process of recycling unconverted residual oil or modified oil thereof discharged from the first separation process (generally comprising a vacuum fractionation process) back to the first hydrocracking reaction process for recycling hydrocracking.

The existing suspension bed hydrocracking reaction separation methods of heavy oil or residual oil belong to a reaction separation process, wherein the residual oil suspension bed hydrocracking reaction separation method with industrial operation performance comprises a Canadian CANMET residual oil suspension bed hydrocracking process (which is later integrated into the Uniflex technology of UOP company in the United states) and an EST residual oil suspension bed hydrocracking process of Italy Eny company. Other residual oil suspension bed hydrocracking reaction separation methods include BPVCC technology of British oil company, HDHPLUS technology of Venezuela national oil company (PDVSA), VRSH technology of Chevron in the United states and the like.

If a thermal cracking system of the circulating heavy oil containing solid particles and unconverted residual oil components is arranged, in a first reaction section, in order to prevent the circulating residual oil which is carried by raw oil and is difficult to thermally crack, thermal condensation coke or coke precursors from being excessively accumulated to form high-concentration asphaltene to deteriorate the liquid phase property of the first reaction section (increase the carbon residue value, increase the viscosity value and reduce the average hydrogen content), in order to prevent sulfide solids, other ashes, catalyst solid particles and other solids generated by metals carried by the raw oil from being excessively accumulated to form high-concentration solid-containing circulating residual oil, a certain ratio of vacuum residue discharged by the first separation section is required to be discharged; compared with the fresh residual oil UR10F, the discharged vacuum residual oil of the first separation section has higher solid particle carrying rate, higher asphaltene concentration and more difficult hydrocracking cracking of the asphaltene.

The residual oil suspension bed hydrogenation thermal cracking reaction separation method only provided with one reaction separation process and a solid particle-containing unconverted residual oil circulating thermal cracking system comprises an EST residual oil suspension bed hydrogenation thermal cracking process of Italy Eny company.

In the EST residual oil suspension bed hydrogenation thermal cracking process, the vacuum residual oil is subjected to hydrocracking in a suspension bed reactor under the existence of a molybdenum-based catalyst and mild operating conditions (the temperature is about 400-425 ℃, and the pressure is about 15-17 MPa), and the vacuum residual oil is converted into a light product. The converted oil enters a separation system to recover gas, naphtha, middle distillate oil and wax oil, a hydrogen-containing gas phase product enters an amine washing part after a light product is separated, clean hydrogen-containing gas is recycled to the reaction process after being compressed again and supplemented with hydrogen, and the distillate oil is recovered from a liquid phase. In the EST residual oil suspension bed hydrogenation thermal cracking process, an oil-soluble matrix is used for being converted into non-carrier MoS in a form of nano-scale thin layer in a reactor₂. Because the sulfide generated in the reaction process of the metal carried in the raw material residual oil is deposited to form a separate phase, the bare MoS is not interfered₂The active center, so the catalyst is not changed in the whole operation process, and the catalyst is not aged and can be recycled for a plurality of times. Because the catalyst exists in a nano-scale thin layer form, the catalyst has extremely large external surface area and extremely high dispersity, and the efficiency of activating hydrogen and inducing the side chain of an aromatic ring to break is very high.

In EST residual oilsIn the suspension bed hydrogenation thermal cracking process, for the catalyst, the influence of coke is small, the surface area is large, and mass transfer diffusion resistance is not existed, so that the EST catalyst has higher activity than the supported catalyst, and the very high specific activity (activity per unit mass) makes the concentration of the EST catalyst only need to be maintained at several thousand mug/g level, so that it can form good catalytic reaction condition₂The distance between the thin layers is several orders of magnitude smaller than the distance between the supported catalyst and the oil molecules, thus shortening the free radical generation time and the time required for the free radicals to reach the catalyst surface and complete the hydrogenation stabilization process, and reducing coking. MoS₂The catalyst has the capability of catalyzing hydrogen to convert the hydrogen into active hydrogen (hydrogen atoms) and activating aromatic rings, so that the hydrogenation of aromatic hydrocarbon and the reduction of carbon residue are realized, and the removal of heteroatom through the reactions of hydrodesulfurization, hydrodenitrogenation, hydrodeoxygenation, hydrodemetallization and the like can be realized through the hydrogenolysis of C-heteroatom bonds.

In the EST residual oil suspension bed hydrogenation thermal cracking process, unconverted residual oil is recycled to a reactor together with dispersed catalyst and other solids. Optimizing process severity (reaction time and reaction temperature) according to the quality of raw material residual oil, mixing part of unconverted residual oil (the quality of which is poorer than that of the fresh raw material residual oil in the same boiling range, lower hydrogen content and higher carbon residue value) with the fresh residual oil for circulating hydrogenation thermal cracking reaction, on one hand, enabling the asphaltene capable of being lightened to realize deep conversion lightening through multiple reactions, on the other hand, enabling the asphaltene difficult to lighten to realize multiple thermal reactions to form heat condensation product accumulation, under the condition of maintaining a certain discharged residual oil ratio, discharging part of asphaltenes or coke precursors which are difficult to convert in time, the properties of the circulating residual oil are in a stable state for a long time, the phenomena of coke formation and equipment scaling caused by asphaltene precipitation of the circulating residual oil and the liquid phase of the first reaction section are avoided, cyclic hydrocracker cracking is achieved as such until near complete conversion, where near complete conversion refers to less discharged tail oil than 100% lightening of the residuum feedstock. In order to limit the recycle accumulation of metal (mainly vanadium, nickel, iron) sulfides from the raw resid, a small amount (about 3 wt.% ratio to fresh resid) of unconverted resid, containing resid hydrocarbons, entrained metal sulfide particles, and other solids, must be vented.

Taking the EST residual oil suspension bed hydrogenation thermal cracking process of Italy Eny company as an example, the amount of the hydrocarbon oil in the discharged solid residual oil VR-OUTS is usually not more than 5%, and is generally only 2.5-3.3 wt% of the weight of the raw residual oil, and the amount is calculated as 3 wt% below, and the once-through conversion rate of the residual oil is 68 wt%. For 100 ten thousand tons/year residual oil raw material, the quantity is 3 ten thousand tons/year, and because the proportion of the quantity of the molybdenum disulfide solid catalyst in the liquid phase to the quantity of the raw oil in the reaction process is a few thousandths, calculated by two thousandths, the discharged solid residual oil VR-OUTS contains 3 ten thousand tons/year hydrocarbon oil and catalyst molybdenum disulfide solid particles (187.5 tons/year) which are 0.00625 times of the hydrocarbon oil.

Taking EST residual oil suspension bed hydrogenation thermal cracking process of Italy Eny company as an example, supposing that the discharged residual oil ratio is reduced to 0.3 wt% of the weight of the raw residual oil, for 100 ten thousand tons/year residual oil raw material, the quantity of the circulating residual oil is increased from 42.65 ten thousand tons/year to 46.62 ten thousand tons/year (the weight increase multiplying factor is only 1.09), the quantity of the discharged tail oil is sharply reduced from 3 ten thousand tons/year to 0.3 ten thousand tons/year (the weight reduction multiplying factor is up to 10 times), the ratio of the quantity of the molybdenum disulfide solid catalyst in the liquid phase to the quantity of the raw oil in the reaction process is two thousandths, so that the discharged solid-containing residual oil VR-OUTS contains 0.3 ten thousand tons/year hydrocarbon oil and the molybdenum disulfide solid catalyst particles (18.75 tons/year) which are 0.00625 times of hydrocarbon oil, from the surface of the calculation result, it seems that the improvement of the utilization ratio of the raw residual oil, the, And at the same time, to reduce the fresh catalyst consumption.

However, the above assumptions do not take into account the following 3 factors caused by the change in cycle oil carbon residue values (related to asphaltene species and concentration):

① on the basis of 97 wt% of the total conversion rate of residual oil, along with the small increase of the total conversion rate of residual oil, the amount of discharged residual oil is sharply reduced, the concentration and carbon residue value of asphaltene in unconverted residual oil can be rapidly increased to form a concentrated solution of asphaltene which is difficult to convert, thereby leading the circulating residual oil to become circulating residual oil with high carbon residue and high concentration of heavy asphaltene which is difficult to hydrocrack;

② the above analysis also shows that if trying to increase the total conversion of the residue oil by 97 wt%, it will cause the stability of the liquid phase in the reaction process to be difficult to control, because the asphaltene concentration and carbon residue value of the cycle residue oil will increase rapidly and it is difficult to control the concentration accurately, for example, if the amount of cycle residue oil is increased by only 9%, the carbon residue content of the cycle residue oil will increase by 0.50-1.0 times or more;

③ on the basis of 97 wt% of the total conversion rate of the residual oil, the amount of discharged residual oil is reduced sharply with a small increase of the total conversion rate of the residual oil, and the concentration of the solid in the unconverted residual oil is multiplied, thereby causing the cycle residual oil to be cycle residual oil with high solid concentration, and there are two influences, on one hand, the concentration effect is favorable for the increase of the concentration of components such as molybdenum sulfide, nickel sulfide, iron sulfide and the like in the liquid phase of the catalyst in the reaction process, and on the other hand, the concentration effect is harmful, namely the concentration effect causes the increase of the concentration of the organic metal sulfide solid particles (metal sulfide without hydrogenation catalyst) carried by the fresh raw material residual oil in the liquid phase of the reaction process, thereby causing the increase of the concentration of the solid entrained in the distillate oil in the fractionation process, and on the other hand, the effect is more obvious, when the conversion rate of the organic metal sulfide (non-molybdenum metal) is 800 μ g/g (calculated as about 1200 μ g/g of sulfide), the extremely difficult processing residual oil (total amount of metal Ni + V is 200 μ g/97 wt% of residual oil, when the conversion rate of the metal is not carried by 0.04.

Due to the limitation of the reasons, the total conversion rate of the raw material residual oil of the EST residual oil suspension bed hydrocracking process is constrained to be 95.0-97.5 wt%, that is, the amount of the hydrocarbon oil discharged OUT of the solid residual oil VR-OUT is usually 5%, and is generally only 2.5-3.3 wt% of the weight of the raw material residual oil, and is calculated according to 3 wt%, and the once-through conversion rate of the residual oil is 68 wt%, and for 100 ten thousand tons/year of residual oil raw material, the amount of the residual oil raw material is 3 ten thousand tons/year, and because the ratio of the amount of molybdenum disulfide solid catalyst in the liquid phase to the amount of the raw material oil in the reaction process is several thousandths, the discharged solid residual oil VR-OUTS contains 3 ten thousand tons/year of hydrocarbon oil and molybdenum disulfide solid particles (187.5 tons/year) carrying 0.00625 times of the weight of the hydrocarbon oil. Because the solid content of the discharged residual oil is too high and the solid oxide of the combustion product of partial sulfide belongs to low-melting-point oxide, the utilization mode of the discharged residual oil can only be used as fuel oil admixture, delayed coking admixture, cement fuel, coal blending coking process, asphalt admixture and the like according to the different properties of the hydrocarbon oil containing the discharged solid residual oil VR-OUTS, and the utilization values of the approaches are not high.

Therefore, the EST residual oil suspension bed hydrogenation thermal cracking process has the following defects:

① because it is difficult to realize high-value application for discharged solid residue oil VR-OUT, the price difference is about 2000 yuan/ton compared with low boiling point hydrocarbon oil, for example, 90% of tail oil can be recovered, i.e. 2.7 million ton/year oil can be increased by 0.54 billion yuan RMB/year;

②, it is difficult to further reduce the catalyst replenishment amount to a large extent;

③ in the process of recycling the metal in the discharged solid residue oil VR-OUT, the problem of the utilization of combustible hydrocarbon oil must be solved, so that the economy of the metal recycling process is low and the working procedure is complex;

④ for extremely difficult processing residual oil with high carbon residue content and high metal content, if trying to pursue the total conversion rate of the residual oil to reach an ideal value such as 97-98 wt%, the stability of the liquid phase in the reaction process is difficult to control, in fact, the total conversion rate is forced to decrease, and the consumption of externally supplied catalyst is rapidly increased.

In fact, the root of the above technical problem lies in the presence of the recycled residue in a single reaction separation section, which has the dual main functions of both recycled residue and recycled catalyst, while reducing the amount of catalyst added requires reducing the rate of discharged residue as much as possible, but reducing the rate of discharged residue inevitably leads to enrichment of unconverted residue with asphaltenes which are difficult to convert. Since engineering technology of the residual oil suspension bed hydrocracking process needs to be optimized and considered by integrating multiple factors, the analysis suggests that the two accumulation effects need to be utilized or responded respectively for the residual oil which is difficult to process and the residual oil which is extremely difficult to process, so that classified processing or classified combined processing to a certain extent is formed. The above analysis suggests that the two main functions of the circulating resid and the circulating catalyst need to be decoupled, one decoupling method is to process the primary hydrocracked product resid separately to form the second reactive separation stage.

The EST residual oil suspension bed hydro-thermal cracking process of Italy Eyni company lacks a secondary oil-solid separation process (avoiding a large amount of circulation of asphaltene) which discharges solid residual oil VR-OUT and has high selectivity and aims at concentrating solid particles of metal sulfides (sulfides of metals carried by fresh residual oil and molybdenum disulfide as a catalyst), and in order to improve the value of hydrocarbon oil in VR-OUT, the oil-solid separation process is a chemical separation process for realizing the lightening (hydrogenation) of residual oil, and a second heavy oil suspension bed hydrogenation reaction separation process is arranged to obtain solid residual oil with higher proportion of light conversion distillate oil and molybdenum disulfide as a catalyst in the conventional heavy oil suspension bed hydrogenation reaction separation process.

A segmented heavy oil suspension bed hydro-thermal cracking reaction separation method can be used, wherein a reaction product obtained by a heavy oil UR10F mainly comprising hydrocarbons with the conventional boiling point higher than 530 ℃ in a first reaction section of a first reaction separation section is separated into a first separation section discharged heavy oil containing a residual oil component and catalyst solid particles in a first separation section US10, and a reaction product obtained by the first separation section discharged heavy oil in a second reaction section of a second reaction separation section is separated into a second separation section residual oil US20-VR containing a residual oil component and catalyst solid particles in a second separation section US 20; US20-VR may be used partly as export heavy oil from the second separation section and partly as long cycle heavy oil entering the first reaction section; compared with a single-stage process, most of residual oil in the discharged heavy oil of the first separation section can be lightened into distillate oil, the consumption of a fresh catalyst can be obviously reduced, the residual oil with higher metal content or higher carbon residue content can be processed, and various combined processes can be formed.

In essence, in the EST resid suspension bed hydrocracking process, the primary hydrocracked product resid (also primary unconverted resid with low catalyst solids concentration) of the resid feedstock and the secondary hydrocracked product resid (which becomes secondary unconverted resid with high catalyst solids concentration when processed separately) of the resid discharged from the first separation section are mixed together, and the process is characterized in that:

① its advantages are that the circulating residual oil and the first residual oil are mixed to proceed suspension bed hydrogenation thermal cracking reaction and product separation, so the process is simple and the investment is saved;

② the second advantage is that the concentration of catalyst solid particles in the liquid phase during the hydro-thermal cracking process of the suspension bed of the primary raw material residual oil (fresh raw material residual oil) is improved, thereby being beneficial to the hydro-thermal cracking reaction of the suspension bed of the primary raw material residual oil (fresh raw material residual oil);

③ has one of the disadvantages that in order to achieve a large reduction in the amount of discharged residue, the proportion (or concentration) of solids in the unconverted residue stream is necessarily increased, and that in turn the proportion (or concentration) of solids in the liquid phase in the entire circulation flow range through which the circulating residue flows is necessarily increased, for example, when the discharge rate of residue is reduced from 3 wt% to 0.3 wt%, the carrying proportion of solid particles in the circulating residue, which are generated by the organometals carried by the raw materials and are 300 mug/g, is increased from 0.01: 1g/g to 0.10: 1g/g, resulting in that the carrying proportion of the corresponding solid particles in the mixed residue of the starting raw materials in the reaction process is increased from 0.0032: 1g/g to 0.032: 1 g/g;

when the raw material residual oil is difficult-to-process residual oil (the total amount of metal Ni and V is 200-800 mug/g) and extremely difficult-to-process industrial residual oil (the total amount of metal Ni and V is 200-800 mug/g), the metal content is 800 mug/g, the concentration of the metal sulfide is corresponding to the value that the metal carried by the raw material residual oil generates about 1200 mu g/g solid particles, the carrying ratio in the circulating residual oil is increased from 0.04: 1g/g to 0.40: 1g/g, which causes the carrying ratio of corresponding solid particles in the mixed residual oil of the initial raw materials of the reaction process to be increased from 0.0128: 1g/g to 0.128: 1g/g, therefore, the high-particle-concentration liquid phase condition is caused in the circulating process of circulating residual oil, the erosion amplitude of instruments such as a high-pressure-drop valve, a high-flow-rate valve, a flow meter and the like is greatly increased, and the problem of preventing the washing of a large number of particles from depositing is caused to the operation of a liquid level measuring and monitoring instrument;

④ the other disadvantage is that, in the fractionation process of the suspension bed hydrocracking product, too high weight concentration (up to more than ten percent) of solid metal sulfide particles increases the probability of carrying solid particles by the distillate oil, thereby increasing the difficulty of removing solid particles in the fractionation process, increasing the concentration of solid particles in the distillate oil product, and affecting the product quality;

⑤ the third disadvantage is that during the fractionation of hydrocracking products in suspension bed processing of high asphaltene content carbon residue, too low yield of unconverted residue will result in too high asphaltene concentration in the unconverted residue, and such recycling residue with too high asphaltene concentration will seriously deteriorate the thermally cleavable properties of the fresh residue in the suspension bed hydrocracking reaction process;

these recycled residue components are not necessarily present as required for material balance, but are merely the result of the cumulative recycle due to the mode of operation; however, in order to process these extrinsic materials, it is usually necessary to use a certain liquid phase molecular concentration of hydrogen donor solvent oil to inhibit coking or to use a certain amount of aromatic-rich wax oil to dilute the residual oil, and the existence of such poor cycle residual oil in large quantities requires the use of a large amount of hydrogen donor solvent oil or wax oil diluent, thereby greatly increasing the operation cost; the first hydrogenation thermal cracking product residual oil is separately processed to form a second reaction separation section, so that the hydrogen supply solvent oil corresponding to the fresh residual oil can be reduced, the quantity of the hydrogen supply solvent oil is greatly reduced, the thermal cracking loss rate of the hydrogen supply solvent oil can be reduced, and the process economic benefit is improved;

in this case, the first reaction separation stage is used as the first hydrocracking process for very poor residues, with the aim of converting the majority of the hydrocarbon oil (but without excessively pursuing the conversion to prevent excessive deterioration of the properties of the primary unconverted residue) under more optimal reaction conditions (medium or low solids content in the reaction liquid phase, medium or low asphaltene concentration); the second reaction separation section is used as a hydro-thermal cracking process of the primary unconverted residual oil, and aims to convert most of hydrocarbon oil under harsh reaction conditions (high solid content in reaction liquid phase, high asphaltene concentration, large amount of hydrogen supply solvent, low conversion per pass and the like), but does not excessively pursue conversion rate to prevent extreme deterioration of properties of the secondary unconverted residual oil, and then discharge the secondary unconverted residual oil with ultrahigh solid concentration in the separation or fractionation process of the second reaction separation section, so as to finally realize relative separation of the hydrocarbon oil and the solid in the primary unconverted residual oil;

surprisingly, if the unconverted residual oil of the second separation section is used as the catalyst circulating material, namely the long circulating residual oil, and returns to the first reaction section, the consumption of the fresh catalyst can be greatly reduced, and the absolute quantity of the poor residual oil brought into the first reaction section can be greatly reduced, so that the first reaction separation section short circulating residual oil is responsible for the circulating residual oil with better circulating property, and the second reaction separation section long circulating residual oil is responsible for efficiently circulating the catalyst, thereby achieving the aim of reducing the consumption of the catalyst.

Compared with EST (expressed sequence tag) residual oil suspension bed hydrocracking process of Italy England company, the sectional process can convert inferior residual oil (with higher metal content or higher asphaltene content), thereby enlarging the residual oil processing range of the suspension bed hydrocracking process.

The invention can use a sectional type heavy oil suspension bed hydrogenation thermal cracking reaction separation method.

The two reaction separation processes comprise a first reaction separation section and a second reaction separation section which takes the discharged vacuum residue US10-VR-OUT of the first separation process US10 of the first reaction separation section UT10 as raw oil UR 20F. A second reaction separation process comprising a second hydrocracking reaction process of raw oil UR20F and a second separation process (usually comprising vacuum fractionation) of heavy oil components and lower boiling hydrocarbon components of the second hydrocracking reaction product, which may include a short circulation process for circulating the unconverted residual oil or modified oil thereof discharged from the second separation process (usually comprising vacuum fractionation) back to the second hydrocracking reaction process for cyclic hydrocracking, in order to prevent the accumulation of sulfide solids, ash, solid particles of catalyst, etc. generated by metals carried by raw oil UR20F, during the second hydrocracking reaction, the second separation process (usually including the second vacuum fractionation process) of the second reaction separation section must discharge discharged heavy oil containing solid particles, such as vacuum residue US20-VR-OUT, mainly composed of unconverted residue. In order to simplify the process, the gas phase material flow of the second reaction process can be recovered through the first reaction process or combined with the gas-containing material of the first reaction process, the gas phase material flow of the separation process of the second reaction product can be recovered through the gas-containing material of the separation process of the first reaction product, and the unconverted residual oil of the second reaction process can be recycled to the first reaction process to reduce the consumption of the catalyst, thereby forming the combined process.

The three reaction separation section processes comprise a first reaction separation section and a second reaction separation section, and also comprise a third reaction separation section which takes the discharged vacuum residue US20-VR-OUT of a second section separation process US20 of the second reaction separation section UT20 as raw oil UR 30F. A third reaction separation process, a third hydrocracking reaction process UR30 containing raw oil UR30F and a third separation process (usually including vacuum fractionation) for separating wax oil component and residual oil component of the third hydrocracking reaction product, which can include a short circulation process for circulating the unconverted residual oil or modified oil thereof discharged from the third separation process (usually including vacuum fractionation) back to the third hydrocracking reaction process UR30 for circulating hydrocracking, in order to prevent the accumulation of sulfide solids, ash, solid particles of catalyst, etc. generated by metals carried in raw oil UR30F, and solid particles existing in the third hydrocracking reaction process, the third separation process (usually including the third vacuum fractionation process) of the third reaction separation section must discharge discharged vacuum residue US30-VR-OUT containing solid particles and mainly composed of unconverted residue. In order to simplify the process, the gas phase material flow of the third section of reaction process can be recovered through the upstream reaction section reaction process or jointly with the gas-containing material of the upstream reaction section reaction process, the gas phase material flow of the third section of reaction product separation process can be jointly recovered through the gas-containing material of the upstream reaction section reaction product separation process, and the third section of unconverted residual oil can be recycled to the upstream reaction section reaction process to reduce the consumption of the catalyst, so that the combined process is formed.

The present invention, if necessary, can constitute a process comprising four or more reaction separation stages, and generally, satisfactory results can be obtained by using two reaction separation stages.

When the sectional type heavy oil suspension bed hydrocracking reaction separation method is applied to the heavy oil hydrocracking process, compared with the prior EST process, the method mainly aims to obtain 1 or more of the following target effects:

① A secondary suspension bed hydrogenation reaction separation process is set for the discharged residual oil of the first separation section, the hydrocarbon oil in the recovered tail oil is converted in a chemical reaction mode, for 100 ten thousand tons/year residual oil raw material, the total conversion rate of the residual oil of the first reaction section is assumed to be 97 wt%, and 0.54 million yuan RMB/year value can be increased if 90% of the recovered tail oil, namely 2.7 ten thousand tons/year oil product, and the benefit is huge;

② if 90% of the discharged residual oil of the first separation section is recovered by the hydro-conversion of the suspension bed of the second reaction section (namely the hydro-thermal cracking conversion rate of the second reaction section is 90 wt%), the weight ratio of the molybdenum disulfide solid particles as the catalyst in the discharged residual oil of the second separation section to the hydrocarbon oil is concentrated by 10 times;

if 95% of the discharged residual oil of the first separation section is recovered through the hydroconversion of the suspension bed of the second reaction section (namely the hydroconversion rate of the second reaction section is 95 wt%), the weight ratio of the molybdenum disulfide solid particles as the catalyst in the discharged residual oil of the second separation section to the hydrocarbon oil is concentrated by 20 times;

the data show that the solid in the discharged residual oil of the second separation section can be concentrated at a high rate, so that the recovery of metal is facilitated, and meanwhile, in the second reaction section, the molybdenum disulfide solid particles which are catalysts with ultrahigh concentration are distributed, so that the processing of the circulating residual oil of the second separation section is facilitated;

the above data also indicate that, in order to achieve high-rate concentration of the molybdenum disulfide solid particles in the discharged residual oil of the second separation section, the second reaction separation section is preferably not diluted by other high-boiling-point hydrocarbon materials with low solid concentration, so as to prevent reduction of the concentration rate of the molybdenum disulfide solid particles in the tail oil of the second reaction separation section and increase of the loss rate of the residual oil;

③ the above data also indicate that this does not necessarily require or limit the manner in which the second reaction separation stage can be operated for further processing or recovery of material (gas phase material or liquid phase material) that is free or substantially free of solid particles, and that the second reaction separation stage can be operated for further processing or recovery of material (gas phase material or liquid phase material) that is free or substantially free of solid particles, either centrally with the same material as the upstream reaction separation stage or through part or all of the flow path of the upstream reaction separation stage to simplify the process;

the above data also indicate that it is not necessary to require or limit the source and the destination of the hydrogen raw material of the second reaction separation section, therefore, the hydrogen raw material of the second reaction separation section can be jointly heat exchanged or heated with the hydrogen raw material of the first section suspension bed hydrocracking reaction process, or be used in series for the second time to form a combined process;

④ it is also clear that it is not necessary to require or limit the way of the advanced process or recovery process of the solid particle-containing material (gas phase material or liquid phase material) discharged from the second reaction separation section, and the solid particle-containing material (gas phase material or liquid phase material) discharged from the second reaction separation section can be treated together with the same kind of material of the upstream reaction separation section or can be passed through part or all of the flow path of the upstream reaction separation section to form a long circulation flow path between different flow paths, that is, the solid-containing material (or residual oil with ultra-high catalyst solid content) of the downstream flow path section can be mixed and returned to the upstream solid-containing material, but the short-circuit of the solid-containing material with low solid content of the upstream flow path section to the solid-containing material with high solid content of the downstream flow path section is limited;

the invention takes the second reaction separation section as a technical approach, and can improve the catalyst circulation efficiency of the circulating oil by times by improving the catalyst concentration in the long circulating residual oil; the concentration of the catalyst in the discharged residual oil of the second separation section can reach 5-20 times of that of the catalyst in the discharged residual oil of the first separation section and can reach tens of to hundreds of times of that of the catalyst carried in the fresh residual oil of the first reaction section, so that the addition amount of the fresh catalyst can be greatly reduced under the condition of realizing the same concentration of the liquid catalyst of the first reaction section, or higher catalyst concentration can be formed in the first reaction section under the condition of the same catalyst consumption, so that the conversion per pass is improved, the quantity of short circulating oil of the first reaction separation section is obviously reduced, the scale, the investment, the energy consumption and the catalyst consumption of the whole process are reduced, and a total flow with higher economy is formed;

⑤ can improve the ability to process inferior fresh residual oil, and can flexibly adjust the operation condition and conversion rate of the first reaction section and the second reaction section according to the nature of the fresh residual oil;

when processing the residue oil which is difficult to process and the residue oil which is extremely difficult to process, in order to prevent the operation condition of the first reaction separation section from deteriorating, prolong the operation period of the first reaction separation section, improve the distillate oil quality of the first reaction separation section and properly reduce the severity of the first reaction section; in the second reaction separation section, the discharged residual oil of the first separation section can be converted under harsh reaction conditions (high solid content in a reaction liquid phase, high asphaltene concentration, use of a large amount of hydrogen supply solvent, low conversion per pass, and the like), most of the hydrocarbon oil (the conversion rate is not excessively pursued to prevent the extreme deterioration of the properties of the secondary unconverted residual oil) is converted, and then the discharged residual oil of the second separation section with ultrahigh solid concentration is discharged in the separation or fractionation process of the second reaction separation section, so that the relative separation of the hydrocarbon oil and the solid in the discharged residual oil of the first separation section is finally realized;

⑥, the operation stability of the process can be greatly improved, and for the discharged residual oil of the first separation section with the flow rate far lower than that of the fresh residual oil, a small amount of hydrogen supply solvent or diluent oil is used, so that the asphaltene concentration can be obviously reduced, and the carrying proportion of solid particles can be obviously reduced, therefore, the economic process is easy to form;

⑦ the process is simplified and the investment and operation cost are reduced by the combination of processes.

From the view point of material flow quantity and component characteristics, the heavy oil suspension bed hydrocracking unconverted residual oil is a residue without macromolecule lightening or a converted substance or a concentrate of macromolecules of a thermal condensate in fresh heavy oil, because the quantity of the heavy oil suspension bed hydrocracking unconverted residual oil is less than that of the fresh residual oil, such as 0.02-0.30 of the weight of the heavy oil, the hydrogen supply solvent is convenient to use under the condition of high solvent-oil ratio, such as reaching 0.5-2.0, to carry out the suspension bed hydrocracking reaction with mild operation condition and carry out more hydrogenation saturation reactions, the high-concentration catalyst condition formed by the solid concentration effect is also beneficial to the hydrocracking reaction process of high-concentration asphaltene, other enriched solid particles have the carrying capacity of possible coke formation, and the synergistic effect among the above elements is beneficial and objective.

In fact, the heavy oil suspension bed hydrocracking process has roughly 8 key problems:

① prevent coking of heavy oil feedstock furnaces, which relates to the problem of how to reduce the outlet temperature of heavy oil furnaces, and also to the problem of how to use coking inhibitors such as hydrogen donor solvents;

② inhibit thermal condensation of the initial thermal cracking process of heavy oil feedstocks, which relates to the problem of how to rapidly supply active hydrogen and how to use the supply for

The problem of hydrogen solvents;

③ preventing the solution system from generating a super saturated asphalt phase, namely a second liquid phase at the end of the thermal cracking reaction process, which relates to the problem of reasonably controlling the conversion per pass and also relates to the problem of timely discharging light saturated hydrocarbon to prevent the light saturated hydrocarbon from reducing the aromaticity of the solution;

④ reduces the yield of tail oil and improves the process economy, which relates to the problem of improving the thermal cracking rate of the latter half of the primary thermal cracking process of fresh heavy oil and how to use hydrogen-donating solvent;

⑤ reduces the yield of tail oil and improves the process economy, which relates to the problems of tail oil upgrading and recycle hydro thermal cracking, and how to use hydrogen donor solvent;

⑥ reduces the cost of the cycle process of the tail oil hydrocracking and improves the process economy, which relates to the combination method of the tail oil modification and cycle hydrocracking reaction process and the fresh heavy oil suspension bed hydrocracking reaction process, and also relates to the problem of how to use the hydrogen supply solvent efficiently;

⑦ if hydrogen donor solvent is used, the method is adopted to shorten the circulation path, reduce the pollution degree and improve the use efficiency;

⑧ it is a problem how to form high-performance heavy cycle solvent oil (anti-coking, rich hydrogen supply and strong dissolving capacity for the colloid asphaltene) under the operation goals of improving the conversion rate of heavy oil hydrocracking and prolonging the operation period, and how to reduce the concentration of the colloid asphaltene in the liquid phase solution and prevent the colloid asphaltene from being separated out to become a second liquid phase.

The present invention has been developed in response to the present teachings, and in particular, in response to the present teachings, a means for alleviating or eliminating some of the problems set forth above is provided.

The invention can be combined with any other suitable residual oil suspension bed hydrocracking process to form a corresponding combined process, and the possible combined technologies are at least:

① in the process of hydrogenation of fresh heavy oil in a suspension bed, a liquid product circulating reactor is used for transferring heat of initial hydrocracking reaction to raw heavy oil, so that the preheating temperature is reduced to 360-400 ℃, coking of a furnace tube of a heavy oil heating furnace is prevented, and coking can be further prevented by using a small amount (such as 5-10% of the heavy oil) of hydrogen supply solvent;

② the method is an essential active scheme that highly dispersed high activity catalyst such as molybdenum catalyst can be used, and hydrogen donor solvent can be used to rapidly provide active hydrogen and inhibit thermal condensation in the initial thermal cracking process of heavy oil raw material, which can significantly improve the properties of primary hydro-thermal cracking tail oil;

③ preventing the solution system from generating a super-saturated asphalt phase, namely a second liquid phase, in the end stage of the thermal cracking reaction process, reasonably controlling the conversion per pass to be within the range of 65-80%, using a liquid product circulating reactor to improve the aromaticity of liquid phase hydrocarbon at the outlet of the reactor, and simultaneously adopting a reaction zone at the end of the reactor to spray stripping hydrogen to discharge light hydrocarbon with high saturation in time to prevent the reduction of the aromaticity of the solution and prevent the reduction of the liquid yield due to over-circulating thermal cracking;

④ may require the use of highly aromatic wax oil, added to the rear reaction section, to safely control the asphaltene concentration in the liquid phase in the reactor within safe limits and to safely carry unconverted asphaltenes out of the reactor;

⑤ reducing the yield of tail oil, improving the process economy, setting a hydrogenation modification reaction process CR of tail oil using hydrogen-supplying solvent oil, carrying out hydrogenation modification process taking aromatic hydrocarbon hydrogenation saturation reaction as main target reaction under the conditions of high catalyst-oil ratio, high catalyst concentration and low reaction temperature, then leading the CRP into a rear reaction zone ARB in a suspension bed hydrogenation thermal cracking reaction process AR to carry out moderate hydrogenation thermal cracking reaction, preventing the excessive thermal condensation reaction caused by too high single-pass thermal cracking rate of the modified tail oil, controlling the proportion of reaction types (hydrogenation saturation and hydrocracking) in the integral circulation hydrogenation process of the tail oil THC-VR, properly increasing the proportion of hydrogenation saturation reaction, reducing the proportion of hydrogenation thermal cracking reaction, realizing multiple batch conversion by increasing the circulation amount, controlling the stability of the solution, reducing the quantity of the discharged tail oil, and improving the quality of the discharged tail oil, wherein the most ideal result is that the tail oil quantity is very small (for example, less than 3% -5%) and only used for discharging solid particles (including catalyst particles, heavy oil metal sulfide as raw material and extremely small amount of coking particles) to prevent the excessive accumulation in the reaction system;

meanwhile, a batch feeding technology can be adopted, so that the coking tendency in the tail oil circulating hydrogenation process is further improved;

⑥ the combination method of the tail oil circulation hydrogenation modification reaction process and the fresh heavy oil suspension bed hydrogenation thermal cracking reaction process can reduce the cost of the tail oil circulation hydrogenation modification reaction process and improve the process economy, and because the hydrogen supply solvent is formed to be used in series twice or even for many times, the hydrogen supply speed and the total hydrogen supply capacity of the hydrogenation area of the circulation path of the hydrogen supply solvent can be obviously improved, and the recycling efficiency is improved;

⑦ provides a highly efficient circulation path for hydrogen donor solvent, which can shorten the length of the circulation path, reduce the investment and energy consumption of the circulation path, reduce the pollution degree of the hydrogen donor solvent and effectively reduce the circulation cost;

⑧ the special reactivation step of the hydrogen donor solvent can be combined with the distillate oil hydrogenation upgrading step, the process integration level is further improved, and the investment and energy consumption of the overall process are reduced;

⑨ the process of heavy oil hydrogenation uses a suspension bed hydrogenation reactor, preferably a liquid product circulation type suspension bed hydrogenation reactor;

⑩ in the total flow of heavy oil processing, it combines with heavy oil catalytic cracking process or heavy oil coking process to reduce coke yield, increase light oil yield, improve the hydro-thermal cracking conversion rate of heavy oil raw material in the hydro-thermal cracking reaction process, and prolong the operation period of heavy oil hydro-thermal cracking reaction process.

The invention can form various combined processes by changing the flow forms of each reaction section or separation section, by jointly processing other hydrocarbon-containing materials suitable for joint processing and by combining the subsequent processing methods of hydrocarbon oil in various thermal high-molecular gases.

In the present invention, diluent oil or a hydrogen donor solvent may be used in each reaction stage.

The hydrogen donor solvent precursors used in the present invention, when including heavy oil catalytic cracking products, such as diesel oil and heavy cycle oil, actually constitute a combined process of a heavy oil catalytic cracking process and a heavy oil hydrocracking process.

The invention uses the asphaltene carrier high aromatic hydrocarbon wax oil, when the heavy oil catalytic cracking product wax oil (heavy cycle oil, clarified oil) exists, the combination process of the heavy oil catalytic cracking process and the heavy oil hydrocracking process is actually formed, the hydro-modification and hydro-thermal cracking of the heavy oil catalytic cracking product wax oil (heavy cycle oil, clarified oil) are realized, on one hand, the processing load of the catalytic cracking reaction process is reduced, the coke yield is reduced, and the liquid yield is increased, on the other hand, the hydrogen supply solvent quantity of the heavy oil hydrocracking process is increased, the yield of thermal condensation products (colloid, asphaltene and coke) of the heavy oil hydrocracking reaction process is reduced, the liquid yield is increased, the quantity of the colloid asphaltene carrier solvent oil in the heavy oil hydrocracking product is increased, the heavy oil hydrocracking conversion rate is favorably improved, and on the whole, the distillate oil yield is favorably improved, the coke yield is reduced, and the process economy is obviously improved.

The residual oil hydrocracking product light wax oil can be used in the catalytic cracking reaction process or catalytic cracking reaction process to produce catalytic cracking gasoline or catalytic cracking diesel oil in high yield, and the catalytic cracking diesel oil can be used as a precursor of a light hydrogen-donating solvent.

According to the invention, when the used asphaltene carrier high aromatic hydrocarbon wax oil has the heavy wax oil as a coking product, a combined process of a heavy oil coking process and a heavy oil hydrocracking process is actually formed, and the hydrogenation modification and the hydrogenation thermal cracking of the coked heavy wax oil are realized, so that the processing load in the coking reaction process is reduced, the coke yield is reduced, the liquid yield is increased, the hydrogen supply solvent quantity in the heavy oil hydrocracking process is increased, the yield of thermal condensation products (colloid, asphaltene and coke) in the heavy oil hydrocracking process is reduced, the liquid yield is increased, the quantity of solvent oil carried by colloid asphaltene in the heavy oil hydrocracking product is increased, the heavy oil hydrocracking conversion rate is favorably improved, and in general, the distillate oil yield is favorably improved, the coke yield is reduced, and the process economy is obviously improved.

Particularly, in the heavy oil hydrocracking reaction process using the hydrogen supply solvent, compared with the conventional heavy oil hydrocracking reaction process without the hydrogen supply solvent, under the condition of the same hydrocracking conversion rate, the hydrogen content of the hydrocracking residual oil is obviously increased, the carbon residue value is obviously reduced, the hydrocracking residual oil serving as the circulating hydrocracking residual oil is easier to hydrocrack, the hydrocracking residual oil serving as the discharged oil can be used as a high-quality gasification raw material, and the coke yield is lower when the hydrocracking residual oil serving as the coking process raw material is used, so that the coking distillate oil yield is favorably improved, if a certain proportion of hydrogenation modified oil of high aromatic wax oil is blended in the coking process raw oil such as a delayed coking process serving as the hydrogen supply solvent, a hydrogenation coking reaction process is formed, and the combined process allows the heavy oil hydrocracking reaction process to process worse (higher carbon residue content), Less expensive residua. The combined process of the heavy oil hydrocracking process and the heavy oil coking process is particularly suitable for combining the newly-built heavy oil hydrocracking process with the existing heavy oil coking process.

Heavy wax oil or its hydrogenated modified oil as heavy oil hydrocracking product, wax oil (heavy cycle oil, clarified oil) as heavy oil catalytic cracking product, and heavy wax oil as heavy oil coking product, can be used as hydrogen-supplying solvent oil with certain hydrogen-supplying capability in a decoking reaction process such as a delayed coking reaction process or a fluid coking reaction process or a flexible coking reaction process, is combined with heavy oil hydrogenation thermal cracking residual oil to carry out a coking reaction process, the coking reaction process also typically combines processing of straight run vacuum residua to control the carbon residue content of the overall coker feedstock or ash content, metal content in the product coke, and after the hydrogen supply solvent is used in the heavy oil hydrocracking process, the use amount of the heavy oil hydrocracking catalyst solid and the coke carrier solid can be reduced, so that the operation effect of the coking reaction process of the heavy oil hydrocracking residual oil can be optimized.

The coal tar pitch-containing hydrocarbon oil R10F refers to a hydrocarbon oil material containing coal tar pitch HDS; the coal tar pitch HDS generally has a hydrocarbon boiling point of > 370 ℃, generally > 400 ℃, in particular > 450 ℃, and contains hydrocarbon components such as colloids, asphaltenes, possibly solid particles, having a conventional boiling point of > 530 ℃.

The upflow type hydrogenation modification reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F refers to an upflow type expanded bed hydrogenation thermal cracking reaction process, such as a suspension bed hydrogenation thermal cracking reaction process, a suspension bed and fluidized bed combined hydrogenation thermal cracking reaction process and the like.

In the upflow expanded bed hydrocracking reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F, hydrodesulfurization reaction, hydrocracking reaction and thermal cracking free radical hydrogenation stabilization reaction of at least part of coal asphalt components HDS are carried out, and at least part of hydrocarbon products with lower sulfur content and lower boiling point are generated; the upflow hydro-upgrading reaction process R10 of the coal-containing tar oil R10F usually does not need to achieve the complete lightening of the single-pass reaction, and usually has a reasonably high single-pass conversion rate of 15% to 55%, so that a certain amount of unconverted coal tar components, such as 45% to 85% of unconverted coal tar components, exist in the hydro-upgrading reaction product R10P, and a hydro-thermal cracking product residual oil THC-VR is formed.

If from the perspective of component structure, the hydrogenated and thermally cracked product residual oil THC-VR is itself a residue of un-lightened macromolecules or a converted matter or a concentrated matter of macromolecules of a thermal condensate in the coal tar pitch component, compared with the same boiling range fraction of the coal tar pitch component HDS of a hydrogenated and thermally cracked precursor thereof, the colloid content, the asphaltene content and the carbon residue content in the hydrogenated and thermally cracked product residual oil THC-VR should not be significantly increased, such as the increase range is limited to be below 5-20%; in fact, it is desirable that the hydrocracked product residue, THC-VR, is itself a residue of unreduced macromolecules or a concentrate of macromolecules of the thermal condensate in the coal tar pitch component, typically having a reduced colloid content, asphaltene content, carbon residue content, such as at least 10% to 25% compared to the same boiling range fraction of the coal tar pitch component HDS, which is a hydrothermally cracked precursor thereof. According to experimental studies, such results limit the hydroconversion of the coal tar pitch HDS to typically 5% to 55%, typically 15% to 45%, and most preferably 20% to 35%.

In order to improve the comprehensive processing efficiency of the device, the hydro-thermal cracking reaction process R10 is preferably used for optimizing the single-pass conversion rate of the coal tar pitch component HDS in the coal tar pitch-containing hydrocarbon oil R10F; the excessive increase of the conversion per pass of the coal tar pitch component HDS in the coal tar pitch-containing hydrocarbon oil R10F inevitably increases the thermal condensation reaction of the super-macromolecules in the coal tar pitch component HDS, increases the quantity of the thermal condensation product colloid, asphaltene and liquid phase coke, increases the gas yield, increases the hydrogenation saturation depth of the hydrocarbons with the conventional boiling point lower than 370 ℃ to form the low-aromaticity hydrocarbons with low dissolving and dispersing capacity to the coal tar pitch, and generates the extraction action to concentrate the coal tar pitch, so that the increase of the quantity of the colloid, the asphaltene and the liquid phase coke and the reduction of the quantity of the suitable solvent oil of the colloid, the asphaltene and the liquid phase coke develop to the super-saturation degree or the critical saturation degree, which can cause the precipitation of the colloid, the asphaltene and the liquid phase coke from a stable colloid solution system to form a super-saturated asphalt phase, namely a second liquid phase, and cause the rapid coking in containers such as a reactor and the like, forcing the plant to shut down.

In fact, the scheme of the invention aims to produce high-quality needle coke raw oil, so that the single-pass conversion rate of the coal asphalt component HDS in the coal-containing asphalt hydrocarbon oil R10F is too high to increase the yield of light distillate oil and reduce the quantity of the target asphalt component, which is basically contradictory and inappropriate to the aim of the invention.

In fact, the scheme of the invention aims to produce more high-quality needle coke raw oil, so that under the condition that the single-pass conversion rate of the coal tar pitch component HDS in the coal tar pitch-containing hydrocarbon oil R10F is the same, and under the condition that the desulfurization rate and the thermal cracking rate are the same, the lower the hydrogenation amplitude is, the better the hydrogenation amplitude is, on one hand, the hydrogen consumption cost can be reduced, and on the other hand, the dehydrogenation task of the delayed coking thermal cracking reaction process for producing the needle coke in the later period can be reduced, so that in general, the higher the selectivity of the desulfurization reaction in the upflow expanded bed hydrogenation thermal cracking reaction process R10 is, the better the high-selectivity desulfurization reaction catalyst is, and the condition of important optimization operation is formed. The present invention preferably uses molybdenum-based catalysts, in particular nano-scale molybdenum-based catalysts.

On the premise of the fact that the increase of the conversion per pass of the coal-containing asphalt hydrocarbon oil R10F inevitably increases the amount of the thermal condensate colloid, the asphaltene and the liquid-phase coke, the method of avoiding the precipitation of the colloid, the asphaltene and the liquid-phase coke inevitably reduces the conversion per pass of the coal-containing asphalt hydrocarbon oil R10F or introduces external solvent oil. When the residual components after separating naphthalene and phenol from high-temperature coal tar are used as raw material oil in an upflow expanded bed hydrocracking reaction process R10, the weight ratio of light fraction (hydrocarbons with the conventional boiling point lower than 370 ℃) to coal tar pitch (hydrocarbons with the conventional boiling point higher than 370 ℃) is usually in the range of 35: 55-40: 50, namely 0.64: 1-0.8: 1, and the adjustment of the conversion per pass of the coal tar pitch-containing hydrocarbon oil R10F is free and passive within a certain range, so that the 'solvent oil (or hydrogen donor) for introducing appropriate external colloid, asphaltene and liquid phase coke and the use method thereof' become an important technical problem in relation to the feasibility, stability and economy of long-term operation of the upflow expanded bed hydrocracking reaction process R10 of the coal tar pitch-containing hydrocarbon oil R10F, and an important object of the invention is that the appropriate colloid for introducing external, Under the condition of solvent oil (or hydrogen donor) of asphaltene and liquid-phase coke, the existing solvent oil use method is improved to reduce the using amount or use cost of the solvent oil (or hydrogen donor), so that the reasonable reduction of the scale of R10 and a product separation system can obviously save investment, reduce hydrogen consumption and energy consumption, and improve the economy of R10.

The diluent KWS in the present invention refers to hydrocarbons entering the upflow hydro-upgrading reaction process R10, which can reduce the viscosity of the overall reaction liquid phase, or reduce the asphalt concentration, or increase the hydrogen supply capacity (coke inhibition) or reduce the reaction temperature, and it is preferable that the introduction of the diluent KWS does not adversely affect or affect the needle coke feedstock fraction, and therefore, the diluent KWS is preferably hydrocarbons having a normal boiling point of less than 320 ℃ and appearing as liquid as possible in the upflow hydro-upgrading reaction process R10, that is, the diluent KWS is preferably hydrocarbons having a normal boiling point of 250 to 320 ℃.

The diluent KWS can be an external supply stream, can be hydrocarbons with a proper boiling range, such as normal hydrocarbons with a boiling point of 250-320 ℃, obtained by separating a reaction product R10P in an upflow hydro-upgrading reaction process R10, and can be hydrocarbons rich in hydrogen-supplying hydrocarbons.

The upflow hydro-upgrading reaction process R10 of the coal-containing bituminous hydrocarbon oil R10F of the present invention is described in detail below.

The coal tar pitch hydrogenation reaction which may be performed by the upflow hydrogenation upgrading reaction process R10 of the coal tar pitch-containing hydrocarbon oil R10F of the present invention is described below.

In the upflow hydrogenation modification reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F, at least part of coal-containing asphalt hydrocarbon oil R10FL undergoes desulfurization reaction or hydrodesulfurization reaction, thermal cracking reaction and thermal cracking free radical hydrogenation stabilization reaction to generate at least part of hydrocarbon products with lower sulfur content and hydrocarbon products with lower boiling point; the upflow hydro-upgrading reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F usually cannot or is not expected to realize the total lightening of a single-pass reaction, namely, a reasonably high thermal cracking single-pass conversion rate is usually existed, so that a certain amount of unconverted coal asphalt exists in a hydro-upgrading reaction product R10P; the unconverted coal tar pitch is separated into refined pitch (light pitch) and heavy pitch in the pitch refining part, and the heavy pitch is used as tail oil; in order to reduce the amount of the discharged tail oil, the hydrocracking reaction process of the tail oil may need to be arranged to produce low-boiling products, and in order to simplify the overall flow and reduce the investment and energy consumption, the hydrocracking reaction process of the tail oil and the up-flow hydrogenation upgrading reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F may form a combined process, i.e., all the processes are combined or part of the processes are combined.

Although the upflow hydrogenation upgrading reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F is aimed at the desulfurization reaction or hydrodesulfurization reaction, thermal cracking reaction, and thermal cracking radical hydrogenation stabilization reaction of macromolecular hydrocarbons, since the hydrogenation catalyst generally used in the upflow hydrogenation upgrading reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F has a hydrofining function itself, and active hydrogen present can also induce the hydrofining reaction of hydrocarbon molecules, some hydrofining reactions (hydrodemetallization reaction, hydrodeoxygenation reaction, hydrodenitrogenation reaction, hydroaromatics partial saturation reaction, and olefins hydrogenation saturation reaction) must also occur in the upflow hydrogenation upgrading reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F.

In the upflow hydro-upgrading reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F, when the supply of active hydrogen is not timely, thermal cracking radicals of colloid and asphaltene undergo condensation reaction to produce molecules or structural groups with higher molecular weight, and the final result of repeated dehydrogenation and condensation reaction is the generation of coke, which is, of course, a negative reaction that needs to be suppressed or reduced.

The main application object of the invention is an up-flow hydrogenation upgrading reaction process R10 of coal-containing asphalt hydrocarbon oil R10F, the number of used reactors can be 1 or 2 or more, and the number of commonly used reactors is 2-4; the reactor operation mode of the upflow type hydro-upgrading reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F can be in any suitable form, and is generally an upflow type expanded bed reactor or an upflow type expanded bed reactor with liquid product circulation, and the whole reaction zone of a single upflow type expanded bed reactor can be artificially divided into 2 or more reaction zones. The control mode of the inlet temperature of any reaction zone of the upflow type expanded bed reactor can be adjusting the temperature or the flow rate of hydrogen, adjusting the temperature or the flow rate of oil products, and certainly, the introduction of a heat exchanger can also be used for heat exchange and cooling.

The upflow hydrogenation modification reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F uses a reactor, the volume ratio of liquid phase and gas phase (or vapor phase) in the reaction space of which can be the case of mainly liquid phase, and defines the 'actual volume of liquid phase/(actual volume of liquid phase + actual volume of gas phase') in the reaction space as the liquid phase fraction KL of the reaction space, the fraction KL is usually more than 0.45, usually more than 0.55, even more than 0.70, so as to form a practically enhanced liquid phase hydrogenation mode, and in order to keep the hydrogen partial pressure of the reaction space high enough, 2 times or more hydrogen addition at different heights of the reactor may be needed.

When the latter half reaction process R10B of the upflow type hydrogenation modification reaction process R10 of the coal-containing bituminous hydrocarbon oil R10F is combined with the heavy oil component CRPVR in the reaction product CRP of the upflow type hydrogenation modification reaction process CR of the inferior heavy hydrocarbon CRF, the residence time of the latter half reaction process R10B usually firstly meets the requirement of controlling the hydrocracking rate of the heavy oil component CRPVR, and the hydrocracking rate upper limit of the heavy oil component CRPVR is usually set to prevent the conversion rate per pass from being too high.

The reactor form of the upflow type hydrogenation modification reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F can be any suitable form, and has various known forms, such as an upflow fixed bed reactor, an upflow type micro-expansion bed reactor, an upflow type moving bed reactor, an upflow type online replacement bed reactor, a fluidized bed reactor, a suspension bed reactor, a combined bed reactor of the fluidized bed and the suspension bed and the combination of specific forms thereof, and most of the coal-containing asphalt hydrocarbon oil R10F has industrial application cases, and forms relatively fixed technical characteristics.

The reaction trend of the aromatic rich wax oil KVGO in the hydrocracking reaction process is described in detail below.

In the upflow hydrogenation thermal cracking reaction process of petroleum heavy oil such as vacuum residue, aromatic hydrocarbon-rich wax oil KVGO is generally used; the aromatic hydrocarbon-rich wax oil KVGO can be an external supply material, and can be product wax oil such as medium wax oil and heavy wax oil in the upflow type hydrogenation thermal cracking reaction process of vacuum residue.

From the perspective of component characteristics, the aromatic-rich wax oil KVGO is a hydrocarbon component which is difficult to hydrogenate and thermally crack than vacuum residue; for the residual oil of the heavy oil hydrocracking product, although the wax oil of the heavy oil hydrocracking product is a fraction rich in aromatic hydrocarbon, the wax oil has much lower colloid and asphaltene contents, a lower carbon residue value and lower thermal condensation sensitivity, and in the hydrocracking reaction process, the hydrogenation reaction and the thermal condensation reaction exist in parallel, but the hydrogenation reaction is dominant.

From the perspective of component characteristics, the wax oil of the heavy oil hydrocracking product is polycyclic aromatic hydrocarbon-enriched hydrocarbon fraction which has the second difficulty of hydrocracking reaction to the hydrocracking reaction of the vacuum residue THC-VR, has higher carbon residue value and thermal condensation sensitivity, and can generate more thermal condensates when repeatedly circulated in the heavy oil hydrocracking reaction process; in the process of hydrogenation thermal cracking reaction, hydrogenation reaction and thermal condensation reaction exist in parallel, namely, parallel hydrogenation lightening reaction and dehydrogenation condensation heaving reaction exist, and the two reaction directions are as follows:

① hydrogenation and lightening reaction, namely the reaction for inhibiting coking, in the direction of 'polycyclic aromatic hydrocarbon → partially saturated polycyclic aromatic hydrocarbon → hydrocarbons with higher hydrogen content';

② dehydrogenation reaction, namely coking reaction, in the direction of "coke ← carbonaceous intermediate phase ← asphaltene component ← colloid ← polycyclic aromatic hydrocarbon".

Therefore, the aromatic hydrocarbon-rich wax oil KVGO is reformed into hydrogen-donating hydrocarbon through a proper hydrogenation modification process and used for the upflow type hydrogenation thermal cracking reaction process of petroleum heavy oil such as vacuum residual oil, has the effects of increasing the quantity of the hydrogen-donating hydrocarbon and reducing the concentration of colloidal asphalt-like components, and has the effects of improving the hydrogenation thermal cracking conversion rate of the residual oil, reducing the gas yield and reducing the coke yield, so that the liquid yield can be increased.

The upflow hydrocracking reaction process CR of the inferior hydrocarbon CRF of the present invention, which may be a heavy oil hydrocracking product vacuum residuum THC-VR, is described in detail below.

The upflow hydrocracking reaction process CR for poor quality hydrocarbon CRF generally includes a thermal cracking reaction for generating thermal cracking radicals, a hydrogenation stabilization reaction for the thermal cracking radicals, and a hydrorefining reaction such as a hydrogenation saturation reaction or a partial hydrogenation saturation reaction for aromatics.

The inferior hydrocarbon CRF upflow hydrogenation modification process CR of the present invention is described in detail below, with the meaning of hydrogenation modification being that partial hydrogenation saturation of aromatics is the desired predominant reaction in the overall hydrogenation reaction.

The inferior hydrocarbon CRF of the present invention, typically the inferior heavy hydrocarbon, generally has the following meanings: under the condition of not using a hydrogen supply solvent and under the same other operating conditions (reaction pressure, reaction temperature, catalyst composition, addition amount, existence amount, retention time, hydrogen-oil volume ratio and reactor operating mode), the coking tendency of the inferior hydrocarbon CRF in the hydrocracking reaction process is more serious, namely the coking rate is higher or the hydrocracking conversion rate is lower than that of the fraction with the same boiling range in the heavy oil R10F; typically, the carbon residue values for hydrocarbons having a normal boiling point above 530 ℃ in poor quality hydrocarbon CRF are higher than the carbon residue values for hydrocarbons having a normal boiling point above 530 ℃ in heavy oil R10F.

Compared with the conventional hydrogenation thermal cracking reaction process, the up-flow hydrogenation modification reaction process CR of the inferior hydrocarbon CRF mainly aims to ensure that the inferior hydrocarbon CRF has more hydrogenation saturation reactions and sufficient thermal cracking free radical hydrogenation stable reactions under the condition of the existence of a catalyst and a possibly existing hydrogen supply solvent, thereby effectively reducing the carbon residue value of the inferior hydrocarbon CRF and leading the hydrocarbons with the conventional boiling point higher than 530 ℃ to become raw materials in the hydrogenation thermal cracking reaction process with proper hydrogenation thermal cracking degree.

In the combined process of the present invention, the first reaction task of the up-flow hydrogenation modification reaction process CR of the inferior hydrocarbon CRF of the present invention is to perform hydrogenation carbon residue removal reaction of the inferior heavy hydrocarbon, i.e. hydrogenation saturation reaction of heavy aromatics, colloids and asphaltenes, and of course, hydrogenation refining reaction (including demetallization hydrogenation hydrogenolysis reaction, olefin hydrogenation saturation reaction, hydrogenation impurity removal (oxygen, sulfur and nitrogen) reaction, hydrogenation aromatic hydrocarbon saturation or partial saturation reaction, hydrogenation carbon residue removal reaction) or hydrocracking reaction can be performed at the same time. The typical feedstock for the upflow hydro-modification process CR of poor quality hydrocarbon CRF is the product residue of the upflow hydrocracking process R10 of heavy oil R10F, which is typically enriched in the bottoms of a vacuum fractionator during product fractionation, such residue THC-VR typically containing added catalyst conversions such as molybdenum sulfide and the like, R10 product metal sulfides from heavy oil R10FL, and coke that may accumulate.

When poor quality hydrocarbon CRF contains hydrocarbons R10-VR with conventional boiling points above 530 ℃ in the R10 product R10P from an upflow hydrocracking process of heavy oil R10F, the carbon residue content of the heavy oil fraction R10-VR is generally higher than the carbon residue content of the same boiling range fraction of the feedstock R10F, or the liquid phase in the hydrocarbon material up-flow type hydrocracking reaction process R10 is suitable for being used as a dispersion solution of colloid, asphaltene and liquid phase coke in the hydrocracking reaction process of the inferior hydrocarbon CRF, therefore, the invention introduces the heavy hydrocarbon in the CR reaction product of the upflow type hydrogenation modification reaction process of poor quality hydrocarbon CRF into the second half of the upflow type hydrogenation thermal cracking reaction process R10 of heavy oil R10F to carry out moderate hydrogenation thermal cracking reaction, the thermal cracking depth of hydrocarbon inferior CRF can be controlled simultaneously to prevent the production of a second liquid phase (bituminous phase) due to the yield of thermal condensate such as asphaltenes exceeding a limit caused by an excessively high thermal cracking rate.

In the up-flow hydrogenation modification reaction process CR of poor quality hydrocarbon CRF, when the supply of active hydrogen is not timely, the thermal cracking free radicals of colloid and asphaltene can produce condensation reaction to produce molecules or structural groups with larger molecular weight, and the final result of repeated dehydrogenation and condensation reaction is the generation of coke, which is, of course, a negative reaction needing to be inhibited or reduced.

The reactor form of the up-flow hydrogenation modification reaction process CR of the inferior hydrocarbon CRF can be any suitable form, and can be one or the combination of a plurality of fluidized bed reactors, suspension bed reactors and combined bed reactors of the fluidized bed and the suspension bed.

The reactor used in the upflow hydrogenation modification reaction process CR of the inferior hydrocarbon CRF can be 1 or 2 or more, the working mode of the reactor can be in any suitable form, generally, the reactor is an upflow expanded bed reactor or an upflow expanded bed reactor with liquid product circulation, and the whole reaction zone of a single upflow expanded bed reactor can be considered to be divided into 2 or more reaction zones. The control mode of the inlet temperature of any reaction zone of the upflow type expanded bed reactor can be the regulation of the temperature or the flow rate of hydrogen, and can be the regulation of the temperature or the flow rate of oil products.

The reactor used in the upflow hydrogenation modification reaction process CR for poor quality hydrocarbon CRF may have a volume ratio of liquid phase to gas phase (or vapor phase) in the reaction space defined as "actual volume of liquid phase/(actual volume of liquid phase + actual volume of gas phase)" in the reaction space, which may be the case of liquid phase being the dominant, as the liquid phase fraction KL of the reaction space, which is usually greater than 0.5, generally greater than 0.65, or even greater than 0.80, creating a virtually enhanced liquid phase hydrogenation mode, and in order to maintain a sufficiently high hydrogen partial pressure in the reaction space, it may be necessary to add hydrogen gas 2 or more times at different elevations of the reactor.

To adjust the reaction feed properties, or to control the liquid phase properties of the reaction process, a portion of the liquid feedstock R10F may be introduced into the upflow, hydro-modification reaction process CR of poor quality hydrocarbon CRF.

To adjust the reaction feed properties, or to control the reaction process liquid phase properties, the intermediate liquid product of the upflow hydrocracking reaction process R10, e.g., a portion of the liquid product of the front reaction section, or the final liquid product, e.g., a portion of the liquid product of the back reaction section, of the heavy oil R10F may be introduced directly into the upflow hydro-modification reaction process CR for poor quality hydrocarbon CRF at elevated temperature and pressure.

In order to shorten the path of residual oil components in the upflow type hydrogenation thermal cracking reaction product R10P of the heavy oil R10F entering the upflow type hydrogenation modification reaction process CR of the poor quality hydrocarbon CRF, when distillate oil (hydrocarbons with the conventional boiling point lower than 530 ℃) contained in the liquid phase of the reaction product R10P is rich in hydrogen donor or hydrogen donor precursor, part of the liquid phase of the reaction product R10P can be directly introduced into the upflow type hydrogenation modification reaction process CR of the poor quality hydrocarbon CRF, and at the moment, a tail oil high-pressure short circulation loop is formed; when the liquid phase of the reaction product R10P contains distillate oil (hydrocarbons with conventional boiling point lower than 530 ℃) containing only a small amount of hydrogen donor or hydrogen donor precursor, it is usually necessary to obtain residual oil or mixed oil of residual oil and heavy wax oil in the produced oil separation and recovery system of the reaction product R10P, and introduce the residual oil or mixed oil of residual oil and heavy wax oil as inferior heavy hydrocarbon into the upflow type hydrogenation modification reaction process CR, and at this time, a conventional long tail oil circulation loop is formed.

A key object of the present invention is to reuse (recycle, serial recycle) the hydrogen donor solvent component in the reaction product CRP as required, so that the heavy oil component in the reaction product CRP can be subjected to the combined hydrocracking in the upflow hydrocracking reaction process R10 (rear reaction stage) of the heavy oil R10F.

The upflow hydrocracking reaction process CR of the inferior heavy hydrocarbon CRF of the present invention, which may be a hydrocracked product vacuum residue THC-VR of the coal tar pitch component, is described in detail below.

The upflow hydrocracking process CR for poor quality heavy hydrocarbon CRF generally includes thermal cracking reactions that generate thermally cracked radicals, hydrogenation stabilization reactions for the thermally cracked radicals, and generally also includes hydrofinishing reactions such as hydrogenation saturation reactions or partial hydrogenation saturation reactions for aromatics.

The inferior heavy hydrocarbon CRF upflow hydrogenation modification process CR of the present invention is described in detail below, with the implication that partial hydrogenation saturation of aromatics is the desired dominant reaction in the overall hydrogenation reaction.

The inferior heavy hydrocarbon CRF of the present invention, typically inferior heavy hydrocarbon, generally has the following meanings: under the condition of not using a hydrogen supply solvent and under the same other operating conditions (reaction pressure, reaction temperature, catalyst composition, addition amount, existence amount, retention time, hydrogen-oil volume ratio and reactor operating mode), the coking tendency of the inferior heavy hydrocarbon CRF in the hydrocracking reaction process is more serious than that of the fraction with the same boiling range in the coal-containing asphalt hydrocarbon oil R10F, namely the coking rate is higher or the coking conversion rate is lower than that of the hydrocracking reaction process; typically, the carbon residue values for hydrocarbons having a normal boiling point above 530 ℃ in poor quality heavy hydrocarbon CRF are higher than the carbon residue values for hydrocarbons having a normal boiling point above 530 ℃ in coal-containing bitumen hydrocarbon oil R10F.

Compared with the conventional hydrogenation thermal cracking reaction process, the up-flow hydrogenation modification reaction process CR of the inferior heavy hydrocarbon CRF mainly aims to ensure that the inferior heavy hydrocarbon CRF has more hydrogenation saturation reactions and sufficient thermal cracking free radical hydrogenation stable reactions under the condition of the existence of a catalyst and a possibly existing hydrogen supply solvent, effectively reduce the carbon residue value of the inferior heavy hydrocarbon CRF and ensure that the hydrocarbons with the conventional boiling point higher than 530 ℃ become the raw materials of the hydrogenation thermal cracking reaction process with proper hydrogenation thermal cracking degree.

In the combined process of the present invention, the first reaction task of the up-flow hydrogenation modification reaction process CR of the inferior heavy hydrocarbon CRF of the present invention is to perform hydrogenation carbon residue removal reaction of the inferior heavy hydrocarbon, i.e. hydrogenation saturation reaction of heavy aromatics, colloids and asphaltenes, and of course, hydrogenation refining reaction (including demetallization hydrogenation hydrogenolysis reaction, olefin hydrogenation saturation reaction, hydrogenation impurity removal (oxygen, sulfur and nitrogen) reaction, hydrogenation aromatic hydrocarbon saturation or partial saturation reaction, hydrogenation carbon residue removal reaction) or hydrocracking reaction can occur at the same time. Typical feedstock for the upflow hydro-upgrading reaction process CR of heavy hydrocarbon CRF of poor quality is the product residue of the upflow hydro-upgrading reaction process R10 of coal-containing asphaltic hydrocarbon oil R10F, which is typically enriched in the bottoms of a vacuum fractionator during product fractionation, such residue THC-VR typically containing added catalyst conversions such as molybdenum sulfide and the like, R10 product metal sulfides from coal-containing asphaltic hydrocarbon oil R10FL, and coke that may accumulate.

When poor quality heavy hydrocarbon CRF comprises hydrocarbons R10-VR with conventional boiling points above 530 ℃ from R10 product R10P from an upflow hydro-upgrading process of coal-containing bitumen hydrocarbon oil R10F, the carbon residue content of the heavy oil fraction R10-VR is generally higher than the carbon residue content of the same boiling range fraction of the feedstock R10F, or the liquid phase in the up-flow hydro-upgrading reaction process R10 of the hydrocarbon material is suitable for being used as a dispersion solution of colloid, asphaltene and liquid-phase coke in the hydro-thermal cracking reaction process of the poor quality heavy hydrocarbon CRF, therefore, the invention introduces the heavy hydrocarbon in the CR reaction product in the upflow hydrogenation modification reaction process of poor quality heavy hydrocarbon CRF into the second half of the upflow hydrogenation modification reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F for moderate hydrogenation thermal cracking reaction, the thermal cracking depth of hydrocarbon heavy hydrocarbon CRF of poor quality can be controlled simultaneously to prevent the production of a second liquid phase (bitumen phase) due to the yield of thermal condensate such as asphaltenes exceeding a limit caused by an excessively high thermal cracking rate.

In the up-flow hydrogenation modification reaction process CR of poor quality heavy hydrocarbon CRF, when the supply of active hydrogen is not timely, the thermal cracking free radicals of colloid and asphaltene can produce condensation reaction to produce molecules or structure groups with larger molecular weight, and the final result of repeated dehydrogenation and condensation reaction is the generation of coke, which is, of course, a negative reaction needing to be inhibited or reduced.

The reactor form of the up-flow hydrogenation modification reaction process CR of the inferior heavy hydrocarbon CRF can be any suitable form, and can be one or the combination of a plurality of suspended bed reactors and combined bed reactors of boiling beds and suspended beds.

The reactor used in the up-flow hydrogenation modification reaction process CR of the poor quality heavy hydrocarbon CRF can be 1 or 2 or more, the working mode of the reactor can be any suitable form, and the reactor is usually an up-flow type expanded bed reactor or an up-flow type expanded bed reactor with liquid product circulation, and the whole reaction area of a single up-flow type expanded bed reactor can be artificially divided into 2 or more reaction areas. The control mode of the inlet temperature of any reaction zone of the upflow type expanded bed reactor can be the regulation of the temperature or the flow rate of hydrogen, and can be the regulation of the temperature or the flow rate of oil products.

The reactor used in the up-flow hydrogenation modification reaction process CR of poor quality heavy hydrocarbon CRF may have volume ratio of liquid phase to gas phase (or vapor phase) in the reaction space, which may be the case of liquid phase being the main, and "actual volume of liquid phase/(actual volume of liquid phase + actual volume of gas phase)" in the reaction space is defined as the liquid phase fraction KL of the reaction space, and the fraction KL is usually greater than 0.5, generally greater than 0.65, even greater than 0.80, so as to form a practically enhanced liquid phase hydrogenation mode, and in order to keep the hydrogen partial pressure of the reaction space sufficiently high, it may be necessary to add hydrogen 2 or more times at different height positions of the reactor.

To adjust the reaction feed properties, or to control the liquid phase properties of the reaction process, a portion of the liquid feedstock R10F may be introduced into the upflow, hydro-modification reaction process CR of poor quality heavy hydrocarbon CRF.

To adjust the reaction feed properties, or to control the liquid phase properties of the reaction process, the intermediate liquid product of the upflow hydro-upgrading reaction process R10, e.g., a portion of the liquid product of the front reaction section, or the final liquid product, e.g., a portion of the liquid product of the back reaction section, of the coal tar pitch-containing hydrocarbon oil R10F may be introduced directly into the upflow hydro-upgrading reaction process CR of heavy hydrocarbon CRF of poor quality at high temperature and pressure.

In order to shorten the path of residual oil components in the upflow type hydrogenation modification reaction product R10P of the coal-containing asphalt hydrocarbon oil R10F entering the upflow type hydrogenation modification reaction process CR of the inferior heavy hydrocarbon CRF, when distillate oil (hydrocarbon with the conventional boiling point lower than 530 ℃) contained in the liquid phase of the reaction product R10P is rich in a hydrogen donor or a hydrogen donor precursor, part of the liquid phase of the reaction product R10P can be directly introduced into the upflow type hydrogenation modification reaction process CR of the inferior heavy hydrocarbon CRF, and at the moment, a tail oil high-pressure conveying short path is formed; when the liquid phase of the reaction product R10P contains distillate oil (hydrocarbons with conventional boiling point lower than 530 ℃) containing only a small amount of hydrogen donor or hydrogen donor precursor, it is usually necessary to obtain residual oil or mixed oil of residual oil and heavy wax oil in the produced oil separation and recovery system of the reaction product R10P, and the residual oil or mixed oil is introduced into the upflow type hydrogenation modification reaction process CR as poor heavy hydrocarbon, and at this time, a conventional long tail oil conveying path is formed.

An important object of the present invention is to reuse (recycle, serial recycle) the hydrogen donor solvent component in the reaction product CRP as required, so that the heavy oil component in the reaction product CRP can be subjected to combined hydrocracking in the upflow hydrocracking reaction process R10 (such as a rear reaction stage) of the coal-containing bituminous hydrocarbon oil R10F.

The hydro-upgrading reaction process R88 using a fixed bed hydrogenation reactor of the present invention is described in detail below.

The hydrogenation upgrading reaction process R88, the number of fixed bed reactors used can be 1 or 2 or more; the operation mode of the catalyst bed layer of the hydrogenation upgrading reaction process R88 can be in any suitable form, can be a combination of two or more different types of reactors, and can be a down-flow fixed bed reactor, an up-flow fixed bed reactor or an up-flow micro-expansion bed, and the operation mode of the catalyst bed layer is generally a down-flow fixed bed. The hydrogenation upgrading reaction process using the fixed bed hydrogenation reactor uses 1, 2 or more fixed bed catalyst beds, and the inlet temperature of the second and subsequent catalyst beds can be controlled by using cold hydrogen or cold oil.

The hydrogenation upgrading reaction process R88 of the fixed bed hydrogenation reactor uses hydrogenation upgrading catalysts which can be 1, 2 or a series combination of a plurality of types of hydrogenation upgrading catalysts, and the hydrogenation activity of the downstream hydrogenation upgrading catalyst is generally equal to or higher than that of the upstream hydrogenation upgrading catalyst along the flow direction of the reactant flow.

In the hydrogenation upgrading reaction process R88 of the fixed bed hydrogenation reactor, the volume ratio of the liquid phase to the gas phase (or vapor phase) in the catalyst bed of the reactor is defined as the liquid phase fraction KL, which may be greater than 0.75, even greater than 0.95, in order to keep the hydrogen partial pressure in the hydrogenation upgrading catalyst bed sufficiently high, and it may be necessary to add hydrogen at the inlet of each upgrading catalyst bed.

Coal tar is described below.

The coal tar of the invention refers to coal tar or fractions thereof from the pyrolysis steps of coal pyrolysis, coal carbonization or coal gas production, and the like, and can be coal tar or fractions thereof at low temperature, which are byproducts of coal gas production, or coal tar or fractions thereof, which are byproducts of coal coking and coal pyrolysis (including low-temperature coking, medium-temperature coking and high-temperature coking), can be mixed oil of the coal tar, and can be extracted oil, such as deasphalted coal tar or fractions thereof, obtained by extracting the coal tar with a light hydrocarbon solvent.

The high-temperature coking belongs to the high-temperature pyrolysis process of coal, and the final temperature of the pyrolysis process is generally more than 900 ℃ and is usually between 1000 and 1400 ℃. The high-temperature coal tar refers to the byproduct crude tar produced in the process of preparing coke and/or urban coal gas by high-temperature pyrolysis of coal. High temperature coal tar in a primary distillation process typically produces the following products: light oil (topping tar), phenol oil, naphthalene oil, light wash oil, heavy wash oil, light anthracene oil, heavy anthracene oil, asphalt and other products, wherein the phenol oil can be further separated into crude phenol and dephenolized oil, and the naphthalene oil can be further separated into crude naphthalene and dephenolized oil. The high-temperature coal tar light fraction refers to: anthracene oil, wash oil, naphthalene oil, decalin oil, phenol oil, dephenolized oil, light oil, and mixtures thereof.

Because the properties of raw coal and the coking or gas-making process conditions are changed within a certain range, the properties of coal tar are also changed within a certain range. The technological conditions and product requirements of the primary distillation process of the coal tar are also changed within a certain range, so that the properties of the light fraction of the coal tar are also changed within a certain range. The specific gravity of the coal tar light fraction is usually 0.92-1.25, the conventional boiling point is usually 60-500 ℃ and is usually 120-460 ℃, the metal content is usually 5-80 PPm, the sulfur content is 0.4-0.8%, the nitrogen content is 0.6-1.4%, the oxygen content is 0.4-9.0%, the water content is usually 0.2-5.0%, and the carbon residue content is usually 0.5-13%.

The medium-low temperature coal tar is a coal tar product from coal pyrolysis or coal gas production or other processes, can be low-temperature coal tar from a low-temperature coking process (the carbonization temperature is lower than 700 ℃) or medium-temperature coal tar (the carbonization temperature is between 700 and 950 ℃) from a medium-temperature coking process or mixed oil of the medium-temperature coal tar and the medium-temperature coal tar, and generally contains a coal tar heavy oil component. As the properties of raw coal and the coking or gas-making process conditions are changed within a certain range, the properties of medium and low temperature coal tar are also changed within a certain range. The medium and low temperature coal tar of the invention has a specific gravity of 0.89-1.15, and usually has a metal content of 5-200 PPm, a sulfur content of 0.1-0.7% and a nitrogen content of 0.6-1.6%. The medium-low temperature coal tar of the invention sometimes has an inorganic water content of 0.2% to 5.0% and sometimes has an organic oxygen content of usually 2.5% to 18%, particularly 3.5% to 10%, more particularly 5% to 10%.

The coal tar generally comprises a mixture of hydrocarbon components with a conventional boiling range of 120-450 ℃ and hydrocarbon components with a conventional boiling range higher than 450 ℃, a light fraction FD1 with a conventional boiling range of 120-260 ℃ contains bicyclic aromatic hydrocarbon fractions, a middle fraction FD2 with a conventional boiling range of 260-370 ℃ contains bicyclic aromatic hydrocarbon fractions and tricyclic aromatic hydrocarbon fractions, a heavy fraction FD3 with a conventional boiling range of 370-450 ℃ contains bicyclic to tetracyclic aromatic hydrocarbon fractions, and a residual fraction FD4 with a conventional boiling range higher than 450 ℃ is a coal pitch fraction.

The coal tar light distillate oil refers to the coal tar distillate oil with the conventional boiling point of 60-480 ℃ and 60-450 ℃ generally, and can be subjected to hydrogenation modification by adopting a fixed bed hydrogenation technology.

The coal tar heavy distillate oil refers to medium-low temperature coal tar distillate with the conventional boiling point usually higher than 370 ℃ and usually higher than 400 ℃, the hydro-thermal cracking process of the coal tar heavy distillate oil refers to a process for producing the coal tar distillate with the molecular weight lower than that of a cracking raw material by at least part of hydro-cracking reaction, the process usually comprises parallel hydro-demetallization, hydro-refining and hydro-thermal cracking reaction, and the proper reactor form is an up-flow expanded bed such as a suspension bed reactor or an ebullated bed reactor.

The residual oil fraction FD4 is usually difficult to realize long-period and high-yield hydrogenation and lightening by adopting a conventional fixed bed technology, so that the conversion is realized by adopting an upflow expansion bed such as a suspension bed or a boiling bed hydrogenation technology, in order to prevent the agglomeration of colloidal asphaltenes from causing unnecessary coking reaction, solvent hydrocarbons with good mutual solubility with the coal residue oil fraction are usually required to be dissolved and dispersed to form a dilute solution of the colloidal asphaltenes, and the solvent hydrocarbons can be heavy fraction FD3 with the conventional boiling range of 370-450 ℃, can also be partially saturated conversion products of hydrogenated aromatic hydrocarbons of the heavy fraction FD3 and the residual oil fraction FD4, and can also be partially saturated conversion products of hydrogenated aromatic hydrocarbons of the middle fraction FD 2. The conversion product of the middle fraction FD2, which is partially saturated with hydrogenated aromatic hydrocarbon, belongs to an excellent hydrogen supply solvent and is rich in hydrogen supply hydrocarbon.

The high aromatic carbon rate inferior hydrocarbon HDS can be 1 kind of wide-cut high aromatic hydrocarbon, and can also be 2 or more kinds of high aromatic hydrocarbons with different boiling ranges; the hydrogenation reaction process carried out in the first hydrogenation reaction process R10 can process 1 kind of inferior hydrocarbon HDS, and can also process 2 or more kinds of inferior hydrocarbon HDS with different boiling ranges; when 2 or more high aromatic hydrocarbons with different boiling ranges are processed in the first hydrogenation reaction process R10, the first hydrogenation reaction process can be 2-path or multi-path feeding, catalyst bed layers or reaction spaces through which raw materials of different paths flow can be the same or different, and can be a parallel hydrogenation relationship, or a series hydrogenation process relationship in which raw materials firstly and secondly enter a plurality of reaction zones respectively, or a relationship in which raw materials are firstly hydrogenated in parallel and secondly hydrogenated products are converged and then hydrogenated jointly, or other more complex combination relationships.

According to the present invention, the coal tar is usually subjected to a process of filtering off solid particles before being subjected to hydrotreating or before being subjected to fractionation.

The coal tar generally contains high-value compounds such as phenol, naphthalene, anthracene and the like, and the high-value compounds can be extracted before entering the hydrogenation process.

The hydro-thermal cracking process of high aromatic hydrocarbons such as coal tar refers to a method for stabilizing the thermal cracking and the hydrogenation of pyrolysis free radical fragments in the presence of hydrogen, and according to the difference of catalysts and the difference of hydrogenation process conditions, a hydrogen donor can be used for inhibiting the thermal condensation of components easy to coke, such as coal pitch, and an up-flow type expansion bed reactor can be adopted.

The high aromatic hydrocarbon can be full distillate oil of coal tar, light distillate oil obtained by fractionation and cutting, heavy distillate oil obtained by fractionation and cutting, or heavy distillate oil containing coal pitch obtained by fractionation and cutting, and the coal pitch is preferably subjected to combined processing in a coal hydrogenation liquefaction process in order to prolong the operation period and optimize the scheduling of raw material hydrocarbon in the hydrogenation process.

The production process of coal tar and coal-based needle coke is described in detail below.

Coal tar is a product of coal coking, coal gas making, coal pyrolysis or other coal thermal processing processes, and because the properties of raw coal change within a certain range, the coking, gas making, pyrolysis or other coal thermal processing process conditions change within a certain range, the properties of coal tar also change within a certain range. The coal tar of the invention has a specific gravity of 0.89-1.30, and generally has a metal content of 5-1200 PPm, a sulfur content of 0.1-1.2%, a nitrogen content of 0.1-1.8% and an aromatic hydrocarbon content of 50-99%. The coal tar of the present invention sometimes has an inorganic water content of 0.2% to 5.0%, and sometimes has an organic oxygen content of 0.05% to 12%, particularly 0.5% to 10%. The coal tar of the invention generally has an ash content of 0.005-5.00%.

The high-temperature coal tar refers to high-temperature coal tar generated in the high-temperature coal coking process, and naphthalene in the high-temperature coal tar is usually recovered before the high-temperature coal tar enters a hydrogenation device due to the high price of the naphthalene component.

The needle coke with excellent performance can be prepared by using coal tar pitch with proper composition and property or adding other blending materials.

The medium-low temperature coal tar is a coal tar product from coal pyrolysis or coal gas production or other processes, can be low-temperature coal tar from a low-temperature coking process (the carbonization temperature is lower than 700 ℃) or medium-temperature coal tar (the carbonization temperature is between 700 and 950 ℃) from a medium-temperature coking process or mixed oil of the medium-temperature coal tar and the medium-temperature coal tar, and generally contains a coal tar heavy oil component. As the properties of raw coal and the coking or gas-making process conditions are changed within a certain range, the properties of medium and low temperature coal tar are also changed within a certain range. The medium and low temperature coal tar of the invention has a specific gravity of 0.89-1.15, and usually has a metal content of 5-200 PPm, a sulfur content of 0.1-0.7% and a nitrogen content of 0.6-1.6%. The medium-low temperature coal tar of the invention sometimes has an inorganic water content of 0.2% to 5.0% and sometimes has an organic oxygen content of usually 2.5% to 11%, particularly 3.5% to 10%, more particularly 5% to 10%.

The high-power and ultrahigh-power graphite electrode made of needle coke has the outstanding advantages of small resistivity, small thermal expansion coefficient, strong thermal shock resistance, high mechanical strength, good oxidation resistance and the like, and compared with the common electrode steel-making, the high-power and ultrahigh-power graphite electrode can shorten the smelting time of electric furnace steel-making by 50-70%, reduce the power consumption by 20-50% and increase the production capacity by 1.3 times. Therefore, high quality needle coke is a high value functional carbon material.

The needle coke is prepared by adopting a liquid-phase carbonization technology, and the coking raw materials are gradually pyrolyzed and polycondensed in the liquid-phase carbonization process to form mesophase spherules. The mesophase globules are fully grown, fused and oriented, and finally cured into a carbon product with a fibrous structure, namely needle coke. The production process of the needle coke comprises 3 parts of raw material pretreatment, delayed coking and calcination.

The coal-based needle coke refers to needle coke prepared from materials obtained by taking coal as a resource, and the invention refers to needle coke prepared from oil products based on coal tar.

A document for recording the raw material properties, processing method and product performance index information of coal-based needle coke is disclosed in publication A01: ①, publication name of coal tar chemical engineering, pages 262 to 268, book code for ② retrieval, ISBN code, 7-5024-.

A document for recording the data about the raw material properties, the processing method and the product performance index information of the coal-based needle coke is disclosed in the publication A02: ① with the names of the publications in delayed coking process and engineering, pages 57-64 and pages 351-372, ② searches for the codes of ISBN, 978-7-80229-456-1, Chinese edition library CIP data core (2007), No. 168082, ③ master code, Dianthus superbus and ④ publication in China petrochemical press.

A document for recording the data about the properties of raw materials, processing methods and performance index information of coal-based needle coke is disclosed in A03: ① publication name, modern coal chemical engineering technical handbook, pages 1408 to 1411, ② retrieval uses a book code, ISBN code, 978-7-122-09636-4, Chinese edition library CIP data core (2010) No. 197010, ③ master code, Hendede, ④ publication and chemical industry Press.

As raw oil for producing needle coke, the composition factors influencing the formation of the intermediate phase mainly comprise:

① active ingredient

If the asphalt contains a small amount of highly reactive components, the crosslinking reaction proceeds smoothly, and the oriented crosslinking and condensation are caused, so that an anisotropic mosaic structure and a fibrous structure are easily obtained. If the highly reactive component present in the pitch controls the crosslinking, it is non-oriented crosslinking and an isotropic carbon and fine mosaic structure is obtained. The latter situation often occurs when low-temperature dry distillation coal tar is used as a raw material for liquid-phase carbonization;

② quinoline insolubles

The quinoline insolubles referred to herein are primary quinoline insolubles such as free carbon and carbon black. Their presence results in a decrease in the activation energy of the reaction and an increase in the reaction rate constant. The presence of primary quinoline insolubles, while beneficial to the nucleation process, is not beneficial to the growth and coalescence of globules. When their mass fraction exceeds 5%, the pellets have a diameter of only a few micrometers and have an onion-type structure. This structure is not easily graphitized.

③ oxygen, sulfur, nitrogen and heterocyclic compounds thereof

It has been reported that the presence of more than 7% oxygen in the pitch results in complete suppression of mesophase conversion, and non-graphitizing coke is produced. The sulfur is added into the asphalt to promote the rapid conversion of the intermediate phase, and when the addition amount exceeds 7 percent, the intermediate phase molecules are widely crosslinked, and finally the non-graphitizable glassy carbon is generated. In addition, sulfur can also cause crystal expansion. The heterocyclic compound containing oxygen, sulfur and nitrogen has high thermal reactivity, participates in the primary generation of mesophase spherule and is enriched in the mesophase spherule, so that a fine grain mosaic structure is generated.

④ metal element

The asphalt mainly contains Na, K, Mg, Ca, Fe, Cu, Al, V, Ni and the like, which can activate asphalt molecules, accelerate the generation and fusion of mesophase spherule to form a mosaic structure and increase the thermal collision coefficient of the needle coke.

The raw material pretreatment process of the coal-based needle coke mainly aims at removing Quinoline Insoluble (QI) in coal tar pitch. The composition of the coal tar pitch has an influence on the formation of the mesophase, for example, active components, quinoline insolubles, metal elements, heterocyclic compounds, and nitrogen, sulfur, oxygen, etc. The main component of the coal tar pitch is aromatic hydrocarbon, but the coal tar pitch contains a certain amount of quinoline insoluble substances, wherein the quinoline insoluble substances include amorphous carbon generated by heating and condensing certain high molecular resin-shaped substances during distillation of coal tar, and coal dust and coke powder brought out from a coke oven chamber along with coal gas. They adhere around the mesophase, hindering the growth and coalescence of the spherical crystals. The needle coke having a good fiber structure cannot be obtained after the coking. Therefore, it is necessary to remove the crude Quinoline Insolubles (QI) that hinder the growth of the spherules and then to conduct the compositional modulation to obtain a raw material that meets the production requirements of needle coke, which is a necessary step for producing needle coke from coal tar pitch. The pretreatment process of the coal-based needle coke raw material mainly comprises a filtration method, a centrifugal separation method, a solvent method and a vacuum distillation method. The solvent method is further classified into a solvent-settling method, a solvent-centrifugation method, a solvent-filtration method, a solvent-flocculation method, a solvent-extraction method, a supercritical extraction method, and the like.

The needle coke production process UT1 based on coal tar pitch without removing quinoline insoluble substances generally comprises a raw material pretreatment process UT1, a delayed coking process UT20 and a calcination process UT 30. The suspension bed hydrogenation modification process R10 of the material R10F containing high-temperature coal pitch by using a hydrogenation catalyst R10C obtains a suspension bed hydrogenation product R10P containing R10C, and the suspension bed hydrogenation product R10P is separated to obtain an oil material R10P-HS-VS containing an asphalt component of the suspension bed hydrogenation modification product containing R10C; according to a conventional solvent method separation method, in a solvent method separation process UT1, oil R10P-HS-VS containing asphalt components of a suspension bed hydrogenation modified product is separated into light asphalt UT1-10-L containing R10C and solvent and heavy asphalt UT1-10-H containing R10C and solvent; the light asphalt UT1-10-L is separated into solvent UT1-10-L-S and needle coke raw material light asphalt tower bottom oil UT1-10-L-MIXP containing R10C in the fractionation process UT1-20, the separation result of the distillation method is that the needle coke raw material light asphalt tower bottom oil UT1-10-L-MIXP contains a large amount of R10C, so that the quality of needle coke is reduced, even needle coke cannot be obtained, meanwhile, the needle coke raw material light asphalt tower bottom oil UT1-10-L-MIXP containing R10C carries a large amount of catalysts R10C, and the fresh catalyst consumption of R10 is greatly increased.

The needle coke production process UT based on coal tar pitch without quinoline insoluble removal generally comprises a raw material pretreatment process UT1, a delayed coking process UT2 and a calcination process UT 3.

In the pretreatment process of the coal-based needle coke by the raw material solvent method, a mixed solvent prepared from aromatic hydrocarbon and aliphatic hydrocarbon is usually used as an extracting agent, and quinoline insoluble substances in the raw material coal pitch are removed through settling separation. There are many methods disclosed for the pretreatment of pitch to obtain needle coke feedstock refined pitch and heavy pitch, and any suitable method can be used in the present invention, typical methods are the following:

① A method for producing coal tar pitch needle coke, which comprises using a mixed solvent of aromatic hydrocarbon and aliphatic hydrocarbon at a ratio of 1: 0.6-1: 1.2 as an extractant;

② A production method and a system of coal-series needle coke of Chinese patent ZL200910198177.6 are characterized in that in the specified process conditions of the pretreatment process of the solvent method, the mass ratio of the coal tar pitch to the solvent is 1: 0.6-1: 2.0, the solvent is formed by mixing aliphatic hydrocarbon and aromatic hydrocarbon, the mass ratio of the aliphatic hydrocarbon to the aromatic hydrocarbon is 1: 0.6-1.4, the initial boiling point of the solvent under normal pressure is more than or equal to 150 ℃, the distillate w/w before 310 ℃ is more than or equal to 95%, the aliphatic hydrocarbon is a petroleum hydrocarbon solvent, the aromatic hydrocarbon is one fraction in the processing process of tar or is formed by mixing several fractions in the processing process of tar, and the distillate w/w of the aromatic hydrocarbon under normal pressure and at the temperature of 235-250 ℃ is 50-85%;

③ Chinese patent ZL201110284485.8 is a process for preparing needle coke raw material by using coal tar pitch, in the specified process conditions of the pretreatment process of the solvent method, pitch and solvent are fully mixed, and insoluble substances in the mixed solution are removed by a physical separation method, wherein the physical separation is centrifugal separation, the viscosity of the mixed solution subjected to centrifugal separation is 40-90 mpa · s, the solvent is a mixture of coal-series light oil and coal-series aromatic oil or a mixture of BTX and coal-series aromatic oil, the mass ratio of the coal-series light oil or the BTX to the coal-series aromatic oil is 20: 80-95: 5, and the mass ratio of the solvent to the pitch is 0.5-10;

④ Chinese patent ZL201110339693.3 is a process for preparing needle coke raw material by utilizing coal tar and heavy phase circulation, and comprises the following steps in the process conditions of the pretreatment process of the coal tar by a solvent method:

step 1, fully mixing coal tar with a first solvent, and removing first solvent insoluble substances in the mixture by adopting a physical separation method to obtain a first clarified liquid;

the first solvent is coal-series light oil, and the mass ratio of the first solvent to the coal tar is 0.05-10;

step 2, separating and removing the first solvent in the first clarified liquid, and hydrogenating the residual components after the first solvent is removed to obtain refined asphalt;

step 3, fully mixing the first solvent insoluble substance obtained in the step 1 with a second solvent, and removing the second solvent insoluble substance in the mixture by a physical separation method to obtain a second clear liquid;

the second solvent is coal-series medium oil, the mass ratio of the coal-series medium oil to the coal-series light oil is 1: 0.1-1: 2, and the mass ratio of the second solvent to the first solvent insoluble substances is 0.5-50;

step 4, separating and removing the second solvent in the second clarified liquid to obtain a refined heavy component;

step 5, mixing the refined asphalt and the refined heavy component to obtain a raw material for preparing the needle coke;

⑤ patent ZL 201310231391.3A process for producing needle coke, wherein the solvent is a mixture of coal-series light oil and coal-series aromatic oil or a mixture of BTX and coal-series aromatic oil under the specified process conditions of the pretreatment process by a solvent method, the mass ratio of the coal-series light oil or the mixture of BTX and coal-series aromatic oil is 0.25-19, the mass ratio of the solvent to the asphalt is 0.5-10, the coal-series light oil is one or more of gas light oil, tar light oil and naphtha light oil, and the coal-series aromatic oil is one or more of washing oil, absorption oil, cresol oil and anthracene oil;

⑥, the chinese patent application No. 201711084211.8, which is a process for producing needle coke by compounding and blending raw materials including wash oil, anthracene oil and asphalt in medium-low temperature coal tar, is characterized in that in the specified process conditions of the pretreatment process by a solvent method, an extraction agent adopted by solvent extraction is a mixture of aromatic hydrocarbon and alkane according to a mass ratio of 1.2-1.8: 1, wherein the aromatic hydrocarbon is one or more of toluene, furfural and N-methylpyrrolidone, and the alkane is any one of N-hexane, N-heptane and cyclohexane;

⑦ Chinese patent application No. 201711084214.1 discloses a process for producing coal-based needle coke from medium-low temperature coal tar pitch, wherein in the specified process conditions of the pretreatment process by a solvent method, an extracting agent is composed of aromatic hydrocarbon and alkane, the mass ratio of the aromatic hydrocarbon to the alkane is 1: 1-5: 1, wherein the aromatic hydrocarbon is any one of toluene, xylene and wash oil, and the alkane is any one of n-hexane, n-heptane and cyclohexane;

⑧ Chinese patent application No. 201810300971.6 needle coke industrial production raw material pretreatment solvent extraction system and method, in the process condition of the pretreatment process of the solvent method, a mixed extraction agent of aromatic hydrocarbon and aliphatic hydrocarbon is used, the aromatic hydrocarbon is wash oil, dephenolized phenol oil or a mixture of the wash oil and the dephenolized phenol oil, and the aliphatic hydrocarbon is aviation kerosene.

There are many disclosed methods for preparing needle coke from coal tar pitch, and any suitable method can be used in the present invention, and the following 2 methods are typical:

① Chinese patent ZL200510136737.7 describes a process for preparing needle coke by using coal tar soft pitch as a raw material, which is characterized in that the inlet temperature of a coking heating furnace is 320 ℃, the outlet temperature is 420-520 ℃, the raw material enters a coking tower, the temperature is increased from 420 ℃ to 440-450 ℃ at the rate of 5 ℃ per hour, then the raw material is fed at a constant temperature for 2-4 hours, then the temperature is rapidly increased to 460-470 ℃, the raw material is fed at a constant temperature for 4-6 hours, then the temperature is rapidly increased to 490-520 ℃, and the total feeding time is 4-36 hours.

② patent ZL201510015130.7 describes a process for co-production of needle coke, mesophase carbon microspheres and high-quality asphalt, wherein mesophase carbon microspheres are produced by mixing heavy phase asphalt produced by raw material asphalt pretreatment and raw material asphalt, and the by-product asphalt for producing the mesophase carbon microspheres is mixed with refined asphalt obtained by raw material asphalt pretreatment to serve as a needle coke raw material, and comprises the processes of raw material asphalt pretreatment, polymerization reaction, separation and drying of polymerization products, by-product asphalt processing, coking and calcining.

In the raw material solvent method pretreatment process of coal-based needle coke, a mixed solvent prepared from aromatic hydrocarbon and aliphatic hydrocarbon is usually used as an extracting agent, and quinoline insoluble substances in raw material coal pitch are removed through sedimentation separation. Although the solvent method for pretreating raw materials of coal-based needle coke is complex in process and large in investment, the method is used in large quantities because of high yield of refined pitch, good quality of needle coke and stable operation, and is used by the Ministry of heat energy of Mitsubishi chemical industry, China midge Steel group, Anshan mountain.

The coal-based needle coke raw oil refers to raw hydrocarbon oil for preparing needle coke based on coal tar, and can be distilled distillate oil obtained by fractionating coal tar, light components obtained by separating coal pitch, namely the light pitch, and can be hydro-modified oil products of hydrocarbon oil based on coal tar.

The conventional boiling range of the suitable coal-based needle coke raw oil is usually 330-530 ℃, namely the conventional boiling range mainly comprises 3-ring aromatic hydrocarbon, 4-ring aromatic hydrocarbon and 5-ring aromatic hydrocarbon, preferably mainly comprises 3-ring aromatic hydrocarbon and 4-ring aromatic hydrocarbon, and the higher the aromaticity of the hydrocarbon component is, the better the aromaticity is, the better the needle coke yield in the coking process is.

The coal-based needle coke raw oil of the invention generally refers to raw hydrocarbon oil from high-temperature coal tar for preparing needle coke, and can be distillate oil obtained by fractionating high-temperature coal tar, light component obtained by separating high-temperature coal tar coal pitch, namely the light pitch, and can be oil products or isolates thereof after hydrogenation modification of hydrocarbon oil based on high-temperature coal tar.

The raw material for producing needle coke is usually refined pitch (light pitch) obtained by pretreating coal pitch, and the index of the raw material for producing needle coke is generally shown in Table 1.

In the delayed coking process of coal-based needle coke, the raw materials are converted into gas, gasoline, kerosene, diesel oil, wax oil and coke. The technological process and equipment for delayed coking of needle coke are basically the same as those for delayed coking of ordinary petroleum coke, and the main equipment includes heating furnace, coke tower and fractionating tower, but some necessary measures are taken in individual equipment, coking condition and operation, such as controlling temp. raising speed of feeding material, pressure of coking tower, gas injection quantity and regulating circulation ratio to make the oil material maintain a relatively stable state in the coking tower, fully utilizing the plastic flow of intermediate phase material and ordering of molecular arrangement, at the same time making the gas phase product produce shearing force to create the condition of so-called "gas flow coke-drawing" so as to finally form needle coke with streamlined structure. The needle coke produced in the delayed coking process contains high moisture and volatile components, is called green coke, and can be used as a raw material of a high-power and ultrahigh-power graphite electrode only by carrying out high-temperature calcination treatment under the condition of air isolation.

During the calcination process of the coal-based needle coke, the structure and the element composition of the needle coke are changed in a series, so that the physical and chemical properties of the needle coke are improved. The purpose of calcination is to remove moisture and volatiles from the green coke and to improve the carbon content, density, strength, conductivity and chemical stability of the coke. Needle coke calcination is usually carried out in rotary kilns or rotary kilns, where raw coke enters from one end of the kiln and contacts with the exhaust gas of high-temperature calcination, at the outlet end there is a gas-fired oil burner, and the temperature of the calcination zone can be as high as 1500 ℃. The residence time and the heating speed of the coke in the kiln are determined by the rotation speed of the kiln body, the true density of the calcined needle coke is an important evaluation index of the calcination effect, and the calcined needle coke with the true density of more than 2.130 g/cc belongs to needle coke with good quality.

TABLE 1 technical indices of needle coke feedstock

Table 2 shows the technical indexes of coal-based needle coke in the national standard GB/T32158-2015 of the people's republic of China.

As can be seen from Table 2, the increase in ash content from 0.3 mass percent to 0.4 mass percent will result in a reduction in the needle coke product grade from premium to premium, and therefore, in the upgrading of coal tar pitch, the introduction of extraneous ash material such as inorganic solid particles is minimized. That is, the method has great economic significance for effectively reducing the ash content in the coal-based needle coke raw oil by adopting a proper method.

TABLE 2 technical indices of coal-based needle coke (GB/T32158-

In the calcining process of the coal-based needle coke, the release of sulfur needs higher temperature to break C-S chemical bonds, and generally, the sulfur can be released in a large amount in a high-temperature range of 1200-1500 ℃, so that the sulfur content of a calcined needle coke product of high-sulfur needle coke (green coke) produced by high-sulfur-content raw material coal pitch is high, the expansion coefficient of a high-power and ultrahigh-power graphite electrode prepared by the high-sulfur needle coke product is higher, and the crystal expansion phenomenon exceeding the use upper limit is generated, so that the use function is greatly influenced, and as can be seen from table 2, the sulfur content is increased from 0.4 percent (mass fraction) to 0.5 percent (mass fraction), so that the product grade of the needle coke is reduced from primary grade to secondary grade, and the price is greatly reduced. Therefore, the method has great economic significance for effectively reducing the sulfur content in the coal-based needle coke raw oil by adopting a proper method.

At present, the industrial method for preparing needle coke raw oil based on high-temperature coal tar pitch comprises a flash evaporation method and a solvent method, and the main aim is to control the quinoline insoluble content of the needle coke raw oil. The flash evaporation method cuts a section of raw material suitable for producing needle coke from coal tar pitch through vacuum distillation, and separates out residues (including quinoline insoluble QI). The solvent method uses aliphatic hydrocarbon and aromatic hydrocarbon to prepare a mixed solvent according to a certain proportion, the boiling point of the mixed solvent is generally lower than that of the coal pitch to be treated so as to facilitate the distillation separation operation in the later period, the coal pitch is treated by the mixed solvent to remove quinoline insoluble substances, and the method has the advantages of high yield of needle coke raw materials (fine materials), good quality of needle coke products and the defects of complex process, high control requirement of operation parameters and higher investment cost.

Neither the flash evaporation method nor the solvent method belongs to a physical separation method without changing the molecular structure, so that the sulfur content of the coal-based needle coke raw oil is difficult to be greatly reduced, and the medium pitch molecules with a slightly larger molecular weight and a proper structure cannot be converted into the light pitch suitable for the small coal-based needle coke raw oil, so that the sulfur content limits the source of high-temperature coal tar for producing high-quality needle coke, and the utilization rate of the high-boiling-point medium pitch of the high-temperature coal tar with low sulfur content, low metal content, low ash content and low quinoline insoluble content cannot be improved.

Table 3 shows the properties of a typical high quality oil-based needle coke feedstock. ,

TABLE 3 typical Properties of high-quality oil-based needle coke feedstock

The properties of a typical high-quality oil-based needle coke raw material are shown in Table 3, and the catalytic cracking clarified oil is taken as an example, the aromatic content of the catalytic cracking clarified oil is only 61.7 percent and is far lower than the aromatic content of the high-quality coal-based needle coke raw material, which shows that the needle coke raw material with better comprehensive performance can be prepared by reducing the content of impurities (sulfur and nitrogen) as much as possible under the condition of ensuring that the saturation depth of aromatic hydrocarbon is proper and low through a hydrogenation impurity removal (desulfurization and denitrification) reaction process.

By combining the analysis, the invention provides a moderate hydro-upgrading method of high-temperature coal tar or high-temperature coal tar pitch, which reduces the sulfur content or and the metal content or the ash content or the quinoline insoluble content or increases the oil content suitable for producing high-quality coal-based needle coke through a hydro-upgrading reaction process R10, thereby widening the initial raw material range of the high-quality raw oil of the coal-based needle coke or improving the yield of the high-quality raw oil of the coal-based needle coke obtained from the suitable high-temperature coal tar, and simultaneously obtaining the hydro-upgraded heavy pitch. Compared with unhydrogenated initial oil with the same boiling range, the hydro-modified heavy asphalt has high hydrogen content, low viscosity and enhanced liquidity, so that the hydro-modified heavy asphalt can be used as a blending material of other coal asphalt, such as fuel oil, liquid asphalt and the like prepared by blending the hydro-modified heavy asphalt with coal asphalt of medium-temperature coal tar obtained by medium-temperature pyrolysis of long-flame coal, thereby improving the asphalt value and market price of the hydro-modified heavy asphalt.

Because high-temperature coal tar or high-temperature coal tar coal pitch contains more colloid, asphaltene and solid particles, the appropriate operation mode of the hydrogenation modification reaction process R10 is a suspension bed hydrogenation process.

In order to flexibly control the sulfur content and the nitrogen content of the needle coke delayed coking raw material light asphalt, the invention can carry out moderate hydrogenation refining on the suspension bed distillate oil product or and the light asphalt to carry out moderate hydrogenation desulfurization reaction and hydrogenation denitrification reaction, and the project investment can be greatly reduced by completing the tasks in a combined process mode. And the high-temperature coal tar fractionation process or the hydro-upgrading reaction process of distillate oil obtained in the fractionation process of a suspension bed hydrogenation product R10P can be combined together.

The invention can jointly process medium-temperature coal tar distillate or medium-temperature coal tar pitch with proper components and properties, such as the coal pitch of medium-temperature coal tar generated in the fast pyrolysis process of a long-flame coal fluidized bed.

There are many medium-low temperature coal tar suspension bed hydrogenation process methods and high-temperature coal tar suspension bed hydrogenation process methods, and the following are typical 2 methods:

① patent ZL 201010217358.1A coal tar suspension bed hydrogenation method of heterogeneous catalyst, including coal tar raw material pretreatment and distillation separation, coal tar heavy fraction suspension bed hydrocracking and light distillate oil conventional upgrading process, wherein the suspension bed hydrogenation reaction temperature is 320-480 ℃, the reaction pressure is 8-19 MPa, the volume space velocity is 0.3-3.0 h < -1 >, the hydrogen oil volume ratio is 500-2000, the catalyst is a powdery particle coal tar suspension bed hydrogenation catalyst containing molybdenum, nickel, cobalt, tungsten or iron single metal active component or composite multi-metal active component, the addition amount is that the weight ratio of the active component metal amount and the coal tar raw material is 0.1: 100-4: 100, most of the tail oil containing the catalyst after the hydrogenation reaction product is separated out light oil is directly circulated to the suspension bed reactor, after a small part of the tail oil is subjected to catalyst removal treatment, the suspension bed reactor is further lightened, the heavy oil is completely or maximally circulated, the purpose of producing light oil and catalyst recycling at maximum amount from coal tar is realized, and the utilization efficiency of raw material and catalyst is greatly improved.

② Chinese patent ZL201210022921.9 discloses a hydrogenation and lightening method of heavy oil with low hydrogen content using hydrogen-supplying hydrocarbon, wherein the hydrogen-supplying hydrocarbon material flow rich in hydrogen-supplying hydrocarbon is used in the hydrogenation and lightening process of heavy oil such as coal pitch, which has the effects of inhibiting the condensation and coking speed, improving the yield of liquid products in the coal tar heavy oil hydrogenation and transformation process, improving the product quality, reducing the reaction temperature rise and enhancing the operation stability and safety of the device.

Chinese patent ZL201010217358.1, chinese patent ZL201210022921.9 and similar patent processes are mostly targeted methods for producing light distillate to the maximum extent, and therefore, matched catalysts, process conditions and equipment have been developed. The method of the present invention can realize a suitable deep hydrogenation process based on the prior art, and can realize high selectivity for hydrodesulfurization as much as possible, so that the present invention is feasible.

In order to prevent the deposition of high-temperature coal pitch in the reaction system or thermal polycondensation, a diluent or a hydrogen donor is added as necessary during the reaction (to the reaction raw material, the reaction intermediate, the reaction final product) to dilute the coal pitch.

The basic idea of the invention is: a combined method of high aromatic hydrocarbon suspension bed hydrogenation and solvent method needle coke raw material extraction process is characterized in that under the condition of existence of diluted hydrocarbon or and hydrogen-supplying hydrocarbon, a high-temperature coal tar suspension bed hydrogenation modification reaction process R10 is used for reducing sulfur content or and reducing metal content or and reducing ash content or and reducing quinoline insoluble content, and a product R10 in the reaction process R10P is separated to obtain coal-based needle coke raw material oil with proper sulfur content or metal content or ash content or quinoline insoluble content, so that high-quality coal-based needle coke is prepared.

Since the operation goal of the hydro-upgrading reaction process R10 is to obtain hydrocarbons with high aromatic degree and low sulfur content, the ideal hydrogenation depth is a moderate hydrogenation, and the optimized operation conditions (such as catalyst type, catalyst amount, reaction pressure, reaction temperature, reaction time, hydrogen/oil volume ratio, and reactor operation mode) need to be selected according to the material properties (such as sulfur content and colloid asphaltene content) of the specific suspension bed hydrogenation process R10 and the desired product properties (such as sulfur content, colloid asphaltene content and quinoline insoluble content) and reaction depth (such as heavy asphalt cracking rate) of the suspension bed hydrogenation process R10.

Because high-temperature coal tar or high-temperature coal tar coal pitch contains more colloid and asphaltene and only contains a small amount of light distillate oil with low viscosity, the suitable operation mode of the hydrogenation modification reaction process R10 is a suspension bed hydrogenation process.

Since the operation goal of the hydro-upgrading reaction process R10 is to obtain an upgraded fraction with proper properties and a boiling range of 330-530 ℃, in order to reduce the energy consumption for separating other materials added, the added diluted hydrocarbon or hydrogen donor hydrocarbon preferably mainly comprises the conventional hydrocarbon with a boiling point lower than 330 ℃; in order to reduce the flow range of the recycled materials and reduce the investment and energy consumption, the invention recommends a short-flow circulating process, because the operation target of the hydro-upgrading reaction process R10 is to expect to obtain an upgraded fraction with a boiling range of 330-530 ℃ and the larger the boiling point difference is, the better the energy consumption is for reducing the separation of other materials added.

As a method for optimizing the operation, in order to improve and reduce undesired chemical reactions (excessive hydrogenation saturation reaction, excessive thermal cracking reaction, excessive thermal condensation reaction), it is necessary to shorten the reaction time, that is, to shorten the thermal action time, and therefore, it is necessary to use a highly efficient catalyst (highly active catalyst with high degree of dispersion) and to use a process technique for improving the catalyst efficiency (discharging reaction products that suppress the catalyst activity).

The method of the invention can produce high-quality high-valence carbon materials with high value, such as needle coke, and simultaneously obtain coking distillate oil in the coking process of producing the needle coke, and can carry out combined hydrogenation modification together with the distillate oil obtained in the fractionation process of high-temperature coal tar or the fractionation process of a suspension bed hydrogenation product R10P.

In order to flexibly adjust the sulfur content or the nitrogen content of the delayed coking raw material for producing the needle coke, at least a part of light asphalt or component oil thereof can be introduced into a fixed bed hydrogenation modification process R600 to carry out hydrogenation modification with proper hydrogenation depth for carrying out hydrogenation desulfurization reaction and hydrogenation denitrification reaction, and then the product of the hydrogenation modification process R600 is separated to obtain the hydrogenation modification raw material R600-KP for producing the needle coke.

For the present invention, the main purpose of the upflow hydrogenation process of the material of coal pitch containing high temperature coal tar is to perform thermal cracking desulfurization or and hydrodesulfurization in a suitable depth, simultaneously carrying out hydrogenation demetalization reaction to a certain degree, simultaneously carrying out hydrogenation aromatic hydrocarbon partial saturation reaction, hydrogenation thermal cracking reaction, thermal cracking reaction and hydrogenation stabilization reaction of thermal cracking free radicals of a certain degree of high boiling point hydrocarbon components, producing a suspension bed hydrogenation product with low sulfur content, low metal content and more suitable components for needle coke, therefore, the suspension bed hydrogenation product may need to be fractionated firstly, then separating the obtained hydrogenated coal pitch into primary hydrogenated refined pitch and primary hydrogenated heavy pitch, then the primary hydrogenated heavy asphalt is returned to the suspension bed hydrogenation modification process for secondary processing, thereby improving the beneficial overall thermal cracking conversion rate. As for the overall effects of the primary hydrocracking of the coal tar pitch and the secondary hydrocracking of the tail oil of the primary hydrocracking of the coal tar pitch, the hydrogen supply solvent can effectively improve the overall hydrocracking conversion rate, effectively reduce the yield of the externally thrown solid tail oil and further improve the economical efficiency of the process.

The invention can jointly process medium-temperature coal tar distillate or medium-temperature coal tar pitch with proper components and properties, such as the coal pitch of the medium-temperature coal tar generated in the fast pyrolysis process of the long-flame coal fluidized bed, and the composition and the properties of the medium-temperature coal tar are slightly different from those of the high-temperature coal tar.

The basic attributes of a suitable catalyst for the upflow hydro-upgrading process R10 for coal-containing tar oil R10F are described in detail below.

As mentioned above, the main purpose of the upflow hydrogenation modification reaction process R10 of the coal-containing asphalt hydrocarbon oil R10F is to perform thermal cracking desulfurization or hydrodesulfurization and hydrodemetallization at a suitable depth, and simultaneously perform hydrodemetallization reaction to a certain extent, and simultaneously perform hydrogenation aromatics partial saturation reaction, hydrogenation thermal cracking reaction, thermal cracking reaction and hydrogenation stabilization reaction of thermal cracking radicals to a certain extent for high boiling hydrocarbon components, so as to produce a suspension bed hydrogenation product with low sulfur content, low metal content and more suitable components for needle coke, and therefore, naturally put forward the following requirements for the catalyst:

① adopts high-activity catalyst to reduce the use ratio, thereby reducing the carrying capacity in the refined asphalt;

② high dispersity, i.e. small particle size, increases the external surface area of unit weight of catalyst, reduces the use ratio, thereby reducing the carrying amount in refined asphalt;

③ the coking rate on the surface of the catalyst is low, and the coke is not deposited in the coke as much as possible;

④ the difference between the specific gravity of the catalyst and the specific gravity of the coal tar pitch is enough to ensure the sedimentation separation;

⑤ the ability to promote thermal cracking reactions is minimized.

The coal tar of the invention refers to coal tar containing components suitable for hydro-upgrading, which is from coal pyrolysis or coal gas making or coal coking or coal tar products of other processes, the coal tar of the invention can be low-temperature coal tar from a low-temperature coking process or medium-temperature coal tar from a medium-temperature coking process or high-temperature coal tar from a high-temperature coking process, the coal tar of the invention can be a separation product of the coal tar, the coal tar of the invention can be a mixture of the coal tar, and the coal tar of the invention can be a thermal processing product of the coal tar.

The coal tar containing the coal pitch component contains the pitch component from the coal tar.

Coal pyrolysis processes, including fixed bed processes, moving bed processes, fluidized bed processes, rotary kiln processes, ebullated bed processes, and any other suitable process that may be used.

Coal gas-making processes, including fixed bed processes, moving bed processes (e.g., lurgi gasifier processes), fluidized bed processes, rotary kiln processes, ebullated bed processes, and any other suitable process that may be used.

Coal coking processes, including fixed bed processes, moving bed processes, fluidized bed processes, rotary kiln processes, ebullated bed processes, and any other suitable process may be used.

Because the properties of raw coal vary within a certain range, and the coking, gas-making, pyrolysis or other thermal processing conditions vary within a certain range, the properties of coal tar also vary within a certain range. The coal tar of the invention has a specific gravity of 0.89-1.30, and generally has a metal content of 5-1200 PPm, a sulfur content of 0.1-1.2%, a nitrogen content of 0.1-1.8% and an aromatic hydrocarbon content of 50-99%. The coal tar of the present invention sometimes has an inorganic water content of 0.2% to 5.0%, and sometimes has an organic oxygen content of 0.5% to 11%, particularly 3.5% to 10%, more particularly 5% to 10%. The coal tar of the invention generally has an ash content of 0.005-5.00%.

The aroma degree, the content of the miscellaneous elements and the molecular structure of the high-temperature coal tar pitch (the conventional boiling point is more than or equal to 450 ℃) and the aroma degree, the content of the miscellaneous elements and the molecular structure of the medium-low temperature coal tar pitch (the conventional boiling point is more than or equal to 450 ℃) have huge difference, and the hydrogenation test of the suspension bed shows that:

① high-temperature coal tar pitch is difficult to realize high-conversion rate hydro-thermal cracking, the yield of hydrocarbon components with the conventional boiling point of less than or equal to 450 ℃ is low (the hydro-thermal cracking rate of the high-temperature coal tar pitch with the conventional boiling point of more than or equal to 450 ℃ is only 65-70%), and the hydrogen content is low, so that the high-temperature coal tar pitch is poor in property when being used as a raw material for producing clean fuel oil;

the high-temperature coal tar pitch is subjected to shallow hydrogenation modification reaction with appropriate depth, so that a modified needle coke raw material can be obtained, and meanwhile, high-boiling-point coal pitch components can be subjected to hydrogenation modification to a certain degree, namely, the hydrogen content of the high-boiling-point coal pitch components is increased, the viscosity is reduced, so that the use value of the high-temperature coal tar pitch is improved;

②, the low-temperature coal tar pitch is easy to realize high conversion rate hydro-thermal cracking, the yield of hydrocarbon components with the conventional boiling point of less than or equal to 450 ℃ is very high (the hydro-thermal cracking rate of the high-temperature coal tar pitch with the conventional boiling point of more than or equal to 450 ℃ can reach 85% -95%), the hydrogen content is high, and the high-temperature coal tar pitch can be converted into clean fuel oil after further hydro-upgrading;

the medium-low temperature coal tar pitch is unsuitable for being used as needle coke raw materials because macromolecular aromatic hydrocarbons (tetracyclic aromatic hydrocarbons and pentacyclic aromatic hydrocarbons) belong to aromatic hydrocarbons with multiple side chains or long side chains and contain too much organic oxygen.

The above analysis shows that the upflow hydrogenation process of the asphalt material with huge difference in aromaticity needs to be separately arranged to prevent the products (especially the high boiling point product hydrocarbons) from mixing, that is, when the products with respectively proper use difference are needed to be produced, the device scale cannot be increased, the investment can be reduced, and the energy consumption can not be reduced through simple mixing processing.

At present, high-temperature coal tar and medium-and-low-temperature coal tar are produced simultaneously in northern Shanxi, inner Mongolia Ordos areas and inner Mongolia Ughai areas of China. Because the coal tar yield in the coal processing process is low, generally only 2.0-10.0% (of coal raw materials), the coal tar yield of a single coal processing plant is usually very low, even if the sum of the yields of medium and low temperature tars in a region reaches 200 ten thousand tons/year, the quantity of the coal pitch of the corresponding medium and low temperature tars is only 30-40 ten thousand tons/year, even if the sum of the yields of high temperature tars in a region reaches 100 ten thousand tons/year, the quantity of the coal pitch of the corresponding high temperature tars is only 40-45 ten thousand tons/year, therefore, the hydroconversion scales of the coal pitch of the medium and low temperature tars and the coal pitch of the high temperature tars are very low, and the process combination or sharing of the two reaction processes and the separation processes of the reaction products of the two reaction processes is carried out to the maximum extent from the viewpoints of reducing the unit processing amount investment and the unit processing amount energy consumption of a commercial device, or the secondary use of hydrogen and other material flows is carried out, thereby simplifying the flow, reducing the investment, reducing the energy consumption and improving the overall process economy, which becomes the necessity of technical development.

The characteristic parts of the present invention are described below.

① suspension bed hydrogenation reactor;

③ combined hydrogenation reactor of suspension bed and boiling bed;

④ micro-expanded bed.

The general control principle of the gas phase hydrogen sulfide concentration in the hydrogenation reaction process of the present invention is described in detail below.

Any make-up sulfur may be added to any of the hydrogenation processes as desired, but is typically added to the uppermost hydrogenation process inlet to ensure that the minimum hydrogen sulfide concentration necessary for the reaction process, such as 500ppm (v) or 1000ppm (v), or a specified value, to ensure that the hydrogen sulfide partial pressure necessary for the catalyst does not fall below the minimum specified value. The supplementary sulfur may be hydrogen sulfide or a material which can be converted into hydrogen sulfide and has no adverse effect on the hydroconversion process, such as hydrogen sulfide-containing gas or oil, or carbon disulfide or dimethyl disulfide which generates hydrogen sulfide after contacting with high-temperature hydrogen. When the dilute hydrocarbons of the prehydrogenation process R1 are provided as a hydrogenation effluent containing hydrogen sulfide, the sulfur replenisher may not be used if the amount of hydrogen sulfide therein meets the requirements of the prehydrogenation process R1.

The general principles of the high pressure separation process of the hydrogenation reaction effluent of the present invention are described in detail below.

The high-pressure separation process of the hydrogenation reaction effluent usually comprises a cold high-pressure separator, when the density of the hydrocarbon oil in the hydrogenation reaction effluent is high (for example, the density is close to the water density) or the viscosity is high or the hydrocarbon oil is difficult to separate by emulsification with water, a hot high-pressure separator with the operation temperature usually being 150-450 ℃ needs to be arranged, at the moment, the hydrogenation reaction effluent enters the hot high-pressure separator to be separated into hot high-pressure gas mainly comprising hydrogen in volume and hot high-pressure oil liquid mainly comprising conventional liquid hydrocarbon and possibly existing solids, the hot high-pressure gas enters the cold high-pressure separator with the operation temperature usually being 20-80 ℃ to be separated into cold high-pressure oil and cold high-pressure gas, and the following aims are achieved because a large amount of high-boiling-point components enter the hot high-pressure oil liquid: the cold high-fraction oil becomes less dense or less viscous or easily separated from water. The high-pressure separation process of the hydrogenation reaction effluent is provided with the hot high-pressure separator, and the high-pressure separation process also has the advantage of reducing heat loss because the hot high-pressure separation oil liquid can avoid the cooling process of using an air cooler or a water cooler for hot high-pressure separation gas. Meanwhile, part of the hot high-oil-content liquid can be returned to the upstream hydrogenation reaction process for recycling, so as to improve the overall raw material property of the hydrogenation reaction process receiving the circulating oil, or the circulating oil is subjected to circulating hydrogenation.

Before the hydrogenation reaction effluent or hot high-pressure gas enters the cold high-pressure separation part, the temperature is usually reduced (generally, heat exchange with the reaction part feed) to about 220-100 ℃ (the temperature should be higher than the crystallization temperature of the ammonium hydrosulfide in the gas phase of the hydrogenation reaction effluent), then washing water is usually injected into the reaction effluent to form a hydrogenation reaction effluent after water injection, the washing water is used for absorbing ammonia and other impurities such as hydrogen chloride and the like which may be generated, and the water solution after absorbing the ammonia necessarily absorbs the hydrogen sulfide. In the cold high-pressure separation part, the effluent of the hydrogenation reaction after water injection is separated into: a cold high-molecular gas mainly composed of hydrogen in volume, a cold high-molecular oil mainly composed of conventional liquid hydrocarbon and dissolved hydrogen, and a cold high-molecular water mainly composed of water and dissolved with ammonia and hydrogen sulfide. The cold high-moisture water generally contains 0.5-15% (w), preferably 1-8% (w) of ammonia. One purpose of the washing water injection is to absorb ammonia and hydrogen sulfide in the hydrogenation reaction effluent, prevent the formation of ammonia hydrosulfide or ammonia polysulfide crystals from blocking the heat exchanger channels, and increase the pressure drop of the system. The injection amount of the washing water is determined according to the following principle: on the one hand, the washing water is divided into vapor phase water and liquid phase water after being injected into the hydrogenation reaction effluent, and the liquid phase water amount is required to be more than zero, and is preferably 30 percent or more of the total amount of the washing water; in yet another aspect, the wash water is used to absorb ammonia from the hydrogenation effluent, to prevent the high partial gas from having too high an ammonia concentration, and to reduce catalyst activity, and generally the lower the ammonia volume concentration of the high partial gas, the better, the lower the ammonia volume concentration of the high partial gas, the more typically no greater than 200ppm (v), and most preferably no greater than 50ppm (v). The operating pressure of the cold high-pressure separator is the difference between the pressure of the hydrogenation reaction part and the actual pressure drop, and the difference between the operating pressure of the cold high-pressure separator and the hydrogenation reaction pressure is not too low or too high, generally 0.35-3.2 MPa, and generally 0.5-1.5 MPa. The hydrogen volume concentration value of the cold high-molecular gas should not be too low (leading to a rise in the operating pressure of the plant), and should generally be not less than 70% (v), preferably not less than 80% (v), and most preferably not less than 85% (v). At least one part of the cold high-molecular gas, which is 85-100 percent as mentioned above, is returned to the hydrogenation reaction part for recycling so as to provide the hydrogen amount and the hydrogen concentration which are necessary for the hydrogenation reaction part; in order to increase the investment efficiency of the plant, it is necessary to ensure that the recycle hydrogen concentration does not fall below the aforementioned lower limit, for which reason, depending on the specific feedstock properties, reaction conditions, product distribution, a portion of the cold high-molecular gas may be removed to remove methane and ethane produced by the reaction. For discharged cold high-molecular gas, conventional membrane separation process or pressure swing adsorption process or oil washing process can be adopted to realize the separation of hydrogen and non-hydrogen gas components, and the recovered hydrogen is used as new hydrogen.

Fresh hydrogen is fed into the hydrogenation section to replenish hydrogen consumed during the hydrogenation reaction, and the higher the concentration of fresh hydrogen, the better, the more preferably the concentration of fresh hydrogen is not lower than 95% (v), and the more preferably not lower than 99% (v). All of the fresh hydrogen may be introduced into any of the hydrogenation sections, preferably into the prehydrogenation process R1.

Claims

1. The combined method of the upflow hydrogenation reaction processes of different hydrocarbon materials is characterized by comprising the following steps:

2. The method of claim 1, wherein:

3. The method of claim 1, wherein:

4. A method according to claim 1 or 2 or 3, characterized in that:

the upflow hydrogenation reaction process with the hydrogen used in series takes the hydrogen flowing direction as the positive direction, and the working mode is selected from 1 or the combination of several of the following:

5. A method according to claim 1 or 2 or 3, characterized in that:

at least 1 upflow hydrogenation reaction process in the upflow hydrogenation reaction processes used in series of the hydrogen receives the material containing the hydrogen donor.

6. A method according to claim 1 or 2 or 3, characterized in that:

in the upflow type hydrogenation reaction processes in which the hydrogen is used in series, at least 1 upflow type hydrogenation reaction process receives a material containing a hydrogen donor with a conventional boiling point of 230-400 ℃.

7. A method according to claim 1 or 2 or 3, characterized in that:

the different hydrocarbon feeds are selected from the group consisting of coal, mixtures of coal and oil, high aromatic heavy oil, low aromatic heavy oil, which refers to hydrocarbon oils comprising hydrocarbon components having a conventional boiling point above 530 ℃.

8. A method according to claim 1 or 2 or 3, characterized in that:

the different hydrocarbon feeds are selected from the group consisting of coal, mixtures of coal and oil, high aromatic heavy oil, low aromatic heavy oil, which refers to hydrocarbon oils comprising hydrocarbon components having a conventional boiling point above 550 ℃.

9. A method according to claim 1 or 2 or 3, characterized in that:

the different hydrocarbon materials are selected from coal, a mixture of coal and oil, heavy oil with high aromaticity, heavy oil with low aromaticity, the heavy oil refers to hydrocarbon oil containing hydrocarbon components with normal boiling point higher than 450 ℃;

10. A method according to claim 1 or 2 or 3, characterized in that:

11. A method according to claim 1 or 2 or 3, characterized in that:

the upflow type hydrogenation reaction process with the hydrogen used in series takes the hydrogen flowing direction as the positive direction, the reaction product of the downstream-most upflow type hydrogenation reaction process is separated to obtain hydrogen-rich gas RHPV, and at least part of the hydrogen-rich gas RHPV returns to the upstream upflow type hydrogenation reaction process for recycling.

12. A method according to claim 1 or 2 or 3, characterized in that:

the upflow type hydrogenation reaction process with the hydrogen used in series takes the hydrogen flowing direction as the positive direction, the reaction product of the downstream-most upflow type hydrogenation reaction process is separated to obtain hydrogen-rich gas RHPV, and at least part of the hydrogen-rich gas RHPV returns to the upstream-most upflow type hydrogenation reaction process for recycling.

13. A method according to claim 1 or 2 or 3, characterized in that:

the upflow type hydrogenation reaction process with the hydrogen used in series takes the hydrogen flowing direction as the positive direction, the reaction product of the downstream-most upflow type hydrogenation reaction process is separated to obtain hydrogen-rich gas RHPV, at least one part of the hydrogen-rich gas RHPV returns to the upstream upflow type hydrogenation reaction process for recycling, and the volume concentration of the hydrogen is more than 70%.

14. A method according to claim 1 or 2 or 3, characterized in that:

the upflow type hydrogenation reaction process with the hydrogen used in series takes the hydrogen flowing direction as the positive direction, the reaction product of the downstream-most upflow type hydrogenation reaction process is separated to obtain hydrogen-rich gas RHPV, at least one part of the hydrogen-rich gas RHPV returns to the upstream upflow type hydrogenation reaction process for recycling, and the volume concentration of the hydrogen is more than 85%.

15. A method according to claim 1 or 2 or 3, characterized in that:

the operating temperature of the reaction product of the existing upflow hydrogenation reaction process in the thermal high-pressure separation process is 200-480 ℃, and the operating pressure is 6-30 MPa.

16. A method according to claim 1 or 2 or 3, characterized in that:

the operating temperature of the reaction product of the upflow hydrogenation reaction process in the thermal high-pressure separation process is 300-460 ℃, and the operating pressure is 10-25 MPa.

17. A method according to claim 1 or 2 or 3, characterized in that:

the operating temperature of the reaction product of the upflow hydrogenation reaction process in the thermal high-pressure separation process is 350-430 ℃, and the operating pressure is 13-20 MPa.

18. A method according to claim 1 or 2 or 3, characterized in that:

the upflow reactor used in the existing upflow hydrogenation reaction process works in a mode selected from 1 or a combination of several of the following modes:

① suspension bed hydrogenation reactor;

③ combined hydrogenation reactor of suspension bed and boiling bed;

④ micro-expanded bed.