CN112745951B

CN112745951B - Method and system for processing aromatic-rich distillate oil

Info

Publication number: CN112745951B
Application number: CN201911054170.7A
Authority: CN
Inventors: 杨清河; 贾燕子; 孙淑玲; 胡大为; 牛传峰; 戴立顺; 王振; 户安鹏; 任亮; 李大东
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2019-10-31
Filing date: 2019-10-31
Publication date: 2022-06-28
Anticipated expiration: 2039-10-31
Also published as: CN112745951A

Abstract

The invention relates to the field of hydrocarbon oil processing, and discloses a method and a system for processing aromatic-rich distillate oil, which comprise the following steps: (1) introducing the aromatic-rich distillate oil into a third reaction unit, performing hydrogenation saturation, and then fractionating to obtain a first light component and a first heavy component; (2) introducing the deoiled asphalt and the aromatic hydrocarbon-containing material flow into a first reaction unit for hydrogenation reaction, wherein the first reaction unit is a fixed bed hydrogenation unit; (3) fractionating the liquid-phase product from the first reaction unit to obtain a second light component and a second heavy component; (41) introducing a second light component into a second reaction unit for reaction; and (42) introducing the second heavy component into the delayed coking unit for reaction; or the second heavy component as a low sulfur marine fuel oil component. The treatment process provided by the invention can realize high-value DOA utilization.

Description

Method and system for processing aromatic-rich distillate oil

Technical Field

The invention relates to the field of hydrocarbon oil processing, in particular to a method for processing aromatic-enriched distillate oil and a system for processing aromatic-enriched distillate oil.

Background

The high-efficiency conversion of residual oil is the core of oil refining enterprises. The fixed bed residual oil hydrogenation is a key technology for efficiently converting residual oil, and has the characteristics of good product quality, mature process and the like.

However, the high content of asphaltene and metals in the residual oil is a limiting factor of the operation period of the fixed bed residual oil hydrogenation device.

In order to solve the problem, a residual oil solvent deasphalting (demetalization) -hydrotreating-catalytic cracking combined process technology (SHF) developed by the research institute of petrochemical and chemical engineering science is an innovative technology for producing clean fuel for vehicles from low-value vacuum residual oil to the maximum extent and prolonging the operation period, but due to the high softening point of deoiled asphaltene (DOA), the transportation and the utilization are difficult, and the popularization of the SHF technology is limited.

The new combined process for producing more propylene by hydrogenating and catalytic cracking (DCC) chemical transformed residual oil is also limited by the influence of asphaltene and metal in the residual oil, the hydrogenated residual oil has low hydrogen content, the residual oil hydrogenation has short running period and low DCC propylene yield, and the economic benefit of the combined technology is influenced.

In addition, a new low-sulfur ship fuel standard with the sulfur mass fraction not more than 0.5% and a low-sulfur petroleum coke standard with the sulfur mass fraction not more than 3.0% are required to be implemented in 2020, and how to produce the low-sulfur ship fuel (low-sulfur petroleum coke) at low cost is also a problem which needs to be solved urgently at present.

Therefore, the conversion of DOA to low sulfur ship fuel or feedstock for low sulfur petroleum coke production is a technical challenge that needs to be addressed.

Disclosure of Invention

The invention aims to overcome the defects that the pressure grade of the device is high and the investment is large because the device must use high hydrogen partial pressure to improve the hydrogen dissolving amount of the raw material and the treatment efficiency of the device when the existing fixed bed residual oil hydrogenation process is used for treating the poor raw material.

In order to achieve the above object, a first aspect of the present invention provides a method for processing an aromatic-rich distillate oil, the method comprising:

(1) introducing aromatic fraction-rich oil into a third reaction unit, carrying out hydrogenation saturation, and then fractionating to obtain a first light component and a first heavy component, wherein the cutting point of the first light component and the first heavy component is 100-250 ℃, and the aromatic content in the first heavy component is more than or equal to 20 mass%;

(2) introducing deoiled asphalt and an aromatic hydrocarbon-containing stream containing the first heavy component into a first reaction unit to perform hydrogenation reaction, wherein the first reaction unit contains a mineral-rich precursor material and/or a hydrogenation catalyst, the first reaction unit is a fixed bed hydrogenation unit, the mineral-rich precursor material is a material capable of adsorbing at least one metal selected from V, Ni, Fe, Ca and Mg, and the dosage ratio of the deoiled asphalt and the aromatic hydrocarbon-containing stream is that a mixed raw material formed by the deoiled asphalt and the aromatic hydrocarbon-containing stream is in a liquid state at the temperature of not higher than 400 ℃;

(3) Fractionating a liquid-phase product from the first reaction unit to obtain a second light component and a second heavy component, wherein the cutting point of the second light component and the second heavy component is 240-450 ℃;

(41) introducing the second light component into a second reaction unit for reaction to obtain at least one product selected from a gasoline component, a diesel component and a BTX raw material component, wherein the second reaction unit is selected from at least one of a hydrocracking unit, a catalytic cracking unit and a diesel hydro-upgrading unit; and

(42) introducing the second heavy component into a delayed coking unit for reaction to obtain at least one product selected from the group consisting of coker gasoline, coker diesel, coker gas oil, and low sulfur petroleum coke; or the second heavy component is used as a low sulfur marine fuel oil component.

A second aspect of the invention provides a system for processing an aromatic-rich distillate oil, the system comprising:

a third reaction unit, wherein the third reaction unit is used for carrying out hydrogenation saturation and fractionation on the aromatic-rich distillate oil to obtain a first light component and a first heavy component;

a first reaction unit which is a fixed bed hydrogenation unit and is in fluid communication with the third reaction unit for subjecting the deoiled asphalt and the aromatic hydrocarbon-containing stream containing the first heavy component from the third reaction unit to a hydrogenation reaction therein;

A separation unit in fluid communication with the first reaction unit for fractionating a liquid phase product from the first reaction unit therein;

a second reaction unit in fluid communication with the separation unit for reacting therein the second light fraction obtained from the separation unit, the second reaction unit being selected from at least one of a hydrocracking unit, a catalytic cracking unit, and a diesel hydro-upgrading unit;

a delayed coking unit in fluid communication with the separation unit for reacting the second heavy component obtained from the separation unit therein to yield at least one product selected from the group consisting of coker gasoline, coker diesel, coker gas oil, and low sulfur petroleum coke;

an outlet in fluid communication with the separation unit for withdrawing the second heavy component obtained from the separation unit as a low sulfur bunker fuel oil component from the system.

The invention is especially suitable for the hydro-conversion of normal slag and slag reduction, and is especially suitable for the hydro-conversion of inferior residual oil with high metal, high carbon residue, high condensed ring substances and high nitrogen content.

The DOA and the aromatic hydrocarbon-containing material flow are subjected to fixed bed hydrotreatment (such as hydrodesulfurization), and the hydrogenated second light component is subjected to hydrocracking (RLG or RLA) to produce BTX, gasoline fraction and diesel oil fraction or catalytic cracking (LTAG) to produce gasoline fraction (and liquefied gas); the hydrogenated second heavy component produces heavy low sulfur ship fuel or low sulfur petroleum coke.

