Method and system for processing aromatic-rich distillate oils

Abstract

The present invention relates to the technical field of hydrocarbon oil processing, and in particular to a method and system for processing aromatic-rich distillate oil. The method and system include: (1) introducing the aromatic-rich distillate oil into a fifth reaction unit for hydrogen saturation, followed by fractionation to provide a first light component and a first heavy component; (2) introducing the deoiled asphalt and an aromatic-containing stream containing the first heavy component into a hydrogen dissolution unit and mixing them with hydrogen, and introducing the mixed material into a first reaction unit for hydrogenation reaction, where the first reaction unit is a liquid-phase hydrogenation reaction unit; (3) fractionating the liquid-phase product from the first reaction unit to provide a second light component and a second heavy component; (41) introducing the second light component into a second reaction unit for reaction; (42) introducing the second heavy component into a delayed coking unit for reaction; or using the second heavy component as a component of low-sulfur marine fuel oil. The processing method provided by the present invention can realize high-value utilization of DOA.

Description

Detailed Description of the Invention

[0001] [Technical Field] The present invention relates to the field of processing hydrocarbon oils, and in particular to methods and systems for processing aromatic-rich distillate oils.

[0002] [Background technology] Highly efficient conversion of residual oil is the core of petroleum refining enterprises. Fixed-bed residual oil hydrogenation is an important technology for highly efficient conversion of residual oil, and has the characteristics of good product quality, maturation method, etc.

[0003] However, the high content of asphaltenes and metals in residual oil is a limiting factor for the run-time of residual oil hydrogenation on a fixed bed.

[0004] To address this issue, the SINOPEC Research Institute of Petroleum Processing (RIPP) developed a combined solvent deasphalting (demetallization) and hydrotreating catalytic cracking (SHF) technology for residual oil. This innovative technology maximizes the utilization of low-value vacuum residual oil and produces clean vehicle fuel that extends running time. However, the high softening point of deoiled asphaltene (DOA) makes it difficult to transport and use, limiting the widespread adoption of SHF technology.

[0005] A new combined process for producing propylene-rich products by hydrotreating and deep catalytic cracking (DCC) of residual oil is also limited by the influence of asphaltenes and metals in the residual oil. The hydrogen content of the hydrogenated residual oil is low, the operation period of residual oil hydrogenation is short, and the propylene yield from DCC is low, limiting the economic benefits of the combined technology.

[0006] Furthermore, new low-sulfur marine fuel standards, which require sulfur fractions to be no higher than 0.5% by weight, and low-sulfur petroleum coke standards, which require sulfur fractions to be no higher than 3.0% by weight, are scheduled to be implemented in 2020. Technology for inexpensively producing low-sulfur marine fuel (low-sulfur petroleum coke) is also an issue that must be urgently resolved.

[0007] Therefore, the conversion of DOA into a material for low sulfur marine fuel or low sulfur petroleum coke production is a technological challenge that needs to be addressed.

[0008] Summary of the Invention The object of the present invention is to provide a novel method for processing aromatic-rich distillate oils that allows for better hydrotreating results and long-term stable operation of the unit even at lower hydrogen partial pressures, lower hydrogen to oil ratios, and higher space velocities.

