CN110643391B

CN110643391B - Method for converting coal tar/heavy oil into liquefied gas, naphtha and diesel oil

Info

Publication number: CN110643391B
Application number: CN201910822631.4A
Authority: CN
Inventors: 郭强; 田磊; 杨勇; 李永旺; 万会军; 杨强; 魏迎春; 白亮; 余晓忠
Original assignee: SYNEFUELS CHINA Inc; Synfuels China Technology Co Ltd
Current assignee: Zhongke Synthetic Oil Engineering Co., Ltd; Zhongke Synthetic Oil Technology Co Ltd
Priority date: 2019-09-02
Filing date: 2019-09-02
Publication date: 2021-10-01
Anticipated expiration: 2039-09-02
Also published as: CN110643391A

Abstract

The invention relates to a processing method of coal tar/heavy oil, in particular to a method for converting coal tar/heavy oil into liquefied gas, naphtha and diesel oil. Mixing coal tar/heavy oil (and/or tail oil at the bottom of a refined fractionating tower) with a hydrocracking catalyst and hydrogen, allowing the mixture to enter a suspension bed reaction part for hydrocracking reaction, and separating and fractionating to obtain unrefined naphtha, unrefined diesel oil, wax oil and a small amount of gas-phase products; and carrying out hydrofining and optional secondary cracking reaction on the obtained unrefined naphtha, unrefined diesel oil and wax oil, and separating, fractionating and absorbing to stably cut liquefied gas, naphtha and diesel oil products. The method of the invention improves the yield of light oil products, prolongs the operation time of sustainable production of equipment, reduces the equipment stop-check times, and provides a processing method for improving the economical efficiency for tar/heavy oil with lower added value.

Description

Method for converting coal tar/heavy oil into liquefied gas, naphtha and diesel oil

Technical Field

The invention relates to a method for converting coal tar/heavy oil into liquefied gas, naphtha and diesel.

Background

In recent years, due to atmospheric pollution, haze is distributed throughout the country, the national requirement for clean energy is more urgent, the quantity of petroleum resources which can be collected is gradually reduced along with the continuous rising of the demand of petroleum products, crude oil is gradually in the development trend of deterioration and heaviness, and the processing and clean utilization of heavy oil also bring more challenges to the oil refining industry. China is a country with more coal and less oil, and if coal can be converted into liquid fuel in a large scale by a high-efficiency clean process method, the shortage of oil supply can be effectively relieved, and the sustainable and stable development of economy is promoted. With the rapid growth of the international and domestic steel industry, the coking industry shows a growing trend, the yield of coal tar is larger and larger, and the clean processing and the effective utilization of the coal tar become the main topics discussed by relevant students.

The coal tar has the same characteristics as heavy oil, such as high boiling point, high density, more heteroatoms, high oxygen atom, high carbon residue, low H/C ratio, high metal content, high asphaltene, high colloid and the like, has high processing difficulty, and needs to be subjected to a pretreatment working section before entering a fixed bed hydrogenation device based on the specificity of the properties of the oil; the pretreated coal tar is converted into light oil with low density and low viscosity by selecting a proper hydrofining and hydrocracking catalyst and adopting a proper process flow to lighten the whole fraction coal tar/heavy oil fraction.

CN101633848B discloses a medium and low temperature coal tar deep processing method, which comprises the steps of fractionating coal tar to obtain light fraction, phenol oil, heavy fraction and tail-removed pitch; the phenol oil is dephenolized to obtain a phenol product and dephenolized oil, and the dephenolized oil, the light fraction and the heavy fraction are subjected to hydrofining and hydrocracking reaction together to obtain dry gas, liquefied gas, hydrogenated naphtha and hydrogenated diesel oil. The process is only suitable for medium and low temperature coal tar with low distillation range, is not suitable for high temperature coal tar and heavy oil, and simultaneously the tail-removed asphalt is not fully utilized, so that the yield of light oil is low.

CN1351130A discloses a method for producing diesel oil by coal tar hydrogenation. The method mainly comprises the steps of fractionating coal tar firstly, wherein the obtained heavy fraction is not used as a raw material for hydrogenation treatment, but naphtha and diesel fraction in the coal tar are subjected to hydrogenation treatment, and the fraction with the distillation range larger than that of the diesel in the coal tar is not treated and utilized, so that the yield of the light oil is low.

CN104946306A discloses a combined method of suspension bed hydrocracking and fixed bed hydro-upgrading of coal tar whole fraction. The method comprises the steps of fractionating raw materials, sending obtained heavy distillate oil to a suspension bed for hydrocracking, and carrying out hydrofining reaction on the obtained light distillate oil and a suspension bed hydrocracking product in a fixed bed for hydrogenation modification. The process does not reprocess the vacuum distillate oil, the yield of the light oil is lower, the operating pressure of the hydrocracking part of the suspension bed is high, the investment is higher, and the consumption is large.

In view of the foregoing, there is a need for an improved coal tar/heavy oil treatment process.

Disclosure of Invention

In view of the above problems, it is an object of the present invention to provide a method for economically and efficiently converting coal tar/heavy oil into liquefied gas, naphtha and diesel.

Accordingly, the present invention provides a process for converting coal tar or heavy oil into liquefied gas, naphtha and diesel, the process comprising:

(1) hydrocracking: mixing coal tar or heavy oil with a hydrocracking catalyst and hydrogen, entering a suspension bed reaction part for hydrocracking reaction to obtain a gas-phase product and a liquid-phase product, and respectively entering the gas-phase product and the liquid-phase product into a suspension bed separation part for separation;

(2) and (3) fractionation after cracking: the liquid phase obtained by separation of the separation part of the suspension bed enters a suspension bed normal pressure fractionating tower for fractionation to obtain unrefined naphtha, unrefined diesel oil and tower bottom oil of the suspension bed normal pressure fractionating tower; the bottom oil of the suspension bed atmospheric fractionating tower enters a suspension bed reduced pressure fractionating tower for fractionation to obtain wax oil and bottom-reduced residues;

(3) and (3) hydrofining: mixing the unrefined naphtha, the unrefined diesel oil and the wax oil with hydrogen, and allowing the mixture to enter a fixed bed reaction part for hydrofining reaction, so that a hydrofining reaction product enters a fixed bed separation part for separation after heat exchange; and

(4) fractionation and absorption stabilization after refining: the liquid phase obtained from the fixed bed separation part enters a refining fractionation part for fractionation to obtain crude naphtha, diesel oil and tower bottom tail oil of the refining fractionation part; and (3) enabling the gas phase and the crude naphtha obtained by separation in the refining and fractionating part and the gas phase obtained by separation in the step (2) in the suspension bed normal pressure fractionating tower to enter an absorption stabilizing part, and obtaining gas phase dry gas, liquid phase liquefied gas and naphtha.

The liquefied gas, naphtha and diesel oil obtained by the method all meet the national standard.

The invention has the beneficial technical effects that:

(1) the coal tar/heavy oil adopts a full-fraction suspension bed hydrogenation technology, heavy fraction does not need to be cut before entering a suspension bed reactor, and the yield of the light oil as a product is high;

(2) the hydrocracking catalyst used in the suspension bed reaction part is a non-noble metal catalyst, so that the regeneration is not needed, the consumption is low, and the operation cost is low;

(3) the suspension bed has mild reaction conditions, low investment and low operating cost;

(4) the raw material at the fixed bed part is pretreated by a suspension bed, the fraction is lightened, the impurity content is low, the operation condition is mild, and the operation period is long.

Drawings

Fig. 1 is a process flow diagram of one embodiment of a process for converting coal tar/heavy oil into liquefied gas, naphtha and diesel according to the present invention.

1-a suspended bed reactor; 2-suspended bed thermal low pressure separator; 3-a suspension bed cold low pressure separator; 4-a suspended bed temperature high pressure separator; 5-suspension bed cold high pressure separator; 6-a suspension bed recycle hydrogen compressor; 7-suspension bed atmospheric fractionating tower; 8-a suspension bed vacuum fractionating tower; 9-a desulfurization section; 10-a hydrogen recovery section; 11-a protectant reactor; 12-a finishing reactor; 13-fixed bed thermal high pressure separator; 14-fixed bed cold high pressure separator; 15-fixed bed thermal low pressure separator; 16-fixed bed cold low pressure separator; 17-fixed bed recycle hydrogen compressor; 18-a fresh hydrogen compressor; 19-a refining fractionator; 20-a stabilizer column; 21-absorption desorption tower

Fig. 2 is a process flow diagram of another embodiment of a process for converting coal tar/heavy oil into liquefied gas, naphtha and diesel according to the present invention.

1-a suspended bed reactor; 2-suspended bed thermal low pressure separator; 3-a suspension bed cold low pressure separator; 4-a suspended bed temperature high pressure separator; 5-suspension bed cold high pressure separator; 6-a suspension bed recycle hydrogen compressor; 7-suspension bed atmospheric fractionating tower; 8-a suspension bed vacuum fractionating tower; 9-a desulfurization section; 10-a hydrogen recovery section; 11-a protectant reactor; 12-a finishing reactor; 13-fixed bed thermal high pressure separator; 14-fixed bed cold high pressure separator; 15-fixed bed thermal low pressure separator; 16-fixed bed cold low pressure separator; 17-fixed bed recycle hydrogen compressor; 18-a fresh hydrogen compressor; 19-a refining fractionator; 20-a cracking reactor; 21-absorption desorption tower; 22-stabilizer tower

Fig. 3 is a process flow diagram of one embodiment of a method for converting coal tar/heavy oil into liquefied gas, naphtha and diesel according to the present invention.

1-a suspended bed reactor; 2-suspended bed thermal low pressure separator; 3-a suspension bed cold low pressure separator; 4-a suspended bed temperature high pressure separator; 5-suspension bed cold high pressure separator; 6-a suspension bed recycle hydrogen compressor; 7-suspension bed atmospheric fractionating tower; 8-a suspension bed vacuum fractionating tower; 9-a desulfurization section; 10-a hydrogen recovery section; 11-a protectant reactor; 12-a finishing reactor; 13-fixed bed thermal high pressure separator; 14-fixed bed cold high pressure separator; 15-fixed bed thermal low pressure separator; 16-fixed bed cold low pressure separator; 17-fixed bed recycle hydrogen compressor; 18-a fresh hydrogen compressor; 19-a refining fractionator; 20-a stabilizer column; 21-absorption desorption tower; 22-suspended bed thermal high pressure separator; 23-suspension bed temperature low-pressure separator

Detailed Description

The invention provides a method for converting coal tar/heavy oil into liquefied gas, naphtha and diesel oil, which comprises the following steps:

(2) and (3) fractionation after cracking: the liquid phase obtained by separation of the separation part of the suspension bed enters a suspension bed normal pressure fractionating tower for fractionation to obtain unrefined naphtha, unrefined diesel oil and tower bottom oil of the suspension bed normal pressure fractionating tower; the bottom oil of the suspension bed atmospheric fractionating tower enters a suspension bed reduced pressure fractionating tower for fractionation, and wax oil and bottom-reduced residue are obtained after separation by the suspension bed reduced pressure fractionating tower;

(3) and (3) hydrofining: mixing the unrefined naphtha, the unrefined diesel oil and the wax oil with hydrogen, entering a fixed bed reaction part containing a hydrofining catalyst for hydrofining reaction, and entering a fixed bed separation part for separation after heat exchange of a hydrofining reaction product; and

(4) fractionation and absorption stabilization after refining: the liquid phase obtained from the fixed bed separation part enters a refining fractionation part for fractionation to obtain crude naphtha, diesel oil and tower bottom tail oil of the refining fractionation part; the gas phase separated by the suspension bed atmospheric fractionating tower and the gas phase and the crude naphtha separated by the refining fractionating part enter an absorption stabilizing part to obtain gas phase dry gas and liquid phase liquefied gas and naphtha.