The treatment process provided by the invention can realize high-value DOA utilization.

The invention can process the inferior raw material under lower hydrogen partial pressure and can realize the long-period stable operation of the device.

Drawings

FIG. 1 is a process flow diagram for processing an aromatic-rich distillate oil in accordance with a preferred embodiment of the present invention.

Description of the reference numerals

1 heavy oil raw material 2 solvent deasphalting unit

3 deasphalted oil 4 deoiled asphalt

5 aromatic compound 6 mixed raw material

7 first reaction unit 8 second light component

9 second heavy component 10 second reaction Unit

11 delayed coking unit 12 BTX feedstock composition

13 gasoline component 14 Diesel component

15 coking gasoline and 16 coking diesel oil

17 coking waxy oil 18 low sulfur petroleum coke

19 separation unit 20 aromatic rich distillate oil

21 third reaction unit 22 first heavy fraction

Detailed Description

The endpoints of the ranges and any values disclosed herein are not limited to the precise range or value, and such ranges or values should be understood to encompass values close to those ranges or values. For ranges of values, between the endpoints of each of the ranges and the individual points, and between the individual points may be combined with each other to give one or more new ranges of values, and these ranges of values should be considered as specifically disclosed herein.

As mentioned above, the first aspect of the present invention provides a method for processing an aromatic-rich distillate oil, comprising:

(1) introducing the aromatic-rich distillate oil into a third reaction unit, carrying out hydrogenation saturation, and then fractionating to obtain a first light component and a first heavy component, wherein the cutting point of the first light component and the first heavy component is 100-250 ℃, and the aromatic content in the first heavy component is more than or equal to 20 mass percent;

(2) introducing deoiled asphalt and an aromatic hydrocarbon-containing stream containing the first heavy component into a first reaction unit to perform hydrogenation reaction, wherein the first reaction unit contains a mineral-rich precursor material and/or a hydrogenation catalyst, the first reaction unit is a fixed bed hydrogenation unit, the mineral-rich precursor material is a material capable of adsorbing at least one metal selected from V, Ni, Fe, Ca and Mg, and the dosage ratio of the deoiled asphalt and the aromatic hydrocarbon-containing stream is such that a mixed raw material formed by the deoiled asphalt and the aromatic hydrocarbon-containing stream is in a liquid state at a temperature of not higher than 400 ℃;

Preferably, the deoiled asphalt and the aromatic hydrocarbon-containing stream are used in a ratio such that a mixed feedstock formed from the deoiled asphalt and the aromatic hydrocarbon-containing stream is in a liquid state at a temperature of not higher than 280 ℃; it is further preferred that the deoiled bitumen and the aromatic hydrocarbon-containing stream are used in such a ratio that the mixed feedstock formed from the deoiled bitumen and the aromatic hydrocarbon-containing stream is in a liquid state at a temperature of not higher than 100 ℃.

According to the invention, the hydrogenation saturation reaction in the third reaction unit is partial hydrogenation saturation, and the cutting point of the first light component and the first heavy component is 180 ℃.

The first light component preferably enters a catalytic cracking unit to produce the low-carbon olefin.

Particularly preferably, the cut point of the second light component and the second heavy component is 350 ℃.

Preferably, in step (2), the deoiled asphalt and the aromatic hydrocarbon-containing stream are used in such a ratio that the viscosity at 100 ℃ of the mixed feedstock formed by the deoiled asphalt and the aromatic hydrocarbon-containing stream is not more than 400mm²S, more preferably not more than 200mm²S, further preferably not more than 100mm²/s。

Preferably, in step (2), the aromatic hydrocarbon-containing stream further contains aromatic hydrocarbon compounds and/or aromatic hydrocarbon oils, and the aromatic hydrocarbon oils are at least one selected from LCO, HCO, FGO (catalytic heavy distillate oil), ethylene tar, coal tar, coker diesel oil, and coker wax oil.

Preferably, the aromatic hydrocarbon compound is selected from one or more of benzene, toluene, xylene, naphthalene, methylnaphthalene, multi-branched naphthalene and aromatic hydrocarbon with more than two rings, and polycyclic aromatic hydrocarbon with no more than three rings or a mixture thereof is preferred. Particularly preferably, the aromatic hydrocarbon compound is selected from benzene, toluene, xylene, naphthalene, a mixture of at least one C _1-6At least one of alkyl substituted naphthalene and aromatic hydrocarbon with more than three rings.

More preferably, the aromatic hydrocarbon content in the aromatic-rich distillate oil is not less than 20% by mass, still more preferably not less than 25% by mass, and still more preferably not less than 40% by mass.

Preferably, in step (2), the de-oiled asphalt is obtained by subjecting a heavy oil feedstock to a solvent de-asphalting process in a solvent de-asphalting unit.

Preferably, in the solvent deasphalting unit, the mass fraction of the yield of the deasphalted asphalt is not more than 50%, more preferably not more than 40%, and still more preferably not more than 30%.

According to a preferred embodiment, in step (2), the mass ratio of the amount of deoiled asphalt to the aromatic hydrocarbon-containing stream is from 1:10 to 50:10, more preferably from 2:10 to 30: 10; further preferably from 3:10 to 15: 10.

Preferably, the method of the present invention further comprises: recycling the coker diesel and/or coker wax obtained in step (42) back to step (2) as at least part of the aromatics-containing stream.

Preferably, in step (1), the third reaction unit is at least one of a fixed bed reactor, a moving bed reactor and an ebullating bed reactor.

Preferably, the operating conditions in the third reaction unit include: the reaction temperature is 200-420 ℃, the reaction pressure is 2-18MPa, and the liquid hourly space velocity is 0.3-10h^-1Hydrogen-oil volume ratio of 50-5000; more preferably, the operating conditions in the third reaction unit include: the reaction temperature is 220 ℃ and 400 ℃, the reaction pressure is 2-15MPa, and the liquid hourly space velocity is 0.3-5h^-1The volume ratio of hydrogen to oil is 50-4000.

Preferred embodiments in the third reaction unit of the present invention are provided below:

the conditions for partial hydrogenation saturation of the aromatic-rich distillate oil in the presence of hydrogen are generally as follows: the partial hydrogenation saturation technology of the aromatic-rich distillate oil is a fixed bed/ebullated bed/moving bed hydrotreating technology. Taking the industrially mature fixed bed diesel oil or wax oil hydrogenation technology as an example, the reactor or the reaction bed layer at least comprises a hydrofining catalyst. The hydrofining catalyst used in partial hydrogenation saturation of the aromatic-rich distillate oil preferably has good and moderate hydrogenation saturation activity to avoid further saturation of the tetralin structure into a decalin or cycloparaffin structure with lower hydrogen supply capacity. These catalysts are generally catalysts in which a porous refractory inorganic oxide such as alumina is used as a carrier, oxides of metals of the VIB group and/or VIII group such as W, Mo, Co, Ni and the like are used as active components, and other various auxiliaries such as P, Si, F, B and the like are selectively added, and RS series pretreatment catalysts developed by the institute of petrochemical engineering science belong to the catalysts. The RS series catalyst is a NiMo catalyst.

According to the invention, the first reaction unit is preferably a medium/low pressure fixed bed hydrogenation unit.