[0009] To achieve the above objectives, a first aspect of the present invention provides a method for processing an aromatic-rich distillate oil, the method comprising: (1) introducing the aromatics-rich distillate oil into a fifth reaction unit for hydrogen saturation, followed by fractionation to provide a first light component and a first heavy component, wherein the first light component and the first heavy component have a cut point of 100 to 250°C, and the aromatics content in the first heavy component is 20% by weight or more; (2) introducing the deoiled asphalt and the aromatics-containing stream containing the first heavy component into a hydrogen dissolution unit and mixing with hydrogen, and introducing the mixed material into a first reaction unit for hydrogenation reaction, the first reaction unit containing a mineral-rich precursor material and / or a hydrogenation catalyst, the first reaction unit being a liquid-phase hydrogenation reaction unit, the mineral-rich precursor material being a material capable of adsorbing at least one metal selected from V, Ni, Fe, Ca and Mg, wherein the deoiled asphalt and the aromatics-containing stream are used in an amount ratio such that the mixed feedstock formed by the deoiled asphalt and the aromatics-containing stream is in a liquid state at a temperature of 400°C or less; (3) fractionating the liquid-phase product from the first reaction unit to provide a second light component and a second heavy component, wherein the cut points for the second light component and the second heavy component are 240-450°C; (41) introducing the second light component into a second reaction unit for reaction to provide at least one product selected from a gasoline component, a diesel component, and a BTX feedstock component, wherein the second reaction unit is at least one selected from the group consisting of a hydrocracking unit, a catalytic cracking unit, and a diesel hydroupgrading unit; and (42) introducing the second heavy component into a delayed coking unit for reaction to provide at least one product selected from the group consisting of coker gasoline, coker diesel, coker wax oil, and low sulfur petroleum coke; or using the second heavy component as a component of low sulfur marine fuel oil.

[0010] A second aspect of the present invention provides a system for processing an aromatics-rich distillate oil, the system comprising: a third reaction unit for hydrogen saturation and fractionation on the aromatics-rich distillate oil to provide a first light component and a first heavy component; a hydrogen dissolving unit in fluid communication with the third reaction unit for mixing the deoiled asphalt and the aromatics-containing stream including the first heavy component from the third reaction unit with hydrogen therein; a first reaction unit in fluid communication with the hydrogen dissolving unit, the first reaction unit being a liquid-phase hydrogenation reaction unit used to carry out a hydrogenation reaction of the mixed material from the hydrogen dissolving unit 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 in the separation unit, the second reaction unit being at least one selected from the group consisting of a hydrocracking unit, a catalytic cracking unit, and a diesel hydroupgrading unit; a delayed coking unit in fluid communication with the separation unit for reacting therein the second heavy component obtained from the separation unit, the delayed coking unit providing at least one product selected from the group consisting of coker gasoline, coker diesel, coker wax oil, and low sulfur petroleum coke; and an outlet in fluid communication with the separation unit for discharging the second heavy fraction obtained from the separation unit from the system as a low sulfur marine fuel oil fraction.

[0011] When the method for processing aromatic-rich distillate oil according to the present invention is used to treat residual oil, the method can be carried out at a lower hydrogen partial pressure, a lower hydrogen-to-oil ratio, and a higher space velocity, and still achieve a relatively good hydrotreating effect and long-term stable operation of the unit.

[0012] The present invention is particularly suitable for the hydroconversion of atmospheric and vacuum residues. It is suitable for the hydroconversion of lean residual oils with high content of metals, high content of carbon residues, high content of fused ring materials and high content of nitrogen.

[0013] The present invention provides a process for de-oiled asphalt (DOA) hydroprocessing that can efficiently convert heavy oil and produce gasoline and BTX feedstock, as well as a system and method for the flexible production of low sulfur marine fuel and low sulfur petroleum coke.

[0014] BRIEF DESCRIPTION OF THE DRAWINGS FIG. 1 is a flow chart for processing aromatic-rich distillate oils according to a preferred embodiment of the present invention.

[0015] FIG. 2 is a flow chart for processing an aromatic-rich distillate oil according to a first embodiment of the present invention.

[0016] Explanation of symbols 1. Heavy oil feedstock 2. Solvent deasphalting unit 3 De-asphalted oil 4 De-oiled asphalt 5 Aromatics 6 Mixed feedstock 7 First reaction unit 8 Second light component 9. Second heavy component 10. Second reaction unit 11 delayed coking unit 12 BTX feedstock components 13 Gasoline components 14 Diesel components 15 Coker Gasoline 16 Coker Diesel 17 Coker waxy oil 18 Low sulfur petroleum coke 19 Separation unit 20 Aromatic-rich distillate 21 Third reaction unit 22 First heavy component 23 Hydrogen dissolution unit 24 Fourth reaction unit 25 DCC unit 26 Propylene 27 LCO 28 HCO 29 Slurry Oil [Embodiment] The endpoints of the ranges and any values ​​disclosed herein are not limited to the exact ranges or values, and these ranges or values ​​should be understood to encompass values ​​close to these ranges or values. For numerical ranges, each range between the endpoints and each individual point value, and each individual point value, may be combined with each other to create one or more new numerical ranges, and such new numerical ranges should be construed as specifically disclosed herein.