In one embodiment, the suspended bed reaction section can comprise a suspended bed reactor. The suspension bed hydrogenation process is characterized by mixing raw oil, catalyst and hydrogen, and making cracking hydrogenation reaction under the condition of high-temp. and high-pressure. The suspension bed hydrogenation process has complex material properties, various reactions, mainly thermal cracking reaction and harsh reaction conditions. The suspension bed hydrogenation reactor adopts the catalyst and the liquid phase to form slurry phase which is the same as the feeding, the utilization efficiency of the catalyst is high, the reactor has the advantages of high conversion rate, uniform temperature distribution and the like, but the fine particle catalyst cannot be retained in the reactor and is easy to discharge together with the material, and the catalyst is difficult to recover. In the invention, because the catalyst used in the suspension bed reactor is a non-noble metal catalyst, part or all of the catalyst can be recycled, or the catalyst can not be recycled, so that the advantages of the suspension bed reactor can be fully exerted.

In one embodiment, the suspension bed separation section may be selected from any one of the following combinations:

(1) a suspended bed temperature high-pressure separator, a suspended bed cold high-pressure separator, a suspended bed hot low-pressure separator and a suspended bed cold low-pressure separator;

(2) a suspended bed hot high-pressure separator, a suspended bed temperature high-pressure separator, a suspended bed cold high-pressure separator, a suspended bed hot low-pressure separator, a suspended bed temperature low-pressure separator and a suspended bed cold low-pressure separator; or

(3) A suspended bed temperature high-pressure separator, a suspended bed cold high-pressure separator, a suspended bed hot low-pressure separator, a suspended bed temperature low-pressure separator and a suspended bed cold low-pressure separator.

In one embodiment, the fixed bed reaction section may be selected from any one of the following combinations: (1) a protective agent reactor and a refining reactor; or (2) a finishing reactor comprising a bed of protective agent. The protective agent reactor and the refining reactor are fixed bed reactors, and the protective agent reactor and the refining reactor can be independently used in parallel or used in series. Fixed bed reactors are typically used in the hydrogenation of light or distillate oils to produce naphtha and diesel products. Under the condition of hydrofining, unrefined naphtha, unrefined diesel oil, wax oil and hydrogen are mixed and enter a reactor, various impurities in the raw materials are removed through hydrofining by a hydrofining catalyst bed layer, refined reaction products are discharged out of the fixed bed reactor from the bottom of the fixed bed reactor, and the naphtha, the diesel oil and the liquefied gas are obtained after subsequent separation and fractionation.

In one embodiment, the fixed bed separation section may be selected from any one of the following combinations: (1) a fixed bed hot high-pressure separator, a fixed bed cold high-pressure separator, a fixed bed hot low-pressure separator and a fixed bed cold low-pressure separator; or (2) a fixed bed cold high pressure separator + a fixed bed cold low pressure separator.

In one embodiment, the suspension bed fractionation section may be selected from any one of the following combinations: (1) a suspension bed atmospheric fractionating tower and a suspension bed vacuum fractionating tower; or (2) a suspension bed stripping tower, a suspension bed normal-pressure fractionating tower and a suspension bed reduced-pressure fractionating tower. In addition, it is preferred that the suspended bed atmospheric fractionation column and/or the suspended bed vacuum fractionation column is provided with one or more side strippers.

In one embodiment, the finishing fractionation section may be selected from any one of the following combinations: (1) a stripping tower and an atmospheric fractionating tower; (2) an atmospheric fractionating tower and a vacuum fractionating tower; (3) a stripping tower, an atmospheric fractionating tower and a reduced pressure fractionating tower; or (4) an atmospheric fractionating column. In addition, it is preferred that the atmospheric fractionation column and/or the vacuum fractionation column is provided with one or more side strippers.

In one embodiment, the absorption stabilizing moiety may be selected from any one of the following combinations: (1) an absorption desorption tower and a stabilizing tower; (2) an absorption tower, a desorption tower and a stabilizing tower; (3) a naphtha stabilizer column; or 4) a debutanizer + deethanizer.

In one embodiment, the coal tar/heavy oil as the raw material is not particularly limited, and includes, for example, one or more of low temperature coal tar, medium temperature coal tar and high temperature coal tar, and the heavy oil is heavy oil remaining after extracting gasoline and diesel oil from crude oil.

In one embodiment, the catalyst used in the suspended bed reactor is a non-noble metal catalyst. The catalyst is added on line along with raw oil in the form of catalyst-containing tar slurry, and the waste catalyst is discharged along with the bottom-reducing residue of the suspension bed vacuum fractionating tower. The non-noble metal catalyst may be selected from base metal catalysts or combinations thereof, such as Cr-based catalysts, Fe-based catalysts, Ni-based catalysts, Mo-based catalysts, W-based catalysts or Co-based compounds. The present example preferably uses a spent Fe-based catalyst from a fischer-tropsch synthesis process. The carrier composition can be one or any combination of more of silica, alumina, zirconia, titania and molecular sieve. Wherein a carrier prepared by mixing two or more components of the above-mentioned carrier composition is called a composite carrier; the alumina carrier can be pretreated by acid or alkali, such as pretreated alumina carrier by phosphoric acid or sulfuric acid or organic acid, which is called modified alumina carrier; the molecular sieve can be beta molecular sieve, MCM series molecular sieve, Y series molecular sieve, ZSM series molecular sieve or SAP0 series molecular sieve, and can also be one or a combination of several molecular sieves. The catalyst is a non-noble metal catalyst, so that the consumption is low, regeneration is not needed, and the operation cost is low. After the suspension bed reaction, separation and fractionation, the catalyst is discharged out of the suspension bed decompression tower along with the bottom-reducing residue. The bottoms reduction residue can be partially or completely recycled to the suspension bed reaction section.

In one embodiment, the catalyst used in the fixed bed reaction section is typically supported on alumina or silica-containing alumina, and a non-noble metal from group VIB (e.g., Mo, W, etc.) and/or VIII (e.g., Ni, Fe, Co, etc.) is used as an active component. The catalyst is first used as a conventional, or as a conventional hydrotreating catalyst. The catalyst is a hydrotreating catalyst well known to those skilled in the art. The catalyst may be prepared according to methods well known to those skilled in the art and may also be selected from a variety of commercially available catalysts such as various hydrofinishing, hydrotreating catalysts developed by the comforting petrochemical institute, and the like, including FZC-100, FZC-105, FZC-106, FZC-204, FZC-33, FZC-14 catalysts, and the like. Such as ZKM and ZKH series catalysts developed by Zhongke synthetic oil technology Co.

In one embodiment, the bottom-reduced residue obtained from the suspension bed vacuum fractionation tower is directly discharged from the suspension bed vacuum fractionation tower, and may be partially or completely recycled to the suspension bed reaction section.

In one embodiment, the bottoms of the finishing fractionation section may be recycled to the slurry reaction section for hydrocracking reactions. In another embodiment, the bottom tail oil of the refined fractionation section can be mixed with hydrogen and recycled to the cracking reactor for secondary cracking reaction. The cracking reactor is a fixed bed reactor, and the recycle hydrogen part, the separation part and the fractionation part of the cracking reactor can be shared with the fixed bed reactor, or can be partially or completely independently arranged.

The invention also provides a method for converting coal tar/heavy oil into liquefied gas, naphtha and diesel, which comprises the following steps:

(1) hydrocracking: mixing coal tar/heavy oil and a hydrocracking catalyst with hydrogen, and feeding the mixture into a suspension bed reactor for hydrocracking reaction; decompressing the liquid phase product of the suspension bed reactor, and separating in a suspension bed hot low-pressure separator; the liquid phase separated by the suspension bed thermal low-pressure separator enters a suspension bed normal-pressure fractionating tower for fractionation; the gas phase separated by the suspension bed hot low pressure separator enters a suspension bed cold low pressure separator for separation; after heat exchange and temperature reduction are carried out on gas-phase products of the suspension bed reactor, the gas-phase products enter a suspension bed temperature high-pressure separator for separation; cooling the gas phase separated by the suspension bed temperature high-pressure separator, and separating in a suspension bed cold high-pressure separator; the liquid phase separated by the suspension bed cold high-pressure separator enters the suspension bed cold low-pressure separator for separation;

(2) and (3) fractionation after cracking: the liquid phase separated by the suspension bed temperature high-pressure separator and the liquid phase separated by the suspension bed cold low-pressure separator enter a suspension bed normal-pressure fractionating tower for fractionation; separating by a suspension bed atmospheric fractionating tower to obtain unrefined naphtha, unrefined diesel oil and tower bottom oil of the suspension bed atmospheric fractionating tower; the bottom oil of the suspension bed atmospheric fractionating tower enters a suspension bed reduced pressure fractionating tower for fractionation, and wax oil and bottom-reduced residue are obtained after separation by the suspension bed reduced pressure fractionating tower;

(3) and (3) hydrofining: mixing unrefined naphtha, unrefined diesel oil, wax oil and hydrogen, and sequentially entering a protective agent reactor containing a hydrofining catalyst and a refining reactor to perform hydrofining reaction; transferring the hydrorefining reaction product after heat exchange into a fixed bed hot high-pressure separator for separation; the liquid phase separated by the fixed bed heat high-pressure separator enters the fixed bed heat low-pressure separator for separation; cooling the gas phase separated by the fixed bed hot high-pressure separator, and separating in a fixed bed cold high-pressure separator; cooling the gas phase separated by the fixed bed hot low pressure separator, and separating by a fixed bed cold low pressure separator; the liquid phase separated by the fixed bed cold high-pressure separator enters a fixed bed cold low-pressure separator for separation; and

(4) fractionation and absorption stabilization after refining: the liquid phase separated by the fixed bed hot low pressure separator and the liquid phase separated by the fixed bed cold low pressure separator enter a refining fractionating tower for fractionation; separating by a refining fractionating tower to obtain crude naphtha, diesel oil and tower bottom tail oil of the refining fractionating tower; the gas phase obtained by the separation of the suspension bed atmospheric fractionating tower and the gas phase obtained by the separation of the refining fractionating tower and the crude naphtha enter an absorption desorption tower; and (4) enabling the tower bottom product of the absorption desorption tower to enter a stabilizing tower, obtaining liquefied gas at the tower top of the stabilizing tower, and obtaining naphtha at the tower bottom of the stabilizing tower.

In one embodiment, the liquid phase product of the suspended bed reactor can enter the suspended bed thermal high-pressure reactor before entering the suspended bed thermal low-pressure separator, the gas phase separated by the suspended bed thermal high-pressure reactor enters the suspended bed thermal high-pressure separator, and the liquid phase separated by the suspended bed thermal high-pressure reactor enters the suspended bed thermal low-pressure separator; the liquid phase separated by the suspension bed temperature high-pressure separator can enter the suspension bed temperature low-pressure separator before entering the suspension bed normal-pressure fractionating tower, the gas phase separated by the suspension bed temperature low-pressure separator enters the suspension bed cold low-pressure separator, and the liquid phase separated by the suspension bed temperature low-pressure separator enters the suspension bed normal-pressure fractionating tower.

In one embodiment, part of the gas phase separated by the suspension bed cold high-pressure separator is mixed with fresh hydrogen after being pressurized by the suspension bed circulating hydrogen compressor and then returned to the suspension bed reactor. In addition, preferably, the mixed gas phase is divided into two streams, the first stream and the feeding material of the suspension bed reactor are mixed and then enter the suspension bed reactor, the second stream enters the bottom of the suspension bed reactor after being heated, and the ratio of the first stream to the second stream is preferably 1: 5-3: 5.

In one embodiment, the other part of the gas phase separated by the suspension bed cold high-pressure separator, the gas phase separated by the suspension bed cold low-pressure separator and the gas phase separated by the fixed bed cold low-pressure separator pass through the desulfurization section and then enter the hydrogen recovery device.

In one embodiment, the bottoms oil from the hydrofinished refinery fractionator is recycled to the suspended bed reactor for hydrocracking reactions. In another embodiment, the bottom tail oil of the hydrofined refined fractionating tower and hydrogen are mixed and circulated to the cracking reactor for secondary cracking reaction, and the cracking reaction product of the cracking reactor enters a fixed bed hot high-pressure separator for gas-liquid separation after heat exchange.

In one embodiment, the gas phase separated by the fixed bed cold high pressure separator is subjected to pressure increase by the fixed bed recycle hydrogen compressor and then is mixed with part of hydrogen generated by the fresh hydrogen compressor to be divided into two parts. One enters the protectant reactor and the other enters the cracking reactor.

In one embodiment, in the step (1), the reaction temperature of the suspension bed reactor is 250-550 ℃, the operation pressure is 2.0-10.0 MPaG, and the total hydrogen-oil ratio at the inlet of the suspension bed reactor is 200-3000; the average residence time of the liquid phase of the suspension bed reactor is 10-180 min, and the addition amount of the catalyst is 0.02-6 wt% of the raw oil. Preferably, the reaction temperature of the suspension bed reactor is 350-500 ℃, the operation pressure is 4.0-8.0 MPaG, and the total hydrogen-oil ratio at the inlet of the suspension bed reactor is 500-2000; the average residence time of the liquid phase of the suspension bed reactor is 60-120 min, and the adding amount of the catalyst is 0.5-4 wt% of the raw oil.

In one embodiment, in the step (1), the operation temperature of the suspended bed hot low-pressure separator is 200 to 550 ℃, and the operation pressure is 0.5 to 4.0 MPaG; preferably, the operation temperature of the suspension bed hot low-pressure separator is 300-500 ℃, and the operation pressure is 1.5-3.0 MPaG.

In one embodiment, in the step (1), the operation temperature of the suspended bed temperature high-pressure separator is 200 to 450 ℃, and the operation pressure is 2.0 to 10.0 MPaG; preferably, the operation temperature of the suspension bed temperature high-pressure separator is 300-400 ℃, and the operation pressure is 3.0-8.0 MPaG.

In one embodiment, in the step (1), the operating temperature of the suspension bed cold high-pressure separator is 20-100 ℃, and the operating pressure is 2.0-10.0 MPaG; preferably, the operating temperature of the suspension bed cold high-pressure separator is 30-60 ℃, and the operating pressure is 3.0-8.0 MPaG.

In one embodiment, in the step (1), the operating temperature of the suspension bed cold low-pressure separator is 20-100 ℃, and the operating pressure is 0.5-4.0 MPaG; preferably, the operating temperature of the suspension bed cold low-pressure separator is 30-60 ℃, and the operating pressure is 2.0-3.0 MPaG.

In one embodiment, in the step (2), the number of theoretical plates of the suspension bed atmospheric fractionating tower is 10-60, the tower top temperature is 60-200 ℃, the tower bottom temperature is 250-480 ℃, and the operation pressure is 0.01-1.0 MPaG; preferably, the number of theoretical plates of the suspension bed atmospheric fractionating tower is 20-40, the temperature at the top of the tower is 80-150 ℃, the temperature at the bottom of the tower is 300-400 ℃, and the operating pressure is 0.05-0.2 MPaG.

In one embodiment, in the step (2), the number of theoretical plates of the suspension bed vacuum fractionating tower is 10-60, the tower top temperature is 30-150 ℃, the tower bottom temperature is 200-450 ℃, and the operation pressure is 0-0.1 MPaG; preferably, the number of theoretical plates of the suspension bed vacuum fractionating tower is 20-50, the temperature at the top of the tower is 50-100 ℃, the temperature at the bottom of the tower is 300-400 ℃, and the operating pressure is-0.08-0.1 MPaG.

In one embodiment, in the step (3), the reaction temperature of the protective agent reactor is 150-450 ℃, the operation pressure is 6.0-23.0 MPaG, and the volume space velocity is 0.1-10 h^-1The volume ratio of hydrogen to oil at the inlet of the protective agent reactor is 200-3000; preferably, the reaction temperature of the protective agent reactor is 300-400 ℃, the operation pressure is 10.0-18.0 MPaG, and the volume space velocity is 1-5 h^-1And the volume ratio of hydrogen to oil at the inlet of the protective agent reactor is 500-2000.

In one embodiment, in the step (3), the temperature of the refining reactor is 250-500 ℃, the operation pressure is 6.0-23.0 MPaG, and the volume space velocity is 0.05-3 h^-1The volume ratio of hydrogen to oil at the inlet of the refining reactor is 200-3000; preferably, the reaction temperature of the refining reactor is 350-400 ℃, the operation pressure is 10.0-18.0 MPaG, and the volume space velocity is 0.1-1 h^-1And the volume ratio of hydrogen to oil at the inlet of the refining reactor is 500-2000.

In one embodiment, in step (3), the operating temperature of the fixed bed hot high pressure separator is 150 to 350 ℃ and the operating pressure is 6.0 to 23 MPaG; preferably, the operating temperature of the fixed bed hot high-pressure separator is 200-300 ℃, and the operating pressure is 10.0-18.0 MPaG.

In one embodiment, in step (3), the operating temperature of the fixed bed hot low pressure separator is 150 to 350 ℃ and the operating pressure is 0.5 to 4.0 MPaG; preferably, the operation temperature of the fixed bed hot low pressure separator is 200-300 ℃, and the operation pressure is 2.0-3.0 MPaG.

In one embodiment, in the step (3), the operation temperature of the fixed bed cold high-pressure separator is 20-100 ℃, and the operation pressure is 6.0-23 MPaG; preferably, the operation temperature of the fixed bed cold high-pressure separator is 40-60 ℃, and the operation pressure is 10.0-18.0 MPaG.

In one embodiment, in the step (3), the operation temperature of the fixed bed cold low pressure separator is 20 to 100 ℃, and the operation pressure is 0.5 to 4.0 MPaG; preferably, the operation temperature of the fixed bed cold low-pressure separator is 40-60 ℃, and the operation pressure is 2.0-3.0 MPaG.

In one embodiment, in the step (4), the number of theoretical plates of the refining fractionating tower is 10-60, the tower top temperature is 60-200 ℃, the tower bottom temperature is 150-450 ℃, and the operation pressure is 0.01-1.0 MPaG; preferably, the number of theoretical plates of the refining fractionating tower is 20-40, the tower top temperature is 80-150 ℃, the tower bottom temperature is 300-400 ℃, and the operation pressure is 0.05-0.2 MPaG.

In a preferred embodiment, the temperature of the cracking reactor is 250-500 ℃, the operating pressure is 6.0-23.0 MPaG, and the total volume space velocity is 0.05-6 h^-1The volume ratio of hydrogen to oil at the inlet of the cracking reactor is 200-3000; preferably, the reaction temperature of the cracking reactor is 300-400 ℃, the operation pressure is 10.0-18.0 MPaG, and the total volume space velocity is 0.2-3 h^-1And the volume ratio of hydrogen to oil at the inlet of the cracking reactor is 500-1500.

In one embodiment, in the step (4), the number of theoretical plates of the absorption desorption tower is 5-40, the tower top temperature is-40-120 ℃, the tower bottom temperature is-40-250 ℃, and the operation pressure is 0.5-6.0 MPaG; the number of theoretical plates of the stabilizer is 10-60, the temperature at the top of the tower is 30-150 ℃, the temperature at the bottom of the tower is 150-300 ℃, and the operating pressure is 0.3-2.0 MPaG. Preferably, the number of theoretical plates of the absorption desorption tower is 20-30, the temperature at the top of the tower is 0-50 ℃, the temperature at the bottom of the tower is 100-200 ℃, and the operating pressure is 1.0-2.0 MPaG; the number of theoretical plates of the stabilizer is 20-40, the temperature at the top of the tower is 40-60 ℃, the temperature at the bottom of the tower is 180-240 ℃, and the operating pressure is 0.8-1.5 MPaG.

The process of the present invention is further illustrated below with reference to FIG. 1:

(1) hydrocracking: mixing coal tar 101, tar slurry 102 containing a hydrocracking catalyst and tower bottom tail oil of a hydrofining fractionating tower with hydrogen, and allowing the mixture to enter a suspension bed reactor 1 for a hydrocracking reaction; the liquid phase product after the reaction in the suspended bed reactor 1 is decompressed and then enters a suspended bed thermal low-pressure separator 2 for separation; the liquid phase separated by the suspension bed heat low-pressure separator 2 enters a suspension bed normal-pressure fractionating tower 7 for fractionation; the gas phase separated by the suspension bed hot low pressure separator 2 enters a suspension bed cold low pressure separator 3 for separation; after heat exchange and cooling are carried out on gas-phase products after the reaction of the suspended bed reactor 1, the gas-phase products are separated by a suspended bed temperature high-pressure separator 4, and suspended bed temperature high-pressure oil obtained by the suspended bed temperature high-pressure separator 4 enters a suspended bed normal-pressure fractionating tower 7 for fractionation after being decompressed; after heat exchange and cooling of the suspension bed temperature high-pressure gas from the suspension bed temperature high-pressure separator 4, the gas enters a suspension bed cold high-pressure separator 5 for gas-liquid-water three-phase separation, and the gas phase at the top of the suspension bed cold high-pressure separator 5 is subjected to pressure increase by a suspension bed circulating hydrogen compressor 6, mixed with part of new hydrogen from a new hydrogen compressor 18 and then returned to the suspension bed reactor 1 for hydrocracking reaction; the water and the oil in the suspension bed cold high-pressure separator 5 are decompressed and then enter the suspension bed cold low-pressure separator 3 for separation; the gas phase of the suspension bed cold low-pressure separator 3 enters a desulfurization part 9 for desulfurization, and the suspension bed cold low-pressure oil of the suspension bed cold low-pressure separator 3 enters a suspension bed normal-pressure fractionating tower 7 for fractionation;

(2) and (3) fractionation after cracking: the suspended bed hot low-fraction oil from the suspended bed hot low-pressure separator 2, the suspended bed hot high-fraction oil from the suspended bed hot high-pressure separator 4 and the suspended bed cold low-fraction oil from the suspended bed cold low-pressure separator 3 enter a suspended bed atmospheric fractionating tower 7 for fractionation, rich gas obtained by fractionation of the suspended bed atmospheric fractionating tower 7 enters an absorption desorption tower 21, extracted unrefined naphtha and unrefined diesel oil enter a protective agent reactor 11 for hydrofining reaction, the tower bottom oil of the suspended bed atmospheric fractionating tower 7 enters a suspended bed vacuum fractionating tower 8, reduced pressure side-line wax oil of the suspended bed vacuum fractionating tower 8 enters the protective agent reactor 11 for hydrofining reaction, and the reduced bottom residue of the suspended bed vacuum fractionating tower 8 is sent out of the device as a coking raw material or circulated to the suspended bed reactor 1 for hydrocracking reaction;

(3) and (3) hydrofining: unrefined naphtha, unrefined diesel oil, reduced-pressure side wax oil and hydrogen obtained from the suspension bed fractionation part are mixed and sequentially enter a protective agent reactor 11 and a refining reactor 12 for hydrofining reaction, the hydrofining reaction is carried out under a hydrofining catalyst and proper temperature and pressure, and materials such as desulfurization, denitrification, deoxidation, demetalization and the like and olefin saturation are carried out; the hydrorefining reaction product is separated by a fixed bed hot high-pressure separator 13, the obtained fixed bed hot high-pressure oil is decompressed and then enters a fixed bed hot low-pressure separator 15 for separation, and the obtained fixed bed hot high-pressure oil enters a fixed bed cold high-pressure separator 14 for gas, liquid and water three-phase separation after heat exchange and cooling; the top gas phase of the fixed bed cold high-pressure separator 14 is subjected to pressure increase by a fixed bed circulating hydrogen compressor 17, mixed with part of fresh hydrogen from a fresh hydrogen compressor 18 and then mixed with a raw material for a hydrofining reaction to form hydrogen mixed oil, and the hydrogen mixed oil sequentially enters the protective agent reactor 11 and the refining reactor 12 to carry out the hydrofining reaction; the liquid phase in the fixed bed thermal low-pressure separator 15 is decompressed and then enters a refining fractionating tower 19 for fractionation; cooling the gas phase in the fixed bed hot low-pressure separator 15, then separating the gas phase in the fixed bed cold low-pressure separator 16, depressurizing the water and oil in the fixed bed cold high-pressure separator 14, then separating the gas phase in the fixed bed cold low-pressure separator 16, feeding the gas phase in the fixed bed cold low-pressure separator 16 into the desulfurization part 9 for desulfurization, feeding the desulfurized gas phase into the hydrogen recovery device 10, and feeding the fixed bed cold low-pressure oil from the fixed bed cold low-pressure separator 16 into the refining fractionating tower 19 for fractionation;

(4) fractionation and absorption stabilization after refining: the fixed bed hot low fraction oil from the fixed bed hot low pressure separator 15 and the fixed bed cold low fraction oil from the fixed bed cold low pressure separator 16 are subjected to heat exchange and then mixed to enter a refining fractionating tower 19 for fractionation, the top gas phase and the crude naphtha of the refining fractionating tower 19 enter an absorption desorption tower 21, the diesel oil generated by the refining fractionating tower 19 is a diesel oil product, and the bottom tail oil of the refining fractionating tower 19 is circulated to the suspended bed reactor 1 for hydrocracking reaction;

the tower top rich gas of the suspension bed atmospheric fractionating tower 7 and the tower top gas of the refined fractionating tower 19 enter an absorption desorption tower 21 after being boosted, and the tower top crude naphtha oil gas of the refined fractionating tower 19 enters the absorption desorption tower 21 after being boosted; the gas phase at the top of the absorption desorption tower 21 is used as a dry gas output device, the deethanized oil at the bottom of the tower enters a stabilizing tower 20, a liquefied gas product is obtained from the top of the stabilizing tower 20, and a naphtha product and a circulating absorbent are obtained from the bottom of the stabilizing tower 20;

(5) desulfurization and hydrogen recovery: part of the suspension bed cold high-pressure gas from the suspension bed cold high-pressure separator 5, the suspension bed cold low-pressure gas from the suspension bed cold low-pressure separator 3 and the fixed bed cold low-pressure gas from the fixed bed cold low-pressure separator 16 enter a desulfurization part 9 for desulfurization, and the desulfurized gas phase enters a hydrogen recovery device 10; the tail gas enters an absorption desorption tower 21 to recover liquefied gas components or is used as fuel gas to be sent out of the device through the recovered hydrogen obtained by the hydrogen recovery device 10.

The process of the present invention is further illustrated below with reference to FIG. 2:

(1) hydrocracking: mixing coal tar 101, tar slurry 102 containing a hydrocracking catalyst and hydrogen, and feeding the mixture into a suspension bed reactor 1 for a hydrocracking reaction; the liquid phase product after the reaction in the suspended bed reactor 1 is decompressed and then enters a suspended bed thermal low-pressure separator 2 for separation; the liquid phase separated by the suspension bed heat low-pressure separator 2 enters a suspension bed normal-pressure fractionating tower 7 for fractionation; the gas phase separated by the suspension bed hot low pressure separator 2 enters a suspension bed cold low pressure separator 3 for separation; after heat exchange and cooling are carried out on gas-phase products after the reaction of the suspended bed reactor 1, the gas-phase products are separated by a suspended bed temperature high-pressure separator 4, and the obtained suspended bed temperature high-pressure oil is decompressed and then is fractionated in a suspended bed normal-pressure fractionating tower 7; after heat exchange and cooling of the suspension bed temperature high-pressure gas from the suspension bed temperature high-pressure separator 4, the gas enters a suspension bed cold high-pressure separator 5 for gas-liquid-water three-phase separation, and the gas phase at the top of the suspension bed cold high-pressure separator 5 is mixed with part of new hydrogen from a new hydrogen compressor 18 after being boosted by a suspension bed circulating hydrogen compressor 6 and then returns to the suspension bed reactor 1 for hydrocracking reaction; the water and the oil in the suspension bed cold high-pressure separator 5 are decompressed and then enter the suspension bed cold low-pressure separator 3 for separation; the gas phase of the suspension bed cold low-pressure separator 3 enters a desulfurization part 9 for desulfurization, and the suspension bed cold low-pressure oil of the suspension bed cold low-pressure separator 3 is sent to a suspension bed normal-pressure fractionating tower 7 for fractionation;

(2) and (3) fractionation after cracking: the suspended bed hot low-fraction oil from the suspended bed hot low-pressure separator 2, the suspended bed temperature high-fraction oil from the suspended bed temperature high-pressure separator 4 and the suspended bed cold low-fraction oil from the suspended bed cold low-pressure separator 3 enter a suspended bed normal-pressure fractionating tower 7 for fractionation; rich gas obtained by separation in the suspension bed atmospheric fractionating tower 7 enters an absorption desorption tower 21, extracted unrefined naphtha and unrefined diesel oil enter a protective agent reactor 11 for hydrofining reaction, tower bottom oil of the suspension bed atmospheric fractionating tower 7 enters a suspension bed vacuum fractionating tower 8 for fractionation, vacuum side wax oil of the suspension bed vacuum fractionating tower 8 enters the protective agent reactor 11 for hydrofining reaction, and bottom-reduced residue of the suspension bed vacuum fractionating tower 8 is taken as a coking raw material and sent out of a device or circulated to the suspension bed reactor 1 for hydrocracking reaction;

(3) and (3) hydrofining: unrefined naphtha, unrefined diesel oil, reduced-pressure side wax oil and hydrogen are mixed and sequentially enter a protective agent reactor 11 and a refining reactor 12 for hydrofining reaction, hydrofining reaction is carried out under a hydrofining catalyst and proper temperature and pressure, and materials such as desulfurization, denitrification, deoxidation, demetalization and the like and olefin saturation are carried out; the hydrorefining reaction products and the secondary cracking reaction products (see below) are separated by a fixed bed hot high pressure separator 13; the fixed bed hot high-pressure oil separated by the fixed bed hot high-pressure separator 13 is decompressed and then enters the fixed bed hot low-pressure separator 15 for separation; the fixed bed hot high-pressure gas obtained by separation in the fixed bed hot high-pressure separator 13 enters a fixed bed cold high-pressure separator 14 for gas, liquid and water three-phase separation after heat exchange and cooling; the top gas phase of the fixed bed cold high-pressure separator 14 is subjected to pressure rise by a fixed bed circulating hydrogen compressor 17 and then is mixed with part of fresh hydrogen from a fresh hydrogen compressor 18 to be divided into two streams, one stream is mixed with raw materials for refining reaction to form hydrogen mixed oil, the hydrogen mixed oil sequentially enters a protective agent reactor 11 and a refining reactor 12 for hydrogenation refining reaction, and the other stream enters a cracking reactor 20 for secondary cracking reaction;

the liquid phase in the fixed bed thermal low-pressure separator 15 is decompressed and then enters a refining fractionating tower 19 for fractionation; the gas phase of the fixed bed hot low pressure separator 15 is cooled and then enters a fixed bed cold low pressure separator 16 for separation; the water and oil in the fixed bed cold high-pressure separator 14 are decompressed and then enter the fixed bed cold low-pressure separator 16 for separation; the gas phase of the fixed bed cold low pressure separator 16 enters the desulfurization part 9 for desulfurization, and the desulfurized gas phase enters the hydrogen recovery device 10; the fixed bed cold low-pressure oil of the fixed bed cold low-pressure separator 16 enters a refining fractionating tower 19 for fractionation;

(4) and (3) refining and then fractionating: the fixed bed hot low-fraction oil from the fixed bed hot low-pressure separator 15 and the fixed bed cold low-fraction oil from the fixed bed cold low-pressure separator 16 are mixed after heat exchange and enter a refining fractionating tower 19 for fractionation; the gas phase at the top of the refining fractionating tower 19 and the crude naphtha enter an absorption desorption tower 21, the diesel oil generated by the refining fractionating tower 19 is a diesel oil product, and the bottom oil of the refining fractionating tower 19 is circulated to a cracking reactor 20 for secondary cracking reaction;

(5) cracking reaction: mixing the bottom oil of the refining fractionating tower 19 with hydrogen, entering a cracking reactor 20 for secondary cracking reaction, and entering a fixed bed thermal high-pressure separator 13 for separation of secondary cracking reaction products and hydrofining reaction products;

(6) and (3) absorption stabilization: the tower top rich gas of the suspension bed atmospheric fractionating tower 7 and the tower top gas of the refining fractionating tower 19 enter an absorption desorption tower 21 after being boosted; the crude naphtha oil gas at the top of the refining fractionating tower 19 enters an absorption desorption tower 21 after being boosted, the gas phase at the top of the absorption desorption tower 21 is taken as a dry gas outlet device, and the deethanized oil at the bottom of the tower is sent to a stabilizing tower 22; obtaining a liquefied gas product from the top of the stabilizer tower 22, and obtaining a naphtha product and a circulating absorbent from the bottom of the stabilizer tower;

(7) desulfurization and hydrogen recovery: part of the suspension bed cold high-pressure gas from the suspension bed cold high-pressure separator 5, the suspension bed cold low-pressure gas from the suspension bed cold low-pressure separator 3 and the fixed bed cold low-pressure gas from the fixed bed cold low-pressure separator 16 are sent to a desulfurization part 9 for desulfurization, and the gas phase is sent to a hydrogen recovery device 10 after desulfurization; the tail gas enters an absorption desorption tower 21 to recover liquefied gas components or is used as fuel gas to be sent out of the device through the recovered hydrogen obtained by the hydrogen recovery device 10.

The process of the present invention is further illustrated below in conjunction with FIG. 3:

(1) hydrocracking: mixing heavy oil 101, tar slurry 102 containing a hydrocracking catalyst, tower bottom tail oil of a hydrofining fractionating tower and hydrogen, and allowing the mixture to enter a suspension bed reactor 1 for hydrocracking reaction; the liquid phase product after the reaction in the suspended bed reactor 1 is decompressed and then enters a suspended bed thermal high-pressure separator 22 for separation; the liquid phase separated by the suspended bed heat high-pressure separator 22 enters the suspended bed heat low-pressure separator 2 for separation; the liquid phase separated by the suspension bed heat low-pressure separator 2 enters a suspension bed normal-pressure fractionating tower 7 for fractionation; gas phase products after the reaction of the suspended bed reactor 1 are subjected to heat exchange and cooling, and gas phase separated from the suspended bed heat high-pressure separator 22 enters the suspended bed temperature high-pressure separator 4 for separation; the suspended bed temperature high-pressure oil obtained by the suspended bed temperature high-pressure separator 4 is decompressed and then enters a suspended bed temperature low-pressure separator 23 for separation; the liquid phase separated by the suspension temperature-penetration low-pressure separator 23 enters a suspension bed normal-pressure fractionating tower 7 for fractionation; the gas phase separated by the suspension bed temperature low-pressure separator 23 is mixed with the gas phase separated by the suspension bed hot low-pressure separator 2, and enters the suspension bed cold low-pressure separator 3 for separation; after the suspension bed temperature high-pressure gas obtained by the suspension bed temperature high-pressure separator 4 is subjected to heat exchange and cooling, the gas enters a suspension bed cold high-pressure separator 5 to carry out gas-liquid-water three-phase separation; the gas phase at the top of the suspension bed cold high-pressure separator 5 is mixed with part of fresh hydrogen from a fresh hydrogen compressor 18 after being pressurized by the suspension bed circulating hydrogen compressor 6 and then returns to the suspension bed reactor 1 for hydrocracking reaction; the water and the oil in the suspension bed cold high-pressure separator 5 are decompressed and then enter the suspension bed cold low-pressure separator 3 for separation; the gas phase of the suspension bed cold low-pressure separator 3 enters a desulfurization part 9 for desulfurization, and the suspension bed cold low-pressure oil of the suspension bed cold low-pressure separator 3 is sent to a suspension bed normal-pressure fractionating tower 7 for fractionation;

(2) and (3) fractionation after cracking: the suspended bed hot low-fraction oil from the suspended bed hot low-pressure separator 2, the suspended bed hot low-fraction oil from the suspended bed hot high-pressure separator 4 and the suspended bed cold low-fraction oil from the suspended bed cold low-pressure separator 3 enter a suspended bed atmospheric fractionating tower 7 for fractionation, rich gas obtained by separation of the suspended bed atmospheric fractionating tower 7 enters an absorption desorption tower 21, extracted unrefined naphtha and unrefined diesel oil enter a protective agent reactor 11 for hydrofining reaction, tower bottom oil of the suspended bed atmospheric fractionating tower 7 enters a suspended bed vacuum fractionating tower 8 for fractionation, reduced pressure side-stream wax oil of the suspended bed vacuum fractionating tower 8 enters the protective agent reactor 11 for hydrofining reaction, and reduced bottom residues of the suspended bed vacuum fractionating tower 8 are sent out of the device as coking raw materials or circulated to the suspended bed reactor 1 for hydrofining reaction;

(3) and (3) hydrofining: unrefined naphtha, unrefined diesel oil, reduced-pressure side wax oil and hydrogen are mixed and sequentially enter a protective agent reactor 11 and a refining reactor 12 for hydrofining reaction, hydrofining reaction is carried out under a hydrofining catalyst and proper temperature and pressure, and materials such as desulfurization, denitrification, deoxidation, demetalization and the like and olefin saturation are carried out; the hydrorefining reaction product is firstly separated by a fixed bed heat high pressure separator 13; the fixed bed hot high-pressure oil obtained by the fixed bed hot high-pressure separator 13 enters a fixed bed hot low-pressure separator 15 for separation after being depressurized; the fixed bed hot high-pressure gas obtained by the fixed bed hot high-pressure separator 13 enters a fixed bed cold high-pressure separator 14 for gas, liquid and water three-phase separation after heat exchange and cooling; the top gas phase of the fixed bed cold high-pressure separator 14 is subjected to pressure increase by a fixed bed circulating hydrogen compressor 17, mixed with part of fresh hydrogen from a fresh hydrogen compressor 18 and then mixed with a raw material for a hydrofining reaction to form hydrogen mixed oil, and the hydrogen mixed oil sequentially enters the protective agent reactor 11 and the refining reactor 12 to carry out the hydrofining reaction;

(4) fractionation and absorption stabilization after refining: the fixed bed hot low-fraction oil from the fixed bed hot low-pressure separator 15 and the fixed bed cold low-fraction oil from the fixed bed cold low-pressure separator 16 are mixed after heat exchange and enter a refining fractionating tower 19 for fractionation; the gas phase at the top of the refining fractionating tower 19 and the crude naphtha enter an absorption desorption tower 21, the diesel oil generated by the refining fractionating tower 19 is a diesel oil product, and the tail oil at the bottom of the refining fractionating tower 19 circulates to the suspension bed reactor 1 for hydrogenation refining reaction;

the tower top rich gas of the suspension bed atmospheric fractionating tower 7 and the tower top gas of the refined fractionating tower 19 enter an absorption desorption tower 21 after being boosted, and the tower top crude naphtha oil gas of the refined fractionating tower 19 enters the absorption desorption tower 21 after being boosted; the gas phase at the top of the absorption desorption tower 21 is taken as dry gas and sent out of the device, the deethanized oil at the bottom of the tower enters a stabilizing tower 20, a liquefied gas product is obtained from the top of the stabilizing tower 20, and a naphtha product and a circulating absorbent are obtained from the bottom of the tower;

(5) desulfurization and hydrogen recovery: part of the suspension bed cold high-pressure gas from the suspension bed cold high-pressure separator 5, the suspension bed cold low-pressure gas from the suspension bed cold low-pressure separator 3 and the fixed bed cold low-pressure gas from the fixed bed cold low-pressure separator 16 enter a desulfurization part 9 for desulfurization, and the desulfurized gas phase enters a hydrogen recovery device 10; and the tail gas enters an absorption desorption tower 21 to recover liquefied gas components or is sent out of the device as fuel gas through a recovered hydrogen outlet device obtained by the hydrogen recovery device 10.

In the present invention, the liquid phase separated by each separator means the oil phase produced by each separator. Particularly, as the three-phase separation of gas, liquid and water is carried out in the suspension bed cold low-pressure separator and the fixed bed cold low-pressure separator, the terms "suspension bed cold low-pressure oil separation" and "fixed bed cold low-pressure oil separation" respectively refer to oil phases separated by the suspension bed cold low-pressure separator and the fixed bed cold low-pressure separator; in other separators, such as a suspended bed temperature low pressure separator, the term "suspended bed temperature low fraction oil" refers to a liquid phase or an oil phase separated by the suspended bed temperature low pressure separator. Thus, in the present invention, one skilled in the art can determine its origin by its oil name.

In the present invention, the total yield of liquefied gas, naphtha and diesel is calculated as follows:

total yield (%) - (mass flow of liquefied gas + mass flow of naphtha + mass flow of diesel oil)/total mass flow

The method for converting coal tar/heavy oil into liquefied gas, naphtha and diesel adopts a combined mode of a suspension bed and a fixed bed, and uses a non-noble metal catalyst which can be discharged along with bottom reducing residues in a coal tar/heavy oil hydrogenation process of the suspension bed reactor, so that the utilization rate of the catalyst is improved. In addition, compared with the reaction conditions of the conventional suspension bed reactor, the suspension bed reactor has mild hydrogenation reaction conditions, so that the operation is easier to control, the equipment investment and the operation cost are low, and the service life of the suspension bed reactor is prolonged. Therefore, in the method of the invention, the coal tar/heavy oil can be converted into light components with more economic benefits in a full-range mode, and the yield of the light oil and the utilization rate of equipment are improved to the maximum extent.

Examples

The present invention is described in further detail below with reference to examples.

The main properties of the coal tar used in examples 1 to 2 are shown in table 1.

TABLE 1 Main Properties of coal tar

Example 1

According to the flow shown in FIG. 1, coal tar 101(35250kg/h), tar slurry 102(3000kg/h) containing 25 wt% of hydrocracking catalyst, and tower bottom tail oil (7245kg/h) of refined fractionating tower 19 are mixed with hydrogen (10200kg/h) and enter into suspension bed reactor 1 for hydrocracking reaction; the method comprises the steps of reacting under hydrogen, a suspension bed hydrocracking catalyst and 6.5MPaG at 425 ℃, wherein the average residence time of a liquid phase in a suspension bed reactor 1 is 60min, the inlet total hydrogen-oil ratio of the suspension bed reactor 1 is 1500, the feeding amount of the hydrocracking catalyst is 2 wt% of raw oil, and the catalyst is a Fe-based waste catalyst of a Fischer-Tropsch synthesis process.

The liquid phase product in the suspension bed reactor 1 is decompressed and then enters a suspension bed hot low-pressure separator 2 for separation, and the separation is carried out at 425 ℃ and 2.4 MPaG; the liquid phase separated by the suspension bed thermal low-pressure separator 2 enters a suspension bed normal-pressure fractionating tower 7 for fractionation; the gas phase separated by the suspension bed hot low pressure separator 2 enters the suspension bed cold low pressure separator 3 for separation.

The gas phase product after the reaction in the suspension bed reactor 1 enters a suspension bed temperature high pressure separator 4 for separation, and the separation is carried out at 330 ℃ and under the pressure of 6.2 MPaG; the liquid phase of the suspension bed temperature high-pressure separator 4 enters a suspension bed normal-pressure fractionating tower 7 for fractionation; cooling the gas phase in the suspension bed temperature high-pressure separator 4 to 50 ℃, and then feeding the gas phase into a suspension bed cold high-pressure separator 5 for gas-liquid-water three-phase separation; the gas phase separated by the suspension bed cold high-pressure separator 5 is pressurized to 7.2MPaG by the suspension bed circulating hydrogen compressor 6, then mixed with hydrogen pressurized to 7.4MPaG by the fresh hydrogen compressor 18 and divided into two strands according to the proportion of 1:3, the first strand is mixed with the reaction feed of the suspension bed and enters the suspension bed reactor 1 for hydrocracking reaction, and the second strand is heated and then directly enters the bottom of the suspension bed reactor 1.

Reducing the pressure of the oil and the water separated by the suspension bed cold high-pressure separator 5 to 2.3MPaG, and then entering the suspension bed cold low-pressure separator 3 for separation; the oil phase separated by the suspension bed cold low-pressure separator 3 enters a suspension bed normal-pressure fractionating tower 7 for fractionation; the water phase separated by the suspension bed cold low pressure separator 3 is discharged as sewage 136; the gas phase of the suspension bed cold-pressing separator 3 enters a desulfurization part 9 for desulfurization.

The suspension bed hot low fraction oil, the suspension bed hot high fraction oil and the suspension bed cold low fraction oil respectively enter a suspension bed normal pressure fractionating tower 7 for fractionation after heat exchange, the theoretical plate number of the suspension bed normal pressure fractionating tower 7 is 18, the tower top temperature is 132 ℃, the tower bottom temperature is 364 ℃, and the operating pressure is 0.1 MPaG; the gas phase at the top of the suspension bed atmospheric fractionating tower 7 enters an absorption desorption tower 21, and unrefined naphtha at the top of the suspension bed atmospheric fractionating tower 7 and unrefined diesel oil at the lateral line are mixed to a protective agent reactor 11 containing a hydrofining catalyst for hydrofining reaction; heating the bottom oil of the suspension bed atmospheric fractionating tower 7 to 370 ℃ by a heating furnace, entering the suspension bed vacuum fractionating tower 8, wherein the theoretical plate number of the suspension bed vacuum fractionating tower 8 is 26, the tower top temperature is 65 ℃, the tower bottom temperature is 320 ℃, and the operating pressure is-0.098 MPaG; the top gas of the suspension bed vacuum fractionating tower 8 enters a heating furnace, and the vacuum side wax oil enters a protective agent reactor 11 for hydrogenation refining reaction; and the bottom residue 135 of the suspension bed vacuum fractionating tower 8 is sent out of the device.

Mixing unrefined naphtha at the top of the suspension bed atmospheric fractionating tower 7, unrefined diesel oil at the lateral line of the suspension bed atmospheric fractionating tower 7, lateral line wax oil at the suspension bed vacuum fractionating tower and hydrogen, heating the mixture to 370 ℃ by a heating furnace, and sequentially entering a protective agent reactor 11 containing a hydrofining catalyst and a refining reactor 12 to carry out hydrofining reaction; the reaction temperature of the protective agent reactor is 320 ℃, the reaction pressure is 19.2MPaG, and the volume space velocity is 2.0h^-1The hydrogen-oil ratio at the inlet of the reactor is 1500; the reaction temperature of the refining reactor is 385 ℃, the reaction pressure is 18.7MPaG, and the volume space velocity is 0.5h^-1The hydrogen-oil ratio at the inlet of the reactor is 1500; the protective agent reactor 11 adopts ZKM series catalyst developed by Zhongke synthetic oil technology limited, and the refining reactor 12 adopts ZKH series catalyst developed by Zhongke synthetic oil technology limited.

The hydrorefining reaction product enters a fixed bed hot high-pressure separator 13, is separated at 263 ℃ and 18MPaG, the fixed bed hot high oil is depressurized to 2.4MPaG and enters a fixed bed hot low-pressure separator 15 for separation, the fixed bed hot high-pressure gas is cooled to 50 ℃ and enters a fixed bed cold high-pressure separator 14 for separation, the top gas phase of the fixed bed cold high-pressure separator 14 is pressurized to 20.5MPaG by a fixed bed circulating hydrogen compressor 17 and is mixed with the hydrogen pressurized to 20.5MPaG by a new hydrogen compressor 18, and then enters a protective agent reactor 11 for hydrorefining reaction; the liquid phase separated by the fixed bed thermal low-pressure separator 15 is decompressed and then enters a refining fractionating tower 19 for fractionation; cooling the gas phase of the fixed bed hot low-pressure separator 15 to 50 ℃, and separating the gas phase in the fixed bed cold low-pressure separator 16; the liquid phase separated by the fixed bed cold high pressure separator 14 is decompressed to 2.3MPaG, and enters the fixed bed cold low pressure separator 16 for gas, liquid and water three-phase separation; the gas phase separated by the fixed bed cold low pressure separator 16 is introduced into the desulfurization section 9 for desulfurization, the oil phase separated by the fixed bed cold low pressure separator 16 is decompressed and introduced into the refined fractionating tower 19 for fractionation, and the water phase separated by the fixed bed cold low pressure separator 16 is discharged as the sewage 136.

The fixed bed cold low fraction oil and the fixed bed hot low fraction oil are respectively subjected to heat exchange, mixed and heated to 340 ℃ by a heating furnace to a refined fractionating tower 19 for separation, the theoretical plate number of the refined fractionating tower 19 is 23, the tower top temperature is 137 ℃, the tower bottom temperature is 331 ℃, the operating pressure is 0.1MPaG, and the tower top gas phase and the crude naphtha of the refined fractionating tower 19 enter an absorption desorption tower 21; steam stripping the normal first-line diesel oil of the refining fractionating tower 19 by a normal first-line diesel oil stripping tower to obtain a diesel oil product with a stream of 134; the bottom tail oil of the refined fractionating tower 19 is returned to the suspension bed reactor 1 as the circulating oil to carry out the hydrocracking reaction.

The gas phase separated by the suspension bed cold low-pressure separator 3, the gas phase separated by a part of the suspension bed cold high-pressure separator 5 and the gas phase separated by the fixed bed cold low-pressure separator 16 pass through the desulfurization part 9 and then enter the hydrogen recovery device 10; the tail gas of the hydrogen recovery device 10 enters an absorption desorption tower 21, and the hydrogen recovery device recovers hydrogen and enters a new hydrogen compressor 18.

The overhead gas of the suspension bed atmospheric fractionating tower 7, the overhead gas of the refined fractionating tower 19, the overhead crude naphtha of the refined fractionating tower 19 and the tail gas of the hydrogen recovery part enter an absorption desorption tower 21, the theoretical plate number of the absorption desorption tower 21 is 17, the tower top temperature is 45 ℃, the tower bottom temperature is 125 ℃, and the operation pressure is 0.8 MPaG; a dry gas stream 130 obtained from the top of the absorption desorption tower 21 is discharged from the device, deethanized oil at the bottom of the tower enters a naphtha stabilizer 20, the theoretical plate number of the naphtha stabilizer 20 is 25, the tower top temperature is 65 ℃, the tower bottom temperature is 193 ℃, the operating pressure is 0.95MPaG, a liquefied gas stream 131 is obtained from the top of the naphtha stabilizer 20 and is discharged from the device, a first part of naphtha obtained from the bottom of the tower is returned to the absorption desorption tower 21 as a circulating absorbent, the rest of the naphtha is discharged from the device as a naphtha stream 132, and the ratio of the first part to the second part of naphtha is 1: 2.

The mass flow rates of the respective raw materials and products in example 1 are shown in table 2.

TABLE 2 raw materials and product data (unit: kg/h) in example 1

The liquefied gas, naphtha and diesel oil obtained in example 1 of the present invention all meet the national standards. The total yield of liquefied gas, naphtha and diesel oil obtained in example 1 was 80.6%.

Example 2

According to the flow shown in FIG. 2, coal tar 101(35250kg/h), tar slurry 102(3000kg/h) containing 25 wt% of hydrocracking catalyst, and tower bottom tail oil (7245kg/h) of refined fractionating tower 19 are mixed with hydrogen (10200kg/h) and enter into suspension bed reactor 1 for hydrocracking reaction; the method comprises the following steps of reacting under hydrogen, a suspension bed hydrocracking catalyst and 6.5MPaG at 425 ℃ and 6.5, wherein the average residence time of a liquid phase in the suspension bed reactor 1 is 60min, the inlet total hydrogen-oil ratio of the suspension bed reactor 1 is 1500, the feeding amount of the hydrocracking catalyst is 2 wt% of raw oil, and the catalyst is a Fe-based waste catalyst of a Fischer-Tropsch synthesis process.

The liquid phase product in the suspension bed reactor 1 is decompressed and then enters a suspension bed hot low-pressure separator 2, and is separated at 425 ℃ and under the pressure of 2.4 MPaG; the liquid phase separated by the suspension bed thermal low-pressure separator 2 enters a suspension bed normal-pressure fractionating tower 7 for fractionation; the gas phase separated by the suspension bed hot low pressure separator 2 enters the suspension bed cold low pressure separator 3 for separation.

The gas phase product after the reaction in the suspended bed reactor 1 enters a suspended bed temperature high pressure separator 4 for separation at 330 ℃ and 6.2 MPaG; the liquid phase of the suspension bed temperature high-pressure separator 4 enters a suspension bed normal-pressure fractionating tower 7 for fractionation; cooling the gas phase in the suspension bed temperature high-pressure separator 4 to 50 ℃, and then feeding the gas phase into a suspension bed cold high-pressure separator 5 for gas-liquid-water three-phase separation; the gas phase separated by the suspension bed cold high-pressure separator 5 is pressurized to 7.2MPaG by the suspension bed circulating hydrogen compressor 6, then mixed with hydrogen pressurized to 7.4MPaG by the fresh hydrogen compressor 18 and divided into two strands according to the proportion of 1:3, the first strand is mixed with the reaction feed of the suspension bed and enters the suspension bed reactor 1 for hydrocracking reaction, and the second strand is heated and then directly enters the bottom of the suspension bed reactor 1.

The suspended bed hot low fraction oil, the suspended bed hot high fraction oil and the suspended bed cold low fraction oil respectively enter a suspended bed normal pressure fractionating tower 7 for fractionation after heat exchange, the theoretical plate number of the suspended bed normal pressure fractionating tower 7 is 18, the tower top temperature is 132 ℃, the tower bottom temperature is 364 ℃, the operating pressure is 0.1MPaG, and the tower top gas phase of the suspended bed normal pressure fractionating tower 7 enters an absorption desorption tower 21; mixing unrefined naphtha at the top of the suspension bed atmospheric fractionating tower 7 and unrefined diesel oil at the side line into a protective agent reactor 11 for carrying out hydrofining reaction; heating the bottom oil of the suspension bed atmospheric fractionating tower 7 to 370 ℃ by a heating furnace, feeding the heated bottom oil into a suspension bed vacuum fractionating tower 8 for fractionation, wherein the theoretical plate number of the suspension bed vacuum fractionating tower 8 is 26, the tower top temperature is 65 ℃, the tower bottom temperature is 320 ℃, and the operating pressure is-0.098 MPaG; the gas at the top of the suspension bed vacuum fractionating tower 8 enters a heating furnace, the wax oil at the reduced pressure side line enters a protective agent reactor 11 for hydrofining reaction, and the residue 135 at the bottom of the suspension bed vacuum fractionating tower 8 is sent out of the device.

The unrefined naphtha at the top of the suspension bed atmospheric fractionating tower 7, the unrefined diesel oil at the side of the suspension bed atmospheric fractionating tower 7, the paraffin oil at the side of the suspension bed vacuum fractionating tower 7 and hydrogen are mixed and heated to 370 ℃ by a heating furnace, and then the mixture sequentially enters a protective agent reactor 11 containing a hydrofining catalyst and a refining reactionThe vessel 12 carries out hydrofining reaction; the reaction temperature of the protective agent reactor 11 is 320 ℃, the reaction pressure is 19.2MPaG, and the volume space velocity is 2.0h^-1The hydrogen-oil ratio at the inlet of the reactor is 1500; the reaction temperature of the refining reactor 12 is 385 ℃, the reaction pressure is 18.7MPaG, and the volume space velocity is 0.5h^-1The hydrogen-oil ratio at the inlet of the reactor is 1500; the protective agent reactor 11 adopts ZKM series catalyst developed by Zhongke synthetic oil technology limited, and the refining reactor 12 adopts ZKH series catalyst developed by Zhongke synthetic oil technology limited. Mixing the bottom oil of the refining fractionating tower 19 with hydrogen, heating the mixture to 390 ℃ by a heating furnace, and entering a cracking reactor 20 for secondary cracking reaction; the reaction temperature is 395 ℃, the reaction pressure is 18.6MPaG, and the space velocity of the total volume is 2.0h^-1The inlet hydrogen-to-oil ratio of the cracking reactor 20 was 1500. The cracking reactor 20 employs FZC-14 catalyst developed by the comforting petrochemical research institute.

Refined reaction products enter a fixed bed hot high-pressure separator 13, are separated at 263 ℃ and 18MPaG, the fixed bed hot high oil is decompressed to 2.4MPaG and enters a fixed bed hot low-pressure separator 15 for separation, the fixed bed hot high-pressure gas is cooled to 50 ℃ and enters a fixed bed cold high-pressure separator 14 for separation, the gas phase at the top of the fixed bed cold high-pressure separator 14 is boosted to 20.5MPaG by a fixed bed circulating hydrogen compressor 17 and is mixed with hydrogen boosted to 20.5MPaG by a new hydrogen compressor 18, and then is divided into two parts, one part enters a protective agent reactor 11 for hydrogenation refining reaction, and the other part enters a cracking reactor 20 for secondary cracking reaction; the liquid phase separated by the fixed bed hot low pressure separator 15 is decompressed and then enters a refining fractionating tower 19 for fractionation, the gas phase of the fixed bed hot low pressure separator 15 is cooled to 50 ℃ and enters a fixed bed cold low pressure separator 16 for separation, the liquid phase separated by the fixed bed cold high pressure separator 14 is decompressed to 2.3MPaG and enters the fixed bed cold low pressure separator 16 for gas, liquid and water three-phase separation; the gas phase separated by the fixed bed cold low pressure separator 16 is introduced into the desulfurization section 9 for desulfurization, the oil phase separated by the fixed bed cold low pressure separator 16 is decompressed and introduced into the refined fractionating tower 19 for fractionation, and the water phase separated by the fixed bed cold low pressure separator 16 is discharged as the sewage 136.

The fixed bed cold low fraction oil and the fixed bed hot low fraction oil are respectively subjected to heat exchange, mixed, heated to 340 ℃ by a heating furnace, enter a refined fractionating tower 19 for fractionation, the theoretical plate number of the refined fractionating tower 19 is 23, the tower top temperature is 137 ℃, the tower bottom temperature is 331 ℃, the operating pressure is 0.1MPaG, and the tower top gas phase and the crude naphtha of the refined fractionating tower 19 enter an absorption desorption tower 21; steam stripping the normal first-line diesel oil of the refining fractionating tower 19 by a normal first-line diesel oil stripping tower to obtain a diesel oil product with a stream of 134; the bottom tail oil of the refined fractionating tower 19 is returned to the suspension bed reactor 1 as the circulating oil to carry out the hydrocracking reaction.

The overhead gas of the suspension bed atmospheric fractionating tower 7, the overhead gas of the refined fractionating tower 19, the overhead crude naphtha of the refined fractionating tower 19 and the tail gas of the hydrogen recovery part enter an absorption desorption tower 21, the theoretical plate number of the absorption desorption tower 21 is 17, the tower top temperature is 45 ℃, the tower bottom temperature is 125 ℃, and the operation pressure is 0.8 MPaG; a dry gas stream 130 obtained from the top of the absorption desorption tower 21 is discharged out of the device, deethanized oil at the bottom of the tower enters a naphtha stabilizer 22, the number of theoretical plates of the naphtha stabilizer 22 is 25, the temperature at the top of the tower is 65 ℃, the temperature at the bottom of the tower is 193 ℃, and the operating pressure is 0.95 MPaG; a liquefied gas stream 131 is obtained from the top of the naphtha stabilizer 22 and returned to the absorption and desorption column 21 as a recycle absorbent, and the rest of naphtha is taken out of the plant as a naphtha stream 132, and the ratio of the first part to the second part of naphtha is 1: 2.

The mass flow rates of the respective raw materials and products in example 2 are shown in table 3.

TABLE 3 raw materials and product data (unit: kg/h) in example 2

The liquefied gas, naphtha and diesel oil obtained in example 2 of the present invention all meet the national standards. The total yield of liquefied gas, naphtha and diesel oil obtained by this example 2 was 79.8%.

Example 3

The main properties of the heavy oil used in example 3 are shown in table 4.

TABLE 4 Main Properties of heavy oils

Item	Unit of
			Density (20 ℃ C.)	g/cm³	0.964
Viscosity (100 ℃ C.)	mm²/S	108.7
			Freezing point	℃	8.0
Glue	m％	27.67
			C7-asphaltenes	m％	0.25
Residual carbon	m％	6.6
			Ash content	m％	0.089
H/C atomic ratio		1.69
			S	m％	0.13
N	m％	0.45
			Distillation range	℃
IBP		275
			10％		358
30％		433
			50％		492
63％		538
			75％		733
Metal content,. mu.g/g
			Ni	μg/g	11.8
V	μg/g	0.35
			Ca	μg/g	346
Fe	μg/g	10.2
			Acid value	/mgKOH.g^-1	6.35

According to the flow chart shown in FIG. 3, heavy oil 101(1840kg/h), heavy oil slurry 102(160kg/h) containing 25 wt% of hydrocracking reactor, bottom tail oil (378kg/h) of refined fractionating tower 19 and hydrogen (510kg/h) are mixed and enter into suspension bed reactor 1 for hydrogenation refining reaction; the reaction is carried out under the conditions of hydrogen, a suspension bed hydrocracking catalyst, the temperature of 440 ℃ and the pressure of 6.2MPaG, the average residence time of a liquid phase in the suspension bed reactor 1 is 60min, the total hydrogen-oil ratio of an inlet of the suspension bed reactor 1 is 1200, and the feeding amount of the hydrocracking catalyst is 2.04 wt% of the raw material oil. The catalyst is Fe-based waste catalyst of Fischer-Tropsch synthesis process.

The liquid phase product in the suspension bed reactor 1 enters a suspension bed hot high-pressure separator 22, and gas-liquid separation is carried out at 440 ℃ and under the pressure of 6.2 MPaG; the liquid phase separated by the suspended bed heat high-pressure separator 22 enters the suspended bed heat low-pressure separator 2 after being decompressed, and is separated at the temperature of 440 ℃ and the pressure of 2.4 MPaG; the liquid phase separated by the suspension bed heat low pressure separator 2 enters a fractionating tower 7 for fractionation; the gas phase separated by the suspension bed hot low pressure separator 2 enters the suspension bed cold low pressure separator 3 for separation after being cooled; the gas phase product in the suspension bed reactor 1 enters a suspension bed temperature high-pressure separator 4 for separation.

Cooling the gas phase separated by the suspended bed heat high-pressure separator 22 to 330 ℃, then feeding the gas phase into the suspended bed temperature high-pressure separator 4, separating at 330 ℃ and under the pressure of 6.2MPaG, decompressing the liquid phase of the suspended bed temperature high-pressure separator 4 to 2.4MPaG, and then sending the liquid phase to the suspended bed temperature low-pressure separator 23 for separation; the gas phase separated by the suspension bed temperature high-pressure separator 4 is cooled to 50 ℃ and enters a suspension bed cold high-pressure separator 5 for gas, liquid and water three-phase separation; the liquid phase separated by the suspension bed temperature low-pressure separator 23 enters a suspension bed normal pressure fractionating tower 7 for fractionation; the gas phase separated by the suspension bed temperature low-pressure separator 23 is cooled to 50 ℃ and then enters the suspension bed cold low-pressure separator 3 for separation.

The gas phase separated by the suspension bed cold high-pressure separator 5 is pressurized to 7.2MPaG by the suspension bed circulating hydrogen compressor 6, then mixed with hydrogen pressurized to 7.4MPaG by the fresh hydrogen compressor 18 and divided into two strands according to the proportion of 1:3, the first strand is mixed with the reaction feed of the suspension bed and enters the suspension bed reactor 1 for hydrocracking reaction, and the second strand is heated and then directly enters the bottom of the suspension bed reactor 1.

The suspension bed hot low-fraction oil, the suspension bed temperature low-fraction oil and the suspension bed cold low-fraction oil respectively exchange heat and then enter a suspension bed normal pressure fractionating tower 7 for fractionation, the theoretical plate number of the suspension bed normal pressure fractionating tower 7 is 18, the tower top temperature is 174 ℃, the tower bottom temperature is 376 ℃ and the operating pressure is 0.08 MPaG; the gas phase at the top of the suspension bed atmospheric fractionating tower 7 enters an absorption desorption tower 21, and unrefined naphtha at the top of the suspension bed atmospheric fractionating tower 7 and unrefined diesel oil at the lateral line are mixed to a protective agent reactor 11 for hydrogenation refining reaction; heating the bottom oil of the suspension bed atmospheric fractionating tower 7 to 380 ℃ by a heating furnace, feeding the heated bottom oil into a suspension bed vacuum fractionating tower 8 for fractionation, wherein the theoretical plate number of the suspension bed vacuum fractionating tower 8 is 26, the tower top temperature is 65 ℃, the tower bottom temperature is 353 ℃, and the operating pressure is-0.0947 MPaG; the gas at the top of the suspension bed vacuum fractionating tower 8 enters a heating furnace, the wax oil at the reduced pressure side line enters a protective agent reactor 11 for hydrofining reaction, and the residue 135 at the bottom of the suspension bed vacuum fractionating tower 8 is sent out of the device.

Mixing the top unrefined naphtha of the suspension bed atmospheric fractionating tower 7, the lateral line unrefined diesel oil of the suspension bed atmospheric fractionating tower 7, the lateral line wax oil of the suspension bed vacuum fractionating tower and the hydrogen, heating the mixture to 3 DEG by a heating furnaceSequentially entering a protective agent reactor 11 containing a hydrofining catalyst and a refining reactor 12 at 70 ℃ to carry out hydrofining reaction; the reaction temperature of the protective agent reactor 11 is 320 ℃, the reaction pressure is 19.2MPaG, and the volume space velocity is 2.0h^-1The hydrogen-oil ratio at the inlet of the reactor is 1500; the reaction temperature of the refining reactor 12 is 385 ℃, the reaction pressure is 18.7MPaG, and the volume space velocity is 0.5h^-1The hydrogen-oil ratio at the inlet of the reactor is 1500; the protective agent reactor 11 adopts FZC-100, FZC-105, FZC-106 and FZC-204 catalysts which are developed by the comforting petrochemical research institute, and the refining reactor 12 adopts FZC-33 catalysts which are developed by the comforting petrochemical research institute.

The refined reaction product enters a fixed bed thermal high-pressure separator 13, is separated at 63 ℃ and 18MPaG, and the fixed bed thermal high oil is decompressed to 2.4MPaG and enters a fixed bed thermal low-pressure separator 15; the fixed bed hot high-pressure gas is cooled to 50 ℃ and enters a fixed bed cold high-pressure separator 14 for separation; the top gas phase of the fixed bed cold high-pressure separator 14 is boosted to 20.5MPaG by the fixed bed circulating hydrogen compressor 17, then mixed with hydrogen boosted to 20.5MPaG by the new hydrogen compressor 18, and then enters the protective agent reactor 11 for hydrorefining reaction; the liquid phase separated by the fixed bed thermal low-pressure separator 15 is decompressed and then enters a refining fractionating tower 19 for fractionation; cooling the gas phase of the fixed bed hot low-pressure separator 15 to 50 ℃, and separating the gas phase in the fixed bed cold low-pressure separator 16; the liquid phase separated by the fixed bed cold high pressure separator 14 is decompressed to 2.3MPaG, and enters the fixed bed cold low pressure separator 16 for gas, liquid and water three-phase separation; the gas phase separated by the fixed bed cold low pressure separator 16 is introduced into the desulfurization section 9 for desulfurization, the oil phase separated by the fixed bed cold low pressure separator 16 is decompressed and introduced into the refined fractionating tower 19 for fractionation, and the water phase separated by the fixed bed cold low pressure separator 16 is discharged as the sewage 136.

The fixed bed cold low fraction oil and the fixed bed hot low fraction oil are respectively subjected to heat exchange, mixed and heated to 340 ℃ by a heating furnace to a refined fractionating tower 19 for fractionation, the theoretical plate number of the refined fractionating tower 19 is 23, the tower top temperature is 137 ℃, the tower bottom temperature is 331 ℃, the operating pressure is 0.1MPaG, and the tower top gas phase and the crude naphtha of the refined fractionating tower 19 enter an absorption desorption tower 21; steam stripping the normal first-line diesel oil of the refining fractionating tower 19 by a normal first-line diesel oil stripping tower to obtain a diesel oil product with a stream of 134; the bottom tail oil of the refined fractionating tower 19 is returned to the suspension bed reactor 1 as the circulating oil to carry out the hydrocracking reaction.

The overhead gas of the suspension bed atmospheric fractionating tower 7, the overhead gas of the refined fractionating tower 19, the overhead crude naphtha of the refined fractionating tower 19 and the tail gas of the hydrogen recovery part enter an absorption desorption tower 21, the theoretical plate number of the absorption desorption tower 21 is 17, the tower top temperature is 45 ℃, the tower bottom temperature is 125 ℃, and the operation pressure is 0.8 MPaG; a dry gas stream 130 obtained from the top of the absorption desorption tower 21 is discharged out of the device, deethanized oil at the bottom of the tower enters a naphtha stabilizer 20, the number of theoretical plates of the naphtha stabilizer 20 is 25, the temperature at the top of the tower is 65 ℃, the temperature at the bottom of the tower is 193 ℃, and the operating pressure is 0.95 MPaG; a liquefied gas stream 131 is obtained from the top of the naphtha stabilizer 20 and returned to the absorption and desorption column 21 as a recycle absorbent, and the rest of naphtha is taken out of the plant as a naphtha stream 132, and the ratio of the first part to the second part of naphtha is 1: 2.

The mass flow rates of the respective raw materials and products in example 3 are shown in table 5.

TABLE 5 raw materials and product data (unit: kg/h) in example 3

The liquefied gas, naphtha and diesel oil obtained in example 3 of the present invention all meet the national standards. The total yield of liquefied gas, naphtha and diesel oil obtained in example 3 was 84.3%.

It can be seen from the experimental results of examples 1-3 that, by the method of the present invention, the coal tar/heavy oil can be directly sent to the suspension bed reactor for hydrocracking reaction without cutting the coal tar/heavy oil, the obtained suspension bed hydrocracking product is subjected to hydrofining reaction in the fixed bed reactor, and the tail oil at the bottom of the refining fractionating tower of the heavy fraction circulates to the suspension bed reactor or enters the secondary cracking reactor for further lightening, so as to achieve the purpose of maximum production of liquefied gas, naphtha and diesel oil from the whole fraction of the coal tar/heavy oil, improve the utilization efficiency of the raw material, and correspondingly improve the yield of the liquefied gas, naphtha and diesel oil. In addition, compared with the reaction conditions of the conventional suspension bed reactor, the hydrocracking reaction conditions of the suspension bed reactor are mild, so that the operation is easier to control, the equipment investment and the operation cost are low, and the service life of the suspension bed reactor is prolonged.

Claims

1. A process for converting coal tar/heavy oil into liquefied gas, naphtha and diesel, said process comprising the steps of:

2. The method of claim 1, wherein the suspended bed reaction section is a suspended bed reactor.

3. The method of claim 2, wherein the reaction conditions of the suspended bed reactor are as follows: the reaction temperature is 250-550 ℃, the operation pressure is 2.0-10.0 MPaG, the total hydrogen-oil ratio at the inlet of the suspension bed reactor is 200-3000, the average residence time of the liquid phase of the suspension bed reactor is 10-180 min, and the adding amount of the hydrocracking catalyst is 0.02-6 wt% of the raw oil.

4. The method of claim 3, wherein the reaction conditions of the suspended bed reactor are as follows: the reaction temperature is 350-500 ℃, the operation pressure is 4.0-8.0 MPaG, the total hydrogen-oil ratio at the inlet of the suspension bed reactor is 500-2000, the average residence time of the liquid phase of the suspension bed reactor is 60-120 min, and the adding amount of the hydrocracking catalyst is 0.5-4 wt% of the raw oil.

5. The method of any one of claims 1-4, wherein the suspension bed separation section is selected from any one of the following combinations: (1) a suspended bed temperature high-pressure separator, a suspended bed cold high-pressure separator, a suspended bed hot low-pressure separator and a suspended bed cold low-pressure separator; (2) a suspended bed hot high-pressure separator, a suspended bed temperature high-pressure separator, a suspended bed cold high-pressure separator, a suspended bed hot low-pressure separator, a suspended bed temperature low-pressure separator and a suspended bed cold low-pressure separator; or (3) a suspended bed temperature high-pressure separator, a suspended bed cold high-pressure separator, a suspended bed hot low-pressure separator, a suspended bed temperature low-pressure separator and a suspended bed cold low-pressure separator.

6. The method of any one of claims 1-4, wherein the suspension bed fractionation section is selected from any one of the following combinations: (1) a suspension bed atmospheric fractionating tower and a suspension bed vacuum fractionating tower; or (2) a suspension bed stripping tower, a suspension bed normal-pressure fractionating tower and a suspension bed reduced-pressure fractionating tower.

7. The process of claim 6, wherein the suspended bed atmospheric fractionation column and/or the suspended bed vacuum fractionation column is provided with one or more side strippers.

8. The method of any one of claims 1-4, wherein the fixed bed reaction section comprises a protectant reactor and a polishing reactor, and the protectant reactor and polishing reactor are fixed bed reactors.

9. The method of claim 8, wherein the reaction conditions of the protectant reactor are as follows: the reaction temperature is 150-450 ℃, the operation pressure is 6.0-23.0 MPaG, and the volume space velocity is 0.1-10 h^-1And the volume ratio of hydrogen to oil at the inlet of the protective agent reactor is 200-3000.

10. The method of claim 9, wherein the reaction conditions of the protectant reactor are as follows: the reaction temperature is 300-400 ℃, the operation pressure is 10.0-18.0 MPaG, and the volume space velocity is 1-5 h^-1And the volume ratio of hydrogen to oil at the inlet of the protective agent reactor is 500-2000.

11. The method of claim 8, wherein the finishing reactor has reaction conditions as follows: the reaction temperature is 250-500 ℃, the operation pressure is 6.0-23.0 MPaG, and the volume space velocity is 0.05-3 h^-1And the volume ratio of hydrogen to oil at the inlet of the refining reactor is 200-3000.

12. The method of claim 11, wherein the finishing reactor has reaction conditions as follows: the reaction temperature is 350-400 ℃, the operation pressure is 10.0-18.0 MPaG, and the volume space velocity is 0.1-1 h^-1And the volume ratio of hydrogen to oil at the inlet of the refining reactor is 500-2000.

13. The process of any one of claims 1 to 4, wherein the fixed bed separation section is selected from any one of the following combinations: (1) a fixed bed hot high-pressure separator, a fixed bed cold high-pressure separator, a fixed bed hot low-pressure separator and a fixed bed cold low-pressure separator; or (2) a fixed bed cold high pressure separator + a fixed bed cold low pressure separator.

14. The process of any one of claims 1-4, wherein the finishing fractionation section is selected from any one of the following combinations: (1) a stripping tower and an atmospheric fractionating tower; (2) an atmospheric fractionating tower and a vacuum fractionating tower; (3) a stripping tower, an atmospheric fractionating tower and a reduced pressure fractionating tower; or (4) an atmospheric fractionating column.

15. The process of claim 14, wherein the atmospheric and/or vacuum fractionation column is provided with one or more side strippers.

16. The method of any one of claims 1-4, wherein the absorption stabilizer is selected from any one of the following combinations: (1) an absorption desorption tower and a stabilizing tower; (2) an absorption tower, a desorption tower and a stabilizing tower; (3) a naphtha stabilizer column; or (4) a debutanizer + deethanizer.

17. The method of any one of claims 1-4, wherein the bottoms are partially or fully recycled to the slurry reaction section and the bottoms from the finishing fractionation section are recycled to the slurry reaction section for hydrocracking reactions.

18. The method of any one of claims 1 to 4, wherein the bottom tail oil of the refined fractionation section is mixed with hydrogen and recycled to the cracking reactor for secondary cracking reaction,

the cracking reactor is a fixed bed reactor.

19. The process of claim 18 wherein the reaction conditions of the cracking reactor are as follows: the temperature of the reaction was 25 deg.C0-500 ℃, the operation pressure of 6.0-23.0 MPaG and the total volume airspeed of 0.05-6 h^-1And the volume ratio of hydrogen to oil at the inlet of the cracking reactor is 200-3000.

20. The process of claim 19 wherein the reaction conditions of the cracking reactor are as follows:

the reaction temperature is 300-400 ℃, the pressure is 10.0-18.0 MPaG, and the total volume airspeed is 0.2-3 h^-1And the volume ratio of hydrogen to oil at the inlet of the cracking reactor is 500-1500.