Preference for the situationIn step (2), the operating conditions in the first reaction unit include: the reaction temperature is 260-500 ℃, the reaction pressure is 2.0-20.0 MPa, the preferable pressure is 2-12 MPa, the volume ratio of hydrogen to oil is 100-1200, and the liquid hourly space velocity is 0.1-1.5 h^-1. The liquid hourly space velocity and reaction pressure are selected according to the characteristics of the material to be treated and the desired conversion and refining depth.

Preferably, in the step (2), the ore-rich precursor material contains a carrier and an active component element loaded on the carrier, wherein the carrier is selected from at least one of aluminum hydroxide, aluminum oxide and silicon oxide, and the active component element is selected from at least one of metal elements in groups VIB and VIII. More preferably, the active component in the mineral-rich precursor material is an oxide and/or sulfide of a metal element selected from groups VIB and VIII.

Preferably, in the step (2), the ignition loss of the ore-rich precursor material is not less than 3 mass%, and the specific surface area is not less than 80m²(ii) water absorption of not less than 0.9 g/g. The ignition reduction refers to the mass percentage of the reduced mass of the ore-rich precursor material after 600 ℃/2h roasting treatment to the mass before roasting; the water absorption rate refers to the percentage of the mass of the ore-rich precursor material which is increased in the soaking water for half an hour at room temperature (for example, 25 ℃) to the mass before soaking.

According to a preferred embodiment, in step (2), the first reaction unit is sequentially charged with a first ore-rich precursor material and a second ore-rich precursor material, in terms of the reactant flow direction, and the first ore-rich precursor material has a firing rate not less than that of the first ore-rich precursor material.

According to the foregoing preferred embodiment, it is further preferred that the ignition loss of the first mineral-rich precursor material is 3 to 15 mass%, and the ignition loss of the second mineral-rich precursor material is not less than 15 mass%.

According to the foregoing preferred embodiment, it is further preferred that the packing volume ratio of the first mineral-rich precursor material to the second mineral-rich precursor material is from 5:95 to 95: 5.

The hydrogenation catalyst of the present invention may be a graded combination of different catalysts, preferably the hydrogenation catalyst is at least capable of catalyzing hydrodemetallization and hydrodesulfurization reactions.

The present invention is not particularly limited with respect to the specific type of catalyst capable of catalyzing the hydrodemetallization reaction, hydrodesulfurization reaction, hydrodeasphaltization reaction, and hydrodecarbonization reaction, and a catalyst capable of catalyzing the above reaction, which is conventionally used in the art, may be used.

The hydrogenation catalyst of the invention can be, for example, a catalyst which takes a porous refractory inorganic oxide as a carrier, takes oxides or sulfides of metals in a VIB group and/or a VIII group as active components, and selectively adds an auxiliary agent.

Preferably, after the first reaction unit runs for a long period, the ore-rich precursor material can be converted into a vanadium-rich material, and the vanadium content in the vanadium-rich material is not less than 10% by mass; particularly preferably, the mineral-rich precursor material is converted into a vanadium-rich material with a V content of 20 mass% or more, so that high-value V can be directly refined₂O₅。

Preferred embodiments in the first reaction unit of the present invention are provided below:

the raw material hydrotreating technology related in the first reaction unit of the invention is a fixed bed hydrotreating technology, taking the mature fixed bed heavy oil and residual oil hydrotreating technology in the industry at present as an example, the reactor or the reaction bed layer at least comprises an ore-rich precursor material and/or a hydrogenation catalyst, and the ore-rich precursor material mainly comprises two parts: the carrier has strong capability of adsorbing vanadium-containing organic compounds in oil, and the active component has hydrogenation activity function. The carrier is mainly obtained by extruding, molding and drying silicon oxide, aluminum hydroxide or aluminum hydroxide/aluminum oxide mixture, the surface of the carrier is rich in-OH, the carrier has strong adsorption capacity on vanadium-containing organic compounds in oil, and the ignition loss of the carrier is not less than 5 mass percent after roasting for 2 hours at 600 ℃. The active component mainly adopts oxides or sulfides of metals of VIB group and/or VIII group such as W, Mo, Co, Ni and the like.

The hydrogenation catalyst referred to in the foregoing preferred embodiment is generally a heavy residue hydrogenation catalyst, and the heavy residue hydrogenation catalyst refers to a combined catalyst having functions of heavy residue and residue hydrodemetallization, hydrodesulfurization, hydrodecarbonization, and the like. These catalysts are generally catalysts in which a porous refractory inorganic oxide such as alumina is used as a carrier, oxides or sulfides of metals of group VIB and/or group VIII such as W, Mo, Co, Ni and the like are used as active components, and other various auxiliaries such as elements P, Si, F, B and the like are selectively added, such as RDM, RCS series heavy metals, residual oil hydrodemetallization catalysts and desulfurization catalysts developed by petrochemical engineering science research institutes. At present, in the fixed bed residual oil hydrogenation technology, a plurality of catalysts are often used together. In the present invention, a rich mineral precursor material, a hydrodemetallation desulfurization catalyst and a hydrodesulfurization catalyst are preferably used, and the loading order is generally such that the raw material is sequentially brought into contact with the rich mineral precursor material, the hydrodemetallation desulfurization and the hydrodesulfurization catalyst. Of course, there is a technique of mixing and packing these catalysts.

Preferably, in step (41), the second reaction unit is a hydrocracking unit, and the operating conditions in the hydrocracking unit include: the reaction temperature is 360-420 ℃, the reaction pressure is 10.0-18.0 MPa, the volume ratio of hydrogen to oil is 600-2000, and the liquid hourly space velocity is 1.0-3.0 h^-1。

Preferably, the hydrocracking unit is packed with at least one hydrotreating catalyst and at least one hydrocracking catalyst.

Preferably, the hydrocracking unit is a fixed bed hydrocracking unit.

Preferred embodiments in the second reaction unit of the present invention are provided below:

in the step (41), the second light component is introduced into a second reaction unit for reaction, and the hydrocracking technology adopted is fixed bed hydrocracking technology. Taking the industrial mature fixed bed wax oil hydrocracking technology as an example, the reactor or the reaction bed layer at least comprises two hydrocracking catalysts, one is a pretreatment catalyst and the other is a hydrocracking catalyst. Because the metal content, the sulfur content, the nitrogen content and the carbon residue value of the material obtained by fractionation after the fixed bed hydrogenation treatment are all higher, the pretreatment catalyst preferably has strong demetallization activity and good desulfurization and denitrification activity so as to ensure the activity of the subsequent hydrocracking catalyst. The hydrocracking catalyst preferably has good hydrocracking activity and high VGO conversion and HDS activity. These catalysts are generally catalysts in which a porous refractory inorganic oxide such as alumina or molecular sieve is used as a carrier, oxides of metals of the VIB group and/or VIII group such as W, Mo, Co, Ni and the like are used as active components, and other various additives such as P, Si, F, B and the like are selectively added, for example, RS series pretreatment catalysts and RHC series hydrocracking catalysts developed by the institute of petrochemical engineering science and technology belong to the catalysts. The RS series catalyst is a NiW catalyst, and the RHC series catalyst is a NiMo molecular sieve catalyst.

Preferably, in step (41), the second reaction unit is a catalytic cracking unit, and the catalytic cracking unit is a Fluid Catalytic Cracking (FCC) unit.

Preferably, the second light component catalytic cracking technology used in the second light component catalytic cracking is a Fluid Catalytic Cracking (FCC) technology, preferably LTAG technology developed by the institute of petrochemical science and technology, which mainly produces gasoline fractions and liquefied gas.

Preferably, the operating conditions in the fluid catalytic cracking unit include: the reaction temperature is 500-600 ℃, the agent-oil ratio is 3-12, and the retention time is 0.6-6 s.

Unless otherwise specified, the oil-to-agent ratios of the present invention all represent oil-to-agent mass ratios.

Preferably, in step (41), the second reaction unit is a diesel hydro-upgrading unit, and the operating conditions in the diesel hydro-upgrading unit include: the reaction temperature is 330-420 ℃, the reaction pressure is 5.0-18.0 MPa, the volume ratio of hydrogen to oil is 500-2000, and the liquid hourly space velocity is 0.3-3.0 h^-1。

Preferably, the diesel hydro-upgrading unit is loaded with at least one diesel hydro-upgrading catalyst.

The diesel hydrogenation upgrading catalyst can be an RS series pretreatment catalyst and an RHC-100 series diesel hydrocracking catalyst which are researched and developed by petrochemical engineering scientific research institute, for example.

Preferably, in step (42), the second heavy component is introduced into a delayed coking unit for reaction to obtain at least one product selected from coker gasoline, coker diesel, coker gas oil and low sulfur petroleum coke, and the operating conditions in the delayed coking unit include: the reaction temperature is 440-520 ℃, and the retention time is 0.1-4 h.

Preferably, in step (42), the sulfur content of the second heavy component is no greater than 1.8 mass%, the second heavy component is introduced into a delayed coking unit for reaction to give low sulfur petroleum coke, more preferably the sulfur content of the low sulfur petroleum coke is no greater than 3 mass%.

Preferably, in step (42), the second heavy component is used as a low-sulfur bunker fuel oil component, and the sulfur content of the low-sulfur bunker fuel oil component is not more than 0.5 mass%.

The present invention is not particularly limited with respect to the specific operation of the solvent deasphalting treatment, and may be carried out by a solvent deasphalting process which is conventional in the art. The operating parameters of the solvent deasphalting process are exemplified in the examples of the present invention and those skilled in the art should not be construed as limiting the invention.

The invention is suitable for the hydroconversion of normal slag and slag reduction, and is particularly suitable for the hydroconversion of poor residual oil of high metal (Ni + V >150 mug/g, particularly Ni + V >200 mug/g), high carbon residue (the mass fraction of the carbon residue is >17 percent, particularly the mass fraction of the carbon residue is >20 percent) and high condensed ring substances.

As mentioned above, a second aspect of the present invention provides a system for processing an aromatic-rich distillate oil, the system comprising:

a third reaction unit, which is used for carrying out hydrogenation saturation and fractionation on the aromatic-rich distillate oil to obtain a first light component and a first heavy component;

a first reaction unit, which is a fixed bed hydrogenation unit and is in fluid communication with the third reaction unit, for subjecting deoiled asphalt and an aromatic-containing stream containing the first heavy component from the third reaction unit to a hydrogenation reaction therein;

a separation unit in fluid communication with the first reaction unit for fractionating therein a liquid phase product from the first reaction unit;

a delayed coking unit in fluid communication with the separation unit for reacting therein the second heavy components obtained from the separation unit to yield at least one product selected from the group consisting of coker gasoline, coker diesel, coker gas oil, and low sulfur petroleum coke;

Preferably, said delayed coking unit is in fluid communication with said first reaction unit for recycling said coker gas oil and/or said coker gas oil obtained in said delayed coking unit back to said first reaction unit as at least part of said aromatic-containing stream.

Preferably, the system further comprises a solvent deasphalting unit, which is in fluid communication with the first reaction unit and is used for solvent deasphalting the heavy oil feedstock therein, and introducing the deasphalted asphalt obtained after the solvent deasphalting into the first reaction unit.

According to a preferred embodiment, in the system of the invention, the second reaction unit is a hydrocracking unit.

According to another preferred embodiment, in the system of the present invention, the second reaction unit is a catalytic cracking unit and the catalytic cracking unit is a fluid catalytic cracking unit.

According to another preferred embodiment, in the system of the present invention, the second reaction unit is a diesel hydro-upgrading unit.

The method of processing an aromatic-rich distillate oil according to the invention is described in further detail below with reference to fig. 1.

As shown in fig. 1, an aromatic-rich distillate oil 20 is introduced into a third reaction unit 21 to be subjected to hydrogenation saturation and then fractionated to obtain a first light component and a first heavy component 22; and the heavy oil raw material 1 enters a solvent deasphalting unit 2 to be subjected to solvent deasphalting treatment to obtain deoiled asphalt 4 and deasphalted oil 3; the deoiled asphalt 4 and the aromatic hydrocarbon-containing material flow containing the first heavy component 22 form a mixed raw material 6 and enter a first reaction unit 7 for hydrogenation reaction, the aromatic hydrocarbon-containing material flow preferably also contains aromatic hydrocarbon compounds 5 from the outside, wherein the first reaction unit contains an ore-rich precursor material and a hydrogenation catalyst capable of catalyzing at least one reaction selected from hydrodemetallization reaction, hydrodesulfurization reaction, hydrodeasphalting reaction and hydrodecarbonization reaction, and the first reaction unit is a fixed bed hydrogenation unit; the liquid-phase product from the first reaction unit 7 enters a separation unit 19 for fractionation to obtain a second light component 8 and a second heavy component 9, wherein the cut points of the second light component and the second heavy component are 240-450 ℃; introducing the second light component 8 into a second reaction unit 10 to react to obtain at least one product selected from a gasoline component 13, a BTX raw material component 12 and a diesel oil component 14, wherein the second reaction unit is selected from at least one of a hydrocracking unit, a catalytic cracking unit and a diesel oil hydrogenation upgrading unit; and introducing the second heavy fraction 9 into a delayed coking unit 11 for reaction to obtain at least one product selected from the group consisting of coker gasoline 15, coker diesel 16, coker gas oil 17, and low sulfur petroleum coke 18; or the second heavy component 9 is used as a low sulfur marine fuel oil component.

The technology of the invention enables high-efficiency conversion of heavy oil and production of gasoline, BTX feedstock, and flexible production of low-sulfur ship fuel and low-sulfur petroleum coke.

Compared with the prior art, the method has the advantages that organic combination of processes such as residual oil hydrogenation, hydrocracking or catalytic cracking is adopted, so that not only is the light petroleum fraction utilized with high value, but also the low-value DOA is converted into the low-sulfur ship fuel component and the low-sulfur raw material which meet the environmental protection requirement, and the high-efficiency, environmental-friendly and comprehensive utilization of the heavy petroleum resource is realized.

In addition, the technology provided by the invention can enable DOA to be efficiently converted in a medium-low pressure fixed bed hydrogenation reactor and can produce gasoline fraction and BTX raw material, and can provide raw material for producing low-sulfur ship fuel and low-sulfur coke products.

The present invention will be described in detail below by way of examples. The following examples were carried out using the process flow shown in FIG. 1, unless otherwise specified. And, without being specifically stated, the following examples have the following common features:

the results of table 3 in the following examples are, without specific mention, the average of the results obtained in the sampling test every 25h in the continuous operation of the apparatus for 100 h.

Aromatic fraction-rich partial hydrogenation saturation experiments were performed on a medium-sized fixed bed diesel hydrotreater, and the total volume of the reactor was 200 mL.

In the following examples, the hydrogenation catalyst and materials used for the partial hydrosaturation of the aromatic-rich distillate oil are RS series hydrogenation catalysts, under the designation RS2100, developed by the institute of petrochemical science.

RS-2100 refined catalyst, RHC-100 diesel oil hydrogenation modified catalyst, RHC-131 hydrocracking catalyst, RG-30B protective catalyst, RDM-33B residual oil demetalization desulfurization transition catalyst, RCS-31 desulfurization catalyst by the petrochemical industry and chemical engineering science research institute development.

The catalytic cracking catalyst MLC-500 was produced by Changjingtie division of China petrochemical catalyst, Inc.

The normal temperature below indicates 25. + -. 3 ℃.

Example A

Preparation of the ore-rich precursor material 1: selecting 2000g of RPB110 pseudo-boehmite produced by Changjingtang division of China petrochemical catalyst limited, wherein 1000g of RPB110 pseudo-boehmite is treated at 550 ℃ for 2h to obtain about 700g of alumina, fully mixing about 700g of alumina and another 1000g of pseudo-boehmite, then adding 40g of sesbania powder and 20g of citric acid, adding 2200g of deionized water, kneading, extruding into strips for forming, drying at 300 ℃ for 3h to obtain about 1730g of carrier, adding 2100mL of solution containing Mo and Ni for saturated impregnation, wherein the Mo content in the solution is MoO ₃5.5% by mass and 1.5% by mass of Ni, based on the mass of NiO, were immersed for half an hour, and then treated at 180 ℃ for 4 hours to obtain an ore-rich precursor material 1, properties of which are shown in table 6.

Preparation of the mineral-rich precursor material 2: selecting 2000g of RPB110 pseudo-boehmite produced by Changling division of China petrochemical catalyst Limited company, adding 30g of sesbania powder and 30g of citric acid, adding 2400g of deionized water, kneading, extruding into strips, drying at 120 ℃ for 5 hours to obtain about 2040g of carrier, adding 2200mL of solution containing Mo and Ni for saturated dipping, wherein the content of Mo in the solution is MoO₃7.5% by mass and 1.7% by mass of Ni, based on the mass of NiO, were immersed for half an hour, and then treated at 200 ℃ for 3 hours to obtain an ore-rich precursor material 2, properties of which are shown in table 6.

Preparation of the ore-rich precursor material 3: selecting 2000g of commercially available silicon oxide, adding 30g of sesbania powder and 30g of sodium hydroxide, adding 2400g of deionized water, kneading, extruding into strips, forming, drying at 120 ℃ for 5 hours to obtain a carrier, adding 2200mL of solution containing Mo and Ni for saturated impregnation, wherein the content of Mo in the solution is MoO₃4.5% by weight and a Ni content of 1.0% by weight of NiO, were immersed for half an hour and then treated at 200 ℃ for 3 hours to obtain a mineral-rich precursor material 3, properties of which are shown in table 6.

Example 1

Third reaction unit: the raw material aromatic-rich distillate oil is LCO1 (properties are shown in a table 1) and comes from a Yangzi refining catalytic cracking device; third reaction unit operating conditions: the reaction temperature is 290 ℃, the reaction pressure is 4MPa, and the liquid is in the liquid stateThe volume space velocity is 1h^-1The volume ratio of hydrogen to oil is 800: 1.

first fractionation: the cut points of the first light component and the first heavy component 1 (properties see table 1) are 180 ℃;

a first reaction unit: the raw material DOA (from the heavy weight reduction slag) and the first heavy component 1 are mixed according to the mass ratio of 1:10, and the properties are shown in a table 2; the medium-sized fixed bed residual oil hydrotreatment device is characterized in that the total volume of a reactor is 200mL, an RG-30B protective catalyst, an ore-rich precursor material 1, an ore-rich precursor material 2, an RDM-33B residual oil demetalization desulfurization transition catalyst and an RCS-31 desulfurization catalyst are sequentially filled in a first reaction unit according to the material flow direction, and the filling ratio is as follows: RG-30B: mineral-rich precursor material 1: ore-rich precursor material 2: RDM-33B: RCS-31 ═ 6: 30: 30: 14: 20 (V/V); the operating conditions were: the reaction temperature is 360 ℃, the reaction pressure is 8MPa, and the liquid hourly space velocity is 0.3h^-1The volume ratio of hydrogen to oil is 800: 1. the product properties after hydrogenation of the mixed feedstock are shown in table 3.

Second fraction : and (3) fractionating the liquid-phase product obtained by the treatment of the first reaction unit to obtain a second light component with the temperature of less than 350 ℃ and a second heavy component with the temperature of more than or equal to 350 ℃, wherein the properties of the second heavy component are shown in Table 4.

The second light fraction was tested in a second reaction unit.

Second reaction unit: a fixed bed hydrocracking device is sequentially filled with RS-2100, RHC-131 is 40: 60(V/V), the operating conditions are: the reaction temperature of the refining section is 370 ℃, the reaction temperature of the cracking section is 385 ℃, the reaction pressure is 10MPa, and the liquid hourly space velocity is 2.0h^-1The volume ratio of hydrogen to oil is: 1200: 1; a hydrocracked product was obtained with the properties shown in Table 5.

Example 2

Third reaction unit: the raw material, aromatic-rich distillate oil is HCO2 (properties are shown in Table 1), and comes from a Zhehai refining catalytic cracking unit; third reaction unit operating conditions: the reaction temperature is 330 ℃, the reaction pressure is 6MPa, and the liquid hourly space velocity is 1h^-1The volume ratio of hydrogen to oil is 800: 1.

first fractionation: first light groupThe cut point of fraction and first heavy fraction 2 (properties see table 1) was 190 ℃;

a first reaction unit: feed, DOA (from gravity reduction slag) and first heavy fraction 2 in a ratio of 5: 10 in a mass ratio, and the properties are shown in table 2; the treatment apparatus and catalyst loading were the same as in example 1, with the operating conditions: the reaction temperature is 380 ℃, the reaction pressure is 10MPa, and the liquid hourly space velocity is 0.3h ^-1The volume ratio of hydrogen to oil is 800: 1. the product properties after hydrogenation of the mixed feedstock are shown in table 3.

Second fraction: and (3) fractionating the liquid-phase product obtained by the treatment of the first reaction unit to obtain a second light component with the temperature of less than 350 ℃ and a second heavy component with the temperature of more than or equal to 350 ℃, wherein the properties of the second heavy component are shown in Table 4.

The second light fraction was tested in a second reaction unit.

Second reaction unit: the conditions were the same as in example 1, and a hydrocracked product was obtained, the properties of which are shown in Table 5.

Example 3

Third reaction unit: the raw material, rich aromatic distillate oil is LCO1 (properties are shown in Table 1), and comes from a Yangzi refining catalytic cracking device; third reaction unit operating conditions: the reaction temperature is 320 ℃, the reaction pressure is 6MPa, and the liquid hourly space velocity is 1h^-1The volume ratio of hydrogen to oil is 800: 1.

first fractionation: the cut points of the first light fraction and the first heavy fraction 3 (properties see table 1) are 190 ℃;

a first reaction unit: the feedstock, DOA (from gravity reduction slag) and first heavy fraction 3 were mixed in a ratio of 10: 10 in a mass ratio, and the properties are shown in table 2; the treatment apparatus and catalyst loading were the same as in example 1, with the operating conditions: the reaction temperature is 370 ℃, the reaction pressure is 6MPa, and the liquid hourly space velocity is 0.3h ^-1The volume ratio of hydrogen to oil is 800: 1. the product properties after hydrogenation of the mixed feed are shown in Table 3.

Second fraction: fractionating the liquid phase product obtained by the first reaction unit treatment to obtain a second light component at a temperature of less than 350 ℃ and a second heavy component at a temperature of greater than or equal to 350 ℃, wherein the second heavy componentThe properties are shown in Table 4.

The second heavy component is subjected to coking reaction at a reaction temperature of 500 ℃ for a residence time of 0.5h to obtain petroleum coke (yield of 30 mass%), wherein the sulfur content is 2.7 mass%.

The second light fraction was tested in a second reaction unit.

Example 4

Third reaction unit: the raw material, aromatic-rich distillate oil is coal tar (properties are shown in table 1), and comes from a certain coal coking device in China; third reaction unit operating conditions: the reaction temperature is 300 ℃, the reaction pressure is 10MPa, and the liquid hourly space velocity is 0.8h^-1The volume ratio of hydrogen to oil is 800: 1.

first fractionation: the cut points of the first light fraction and the first heavy fraction 4 (properties see table 1) are 190 ℃;

a first reaction unit: the feedstock, DOA (from gravity reduction slag) and the first heavy fraction 4 were mixed in a 15: 10 in a mass ratio, and the properties are shown in table 2; the treatment apparatus and catalyst loading were the same as in example 1, with the operating conditions: the reaction temperature is 350 ℃, the reaction pressure is 12MPa, and the liquid hourly space velocity is 0.3h ^-1The volume ratio of hydrogen to oil is 800: 1. The product properties after hydrogenation of the mixed feed are shown in Table 3.

The second light fraction was tested in a second reaction unit.

Example 5

A similar procedure was followed as in example 3, except that:

in this example, the reaction temperature of the first reaction unit was 395 ℃.

The remaining conditions were the same as in example 3.

The resulting secondary heavy components have major physicochemical properties at >350 ℃ as shown in table 3.

Example 6

A similar procedure was followed as in example 4, except that:

after the mixed raw material is subjected to fixed bed heavy oil hydrogenation treatment, the reaction temperature is increased by 3 ℃ every 30 days, and the operation is stopped after the hydrogenation test is carried out for 300 days in total.

The ore-rich precursor material 1 and the ore-rich precursor material 2 which are initially loaded into the reactor are reacted to become the V-rich material 1 and the vanadium-rich material 2, the V content of the V-rich material 1 and the V content of the vanadium-rich material 2 are respectively 62 mass percent and 53 mass percent after roasting analysis, and the V is the refined high-value V ₂O₅High quality material of (2).

Example 7

A similar procedure was followed as in example 3, except that: in the present embodiment, the first and second electrodes are,

second reaction unit: a small-sized catalytic cracking fixed fluidized bed test device is provided, wherein a catalytic cracking catalyst MLC-500 is used, the reaction temperature is 540 ℃, the catalyst-oil ratio is 8, and the retention time is 3 s; as a result, the product gasoline mass yield was 45% and the gasoline RON octane number was 93.

Example 8

A similar procedure to example 1 was used, except that: in the present embodiment, the first and second electrodes are,

and introducing the second heavy component into a delayed coking unit for reaction to obtain the coker gasoline, coker diesel oil and coker gas oil.

The operating conditions of the delayed coking unit were: the reaction temperature is 500 ℃ and the residence time is 0.5 h.

The sulfur content of the coker diesel was 0.16 mass%, the condensation point was-13 ℃, and the cetane number was 49.

The sulfur content of the coker gas oil was 0.76 mass%, and the condensation point was 32 ℃.

The yield of coker gasoline was 15%, the sulfur content was 0.08 mass%, and MON was 60.

And the coker diesel and coker gas oil were recycled to the third reaction unit for blending with LCO for hydroprocessing, the operating conditions of the third reaction unit being the same as in example 1.

First fractionation: the cut points of the first light fraction and the first heavy fraction 8 (properties see table 1) are 180 ℃;

A first reaction unit: the feedstock, DOA (from gravity reduction slag) and first heavy component 8 are mixed in a ratio of 1: 10 in a mass ratio, and the properties are shown in table 2; the treatment apparatus and catalyst loading were the same as in example 1, with the operating conditions: the reaction temperature is 360 ℃, the reaction pressure is 8MPa, and the liquid hourly space velocity is 0.3h^-1The volume ratio of hydrogen to oil is 800: 1. the product properties after hydrogenation of the mixed feedstock are shown in table 3.

Second fraction: and (3) fractionating the liquid-phase product obtained in the first reaction unit to obtain a second light component at a temperature of less than 350 ℃ and a second heavy component at a temperature of more than or equal to 350 ℃, wherein the properties of the second heavy component are shown in Table 4.

The second light fraction was tested in a second reaction unit.

Example 9

the second light fraction was tested in a second reaction unit.

Second reaction unit: the diesel oil hydrogenation modification device is sequentially filled with RS-2100: RHC-100 ═ 40: 60(V/V), the operating conditions are: the reaction temperature is 360 ℃, the reaction pressure is 8MPa, the volume ratio of hydrogen to oil is 800, and the liquid hourly space velocity is 1.2h^-1(ii) a The obtained diesel oil component has the sulfur content of 6ppm, the condensation point of-30 ℃ and the cetane number of 52.4.

Example 10

A similar procedure was followed as in example 1, except that the catalyst loading in the first reaction unit in this example was as follows:

according to the material flow direction, the catalyst filling sequence is a hydrogenation protection catalyst, an ore-rich precursor material 1, a hydrodemetallization desulfurization catalyst and a hydrodesulfurization catalyst. In the first reaction unit, the filling ratio among the catalysts is as follows: RG-30B: rich-ore precursor material 1: RDM-33B RCS-31 ═ 6: 60: 14: 20 (V/V).

The product properties after hydrogenation of the mixed feedstock are shown in table 3.

The second light fraction was tested in a second reaction unit.

Example 11

according to the material flow direction, the catalyst filling sequence is hydrogenation protection catalyst, ore-rich precursor material 2, ore-rich precursor material 1, hydrogenation demetalization desulfurization catalyst and hydrogenation desulfurization catalyst. In the first reaction unit, the filling ratio among the catalysts is as follows: RG-30B, rich mineral precursor material 2: rich-ore precursor material 1: RDM-33B RCS-31 ═ 6: 30: 30: 14: 20 (V/V).

The product properties after hydrogenation of the mixed feed are shown in Table 3.

The second light fraction was tested in a second reaction unit.

Example 12

according to the material flow direction, the filling sequence of the catalyst is hydrogenation protective agent, hydrogenation demetalization desulfurization catalyst and hydrogenation desulfurization catalyst, and the filling ratio among the catalysts is as follows: RG-30B: RDM-33B: RCS-31 ═ 10:40:50 (V/V).

The second light fraction was tested in a second reaction unit.

Example 13

according to the material flow direction, the filling sequence of the catalyst is hydrogenation protection catalyst, ore-rich precursor material 3, hydrogenation demetalization desulfurization catalyst and hydrogenation desulfurization catalyst, and the filling ratio among the catalysts is as follows: RG-30B: rich-ore precursor material 3: RDM-33B: RCS-31 ═ 6: 40: 24: 30 (V/V).

The second light fraction was tested in a second reaction unit.

Comparative example 1

A similar procedure was followed as in example 1, except that:

in this comparative example, the aromatic-rich distillate QY (aromatic content 20 mass%) was directly mixed with DOA without passing through a partial hydrogenation saturation treatment apparatus and then introduced into the same first reaction unit as in example 1. DOA and QY are mixed in a mass ratio of 1:10, and the properties of the mixed raw materials are shown in Table 2.

The product properties of the mixed feedstock after hydroprocessing in the first reaction unit are shown in table 3.

And (3) fractionating the liquid-phase product obtained by the treatment of the first reaction unit to obtain a second light component with the temperature of less than 350 ℃ and a second heavy component with the temperature of more than or equal to 350 ℃, wherein the properties of the second heavy component are shown in a table 4.

The second light fraction was tested in a second reaction unit.

Comparative example 2

A similar procedure was followed as in comparative example 1, except that:

in this comparative example, DOA and QY were mixed at a mass ratio of 2:10, and the properties of the mixed raw materials are shown in Table 2.

The product properties of the mixed feedstock after hydrotreatment in the first reaction unit are shown in table 3.

And (3) fractionating the liquid-phase product obtained by the treatment of the first reaction unit to obtain a second light component with the temperature of less than 350 ℃ and a second heavy component with the temperature of more than or equal to 350 ℃, wherein the properties of the second heavy component are shown in Table 4.

The second light fraction was tested in a second reaction unit.

Comparative example 3

A similar procedure was followed as in comparative example 1, except that:

in this comparative example, DOA and QY were mixed at a mass ratio of 3:10, and the next test could not be carried out because of a large amount of solids (100 ℃ C.) in the raw materials.

Table 1: before and after hydrogenation of aromatic-rich distillate oil

Table 2: mixed raw material properties

	Example 1	Example 2	Example 3	Example 4
					20 ℃ state	Liquid state	Liquid state	Liquid state	Liquid state
C₇Insoluble matter content by mass%	2.09	7.67	13.50	16.80
					Carbon residue, mass%	2.27	8.33	19.50	25.00
Sulfur, mass%	1.4	2.14	3.21	3.85
					Viscosity (100 ℃ C.), (mm)²/s)	1.9	8.6	35.1	36.0
Ni+V,(μg/g)	23	104	153	195

Table 2 (continuation): mixed raw material properties

	Example 8	Comparative example 1	Comparative example 2
				20 ℃ state	Liquid state	Liquid state	Liquid state
C₇Insoluble matter content by mass%	2.18	1.99	3.83
				Carbon residue, mass%	3.7	2.58	4.17
Sulfur, mass%	1.68	1.55	2.47
				Viscosity (100 ℃ C.), (mm)²/s)	3.9	3.1	5.6
Ni+V,(μg/g)	32	25	41

Table 3: the product property after hydrogenation of mixed raw material

Table 4: second heavy component Properties

Table 5: hydrocracking product Properties

Item	Density (20 deg.C), g/cm3	RON	Sulfur content, μ g/g
				Example 1	0.72	＞92	＜10
Example 2	0.72	＞92	＜10
				Example 3	0.72	＞92	＜10
Example 4	0.72	＞92	＜10
				Example 8	0.72	＞92	＜10
Example 10	0.72	＞92	＜10
				Example 11	0.72	＞92	＜10
Example 12	0.72	＞92	＜10
				Example 13	0.72	＞92	＜10
Comparative example 1	>0.72	<92	11
				Comparative example 2	>0.72	<92	12

Table 6: properties of ore-rich precursor material

	Burn and decrease in mass%	Specific surface area, m²/g	Water absorption, g/g
				Mineral-rich precursor material 1	13.5	263	1.08
Mineral-rich precursor material 2	29.9	279	1.22
				Mineral-rich precursor material 3	20.5	99	1.05

From the above results, it can be seen that the technology of the present invention is capable of producing high quality low sulfur marine fuel or low sulfur coke product feedstocks from DOA.

Moreover, the technology of the invention can obtain gasoline products with high quality and meeting national V standards.

The preferred embodiments of the present invention have been described above in detail, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, many simple modifications can be made to the technical solution of the invention, including combinations of various technical features in any other suitable way, and these simple modifications and combinations should also be regarded as the disclosure of the invention, and all fall within the scope of the invention.

Claims

1. A method of processing an aromatic-rich distillate oil, the method comprising:

(2) introducing deoiled asphalt and an aromatic hydrocarbon-containing stream containing the first heavy component into a first reaction unit for hydrogenation reaction, wherein the first reaction unit contains a mineral-rich precursor material and/or a hydrogenation catalyst, the first reaction unit is a fixed bed hydrogenation unit, the mineral-rich precursor material is a material capable of adsorbing at least one metal selected from V, Ni, Fe, Ca and Mg, the dosage ratio of the deoiled asphalt and the aromatic hydrocarbon-containing stream is such that a mixed raw material formed by the deoiled asphalt and the aromatic hydrocarbon-containing stream is in a liquid state at a temperature of not higher than 400 ℃, and the viscosity of the mixed raw material at 100 ℃ is not higher than 400mm ²/s；

2. The process of claim 1, wherein in step (2), the deoiled bitumen and the aromatic hydrocarbon-containing stream are used in a ratio such that the 100 ℃ viscosity of the mixed feedstock formed from the deoiled bitumen and the aromatic hydrocarbon-containing stream is no greater than 200mm²/s。

3. The process of claim 1, wherein in step (2), the deoiled bitumen and the aromatic hydrocarbon-containing stream are used in a ratio such that the 100 ℃ viscosity of the mixed feedstock formed from the deoiled bitumen and the aromatic hydrocarbon-containing stream is no greater than 100mm ²/s。

4. The method according to any one of claims 1 to 3, wherein in step (2), the aromatic hydrocarbon-containing stream further contains aromatic hydrocarbon compounds and/or aromatic hydrocarbon oils, and the aromatic hydrocarbon oils are selected from at least one of LCO, HCO, FGO, ethylene tar, coal tar, coker diesel oil, and coker gas oil.

5. The process according to claim 4, wherein the aromatic hydrocarbon compound is selected from benzene, toluene, xylene, naphthalene, a mixture of at least one C_1-6At least one of alkyl substituted naphthalene and aromatic hydrocarbon with more than three rings.

6. The method according to any one of claims 1 to 3, wherein the aromatic hydrocarbon content in the aromatic-rich distillate oil is 20 mass% or more.

7. The method according to any one of claims 1 to 3, wherein the aromatic content in the aromatic-rich distillate oil is 25 mass% or more.

8. The method according to any one of claims 1 to 3, wherein the aromatic hydrocarbon content in the aromatic-rich distillate oil is 40 mass% or more.

9. The method according to any one of claims 1 to 3, wherein in the step (2), the deoiled asphalt is deoiled asphalt obtained by subjecting a heavy oil raw material to a solvent deasphalting treatment in a solvent deasphalting unit.

10. The process of claim 9, wherein the mass fraction yield of the de-asphalted asphalt in the solvent deasphalting unit is no greater than 50%.

11. The process of claim 9, wherein the mass fraction of yield of the deoiled asphalt in the solvent deasphalting unit is not more than 40%.

12. The process of claim 9, wherein the mass fraction of yield of the deoiled asphalt in the solvent deasphalting unit is not more than 30%.

13. The process of any of claims 1-3, wherein in step (2), the deoiled asphalt and the aromatic hydrocarbon-containing stream are used in a mass ratio of 1:10 to 50: 10.

14. The process of any of claims 1-3, wherein in step (2), the deoiled asphalt and the aromatic hydrocarbon-containing stream are used in a mass ratio of 2:10 to 30: 10.

15. The process of any of claims 1-3, wherein in step (2), the deoiled asphalt and the aromatic hydrocarbon-containing stream are used in a mass ratio of 3:10 to 15: 10.

16. The method of any of claims 1-3, wherein the method further comprises: recycling the coker diesel and/or coker wax obtained in step (42) back to step (2) as at least part of the aromatics-containing stream.

17. The process according to any one of claims 1 to 3, wherein, in step (1), the third reaction unit is at least one of a fixed bed reactor, a moving bed reactor and an ebullating bed reactor.

18. The method of claim 17, wherein the operating conditions in the third reaction unit comprise: the reaction temperature is 200-420 ℃, the reaction pressure is 2-18MPa, and the liquid hourly space velocity is 0.3-10h^-1The volume ratio of hydrogen to oil is 50-5000.

19. The method of claim 17, wherein the operating conditions in the third reaction unit comprise: the reaction temperature is 220 ℃ and 400 ℃, the reaction pressure is 2-15MPa, and the liquid hourly space velocity is 0.3-5h^-1The volume ratio of hydrogen to oil is 50-4000.

20. The process according to any one of claims 1 to 3, wherein in step (2), the operating conditions in the first reaction unit comprise: the reaction temperature is 260-500 ℃, the reaction pressure is 2.0-20.0 MPa, the volume ratio of hydrogen to oil is 100-1200, and the liquid hourly space velocity is 0.1-1.5 h^-1。

21. The method of claim 20, wherein, in step (2), the operating conditions in the first reaction unit comprise: the reaction pressure is 2-12 MPa.

22. The method according to any one of claims 1 to 3, wherein in step (2), the mineral-rich precursor material contains a carrier and an active component element loaded on the carrier, wherein the carrier is selected from at least one of aluminum hydroxide, aluminum oxide and silicon oxide, and the active component element is selected from at least one of metal elements in groups VIB and VIII.

23. The method as claimed in claim 22, wherein, in the step (2), the rich-ore precursor material has a scorch reduction of not less than 3 mass% and a specific surface area of not less than 80m²(ii) water absorption of not less than 0.9 g/g.

24. The method according to claim 23, wherein in step (2), the first reaction unit is sequentially charged with a first ore-rich precursor material and a second ore-rich precursor material, in terms of a reactant flow direction, and a ignition loss of the second ore-rich precursor material is equal to or greater than that of the first ore-rich precursor material.

25. The method according to claim 24, wherein in step (2), the first mineral-rich precursor material has a burn reduction of 3-15 mass%, and the second mineral-rich precursor material has a burn reduction of not less than 15 mass%.

26. The method of claim 24, wherein the packing volume ratio of the first and second rich precursor materials is from 5:95 to 95: 5.

27. The process of any one of claims 1-3, wherein in step (41), the second reaction unit is a hydrocracking unit, and the operating conditions in the hydrocracking unit include: the reaction temperature is 360-420 ℃, the reaction pressure is 10.0-18.0 MPa, the volume ratio of hydrogen to oil is 600-2000, and the liquid hourly space velocity is 1.0-3.0 h^-1。

28. The process of claim 27, wherein the hydrocracking unit is packed with at least one hydrotreating catalyst and at least one hydrocracking catalyst.

29. The process of any one of claims 1-3, wherein in step (41), the second reaction unit is a catalytic cracking unit and the catalytic cracking unit is a fluid catalytic cracking unit.

30. The process of claim 29, wherein the operating conditions in the fluid catalytic cracking unit comprise: the reaction temperature is 500-600 ℃, the agent-oil ratio is 3-12, and the retention time is 0.6-6 s.

31. The method of any one of claims 1-3, wherein in step (41), the second reaction unit is a diesel hydro-upgrading unit and the operating conditions in the diesel hydro-upgrading unit include: the reaction temperature is 330-420 ℃, the reaction pressure is 5.0-18.0 MPa, the volume ratio of hydrogen to oil is 500-2000, and the liquid hourly space velocity is 0.3-3.0 h ^-1。

32. The method of claim 31, wherein the diesel hydro-upgrading unit is loaded with at least one diesel hydro-upgrading catalyst.

33. The process of any one of claims 1-3, wherein in step (42) the second heavy component is introduced into a delayed coking unit for reaction to yield at least one product selected from coker gasoline, coker diesel, coker gas oil, and low sulfur petroleum coke, and the operating conditions in the delayed coking unit comprise: the reaction temperature is 440-520 ℃, and the retention time is 0.1-4 h.

34. The process of any one of claims 1-3, wherein in step (42) the sulfur content of the second heavy component is no greater than 1.8 mass%, and the second heavy component is introduced into a delayed coking unit for reaction to obtain low sulfur petroleum coke.

35. The process of claim 34, wherein the low sulfur petroleum coke has a sulfur content of no greater than 3 mass%.

36. The method according to any one of claims 1 to 3, wherein in step (42) the second heavy component is used as a low sulphur bunker fuel oil component, and the sulphur content in the low sulphur bunker fuel oil component is not more than 0.5 mass%.

37. A system for processing an aromatic-rich distillate oil, wherein the system is adapted to perform the method of any one of claims 1 to 36, and the system comprises:

38. The system of claim 37, wherein the delayed coking unit is in fluid communication with the first reaction unit for recycling the coker gas oil and/or the coker wax oil obtained in the delayed coking unit back to the first reaction unit as at least a portion of the aromatic-containing stream.

39. The system of claim 37 or 38, further comprising a solvent deasphalting unit in fluid communication with said first reaction unit for subjecting a heavy oil feedstock to a solvent deasphalting process therein and introducing the resulting deasphalted asphalt after said solvent deasphalting process into said first reaction unit.