[0017] As noted above, a first aspect of the present invention provides a method for processing an aromatic-rich distillate oil, the method comprising: (1) introducing the aromatics-rich distillate oil into a fifth reaction unit for hydrogen saturation, followed by fractionation to provide a first light component and a first heavy component, wherein the first light component and the first heavy component have a cut point of 100 to 250°C, and the aromatics content in the first heavy component is 20% by weight or more; (2) introducing the deoiled asphalt and the aromatics-containing stream containing the first heavy component into a hydrogen dissolution unit and mixing with hydrogen, and introducing the mixed material into a first reaction unit for hydrogenation reaction, the first reaction unit containing a mineral-rich precursor material and / or a hydrogenation catalyst, the first reaction unit being a liquid-phase hydrogenation reaction unit, the mineral-rich precursor material being a material capable of adsorbing at least one metal selected from V, Ni, Fe, Ca and Mg, wherein the deoiled asphalt and the aromatics-containing stream are used in an amount ratio such that the mixed feedstock formed by the deoiled asphalt and the aromatics-containing stream is in a liquid state at a temperature of 400°C or less; (3) fractionating the liquid-phase product from the first reaction unit to provide a second light component and a second heavy component, wherein the cut points for the second light component and the second heavy component are 240-450°C; (41) introducing the second light component into a second reaction unit for reaction to provide at least one product selected from a gasoline component, a diesel component, and a BTX feedstock component, wherein the second reaction unit is at least one selected from the group consisting of a hydrocracking unit, a catalytic cracking unit, and a diesel hydroupgrading unit; and (42) introducing the second heavy component into a delayed coking unit for reaction to provide at least one product selected from the group consisting of coker gasoline, coker diesel, coker wax oil, and low sulfur petroleum coke; or using the second heavy component as a component of low sulfur marine fuel oil.

[0018] Preferably, the deoiled asphalt and aromatics-containing stream are used in an amount ratio such that the mixed feedstock formed from the deoiled asphalt and aromatics-containing stream is in a liquid state at a temperature of 280° C. or less. The deoiled asphalt and aromatics-containing stream are used in an amount ratio such that the mixed feedstock formed from the deoiled asphalt and aromatics-containing stream is in a liquid state at a temperature of 100° C. or less.

[0019] Preferably, the hydrogen saturation reaction carried out in the third reaction unit is partial hydrogen saturation, and particularly preferably the first light component and the first heavy component have a cut point of 180°C.

[0020] Preferably, the hydrogen dissolution unit of the present invention is operated under the conditions that the volume ratio of the amount of hydrogen supplied to the mixed feedstock formed by the deoiled asphalt and the aromatics-containing stream (i.e., the volume ratio of hydrogen to oil) is 30 to 200, more preferably 50 to 150, the operating temperature is 300 to 450°C, and the pressure is 2 to 20 MPa.

[0021] According to the method of the present invention, the mixed material obtained after mixing with hydrogen in the hydrogen dissolving unit can be supplied to the first reaction unit in either an upward flow mode or a downward flow mode. Preferably, the mixed material obtained after mixing with hydrogen in the hydrogen dissolving unit is supplied to the first reaction unit in an upward flow mode. As a result, the hydrogen dissolved and dispersed in the oil gathers together to form large bubbles, which are substantially prevented from escaping during the reaction. This provides a sufficient hydrogen source for the hydrogenation reaction, resulting in a better hydrotreating effect, further reducing the tendency of the catalyst to coke, maintaining higher catalytic activity of the catalyst, and extending the service life and stable operation period of the device.

[0022] The first light component is preferably fed to a catalytic cracking unit to produce low olefins. The specific operating conditions of the first light component fed to the catalytic cracking unit to produce low olefins are not particularly limited by the present invention.

[0023] Particularly preferably, the second light component and the second heavy component have a break point of 350°C.

[0024] Preferably, in step (2), the deoiled asphalt and aromatics-containing stream are mixed so that the viscosity of the mixed feedstock formed from the deoiled asphalt and aromatics-containing stream is 400 mm at 100°C. 2 / s or less, preferably 200 mm 2 / s or less, and even more preferably 100 mm 2 It is used at a ratio of less than / s.

[0025] In step (2), the aromatics-containing stream further comprises aromatic hydrocarbons and / or aromatic oils, and the aromatic oils are at least one selected from the group consisting of LCO, HCO, FGO (catalytic heavy oil), ethylene tar, coal tar, coker diesel, and coker wax oil.

[0026] Preferably, the aromatic hydrocarbon is one or more selected from benzene, toluene, xylene, naphthalene, methylnaphthalene, polybranched naphthalene, and aromatic hydrocarbons having more than two rings, and preferably polycyclic aromatic hydrocarbons having three or fewer rings, or mixtures thereof. Particularly preferably, the aromatic hydrocarbon is selected from benzene, toluene, xylene, naphthalene, at least one C 1-6 At least one selected from the group consisting of naphthalene substituted with an alkyl group, and tricyclic or higher aromatic hydrocarbons.

[0027] More preferably, the aromatic hydrocarbon content in the aromatic-rich distillate oil is 20% by weight or more, preferably 25% by weight or more, preferably 40% by weight or more, and more preferably 60% by weight or more.

[0028] Preferably, in step (2), the deoiled asphalt is obtained by subjecting the heavy oil feedstock to a solvent deasphalting process in a solvent deasphalting unit.

[0029] Preferably, in the solvent deasphalting unit, the yield of deoiled asphalt is 50 wt% or less, more preferably 40 wt% or less, and even more preferably 30 wt% or less.

[0030] According to a preferred embodiment, in step (2), the aromatics-containing stream is an aromatics-rich distillate oil, and the weight ratio of the amount of deoiled asphalt to the amount of aromatics-containing stream is 1:10 to 50:10, more preferably 2:10 to 30:10, more preferably 3:10 to 15:10.

[0031] Preferably, the process of the present invention further comprises recycling the coker diesel and / or coker gas oil obtained in step (42) to the first reaction unit of step (1) for hydrogen saturation.

[0032] Preferably, in step (1), the third reaction unit is at least one of a fixed bed reactor, a moving bed reactor, and an ebullated bed reactor.

[0033] Preferably, the third reaction unit has a reaction temperature of 200 to 420°C, a reaction pressure of 2 to 18 MPa, and a liquid hourly space velocity of 0.3 to 10 h -1 and the volume ratio of hydrogen to oil is 50 to 5000. More preferably, the third reaction unit is operated under conditions where the reaction temperature is 220 to 400°C, the reaction pressure is 2 to 15 MPa, and the liquid hourly space velocity is 0.3 to 5 h -1 and is operated under conditions where the volume ratio of hydrogen to oil is 50 to 4000.

[0034] A preferred embodiment of the third reaction unit of the fifth modified example is shown below.

[0035] Partial hydrogenation of aromatic-rich distillate oils in the presence of hydrogen is generally carried out under fixed-bed, ebullated-bed, or moving-bed hydrotreating conditions. Taking current industrial fixed-bed diesel or waxy oil hydrogenation technology as an example, the reactor or reaction bed contains at least a hydrotreating catalyst. The hydrotreating catalyst used for partial hydrogenation of aromatic-rich distillate oils preferably has good and moderate hydrogenation activity to avoid further saturation of tetralin-like structures into decahydronaphthalene or cycloalkane structures, which have lower hydrogen-donating capacity. These catalysts generally use porous refractory inorganic oxides, such as alumina or molecular sieves, as supports, and oxides or sulfides of metals from Group VIB and / or Group VIII, such as W, Mo, Co, and Ni, as active components, with optional addition of various other additives, such as elements such as P, Si, F, and B. For example, the RS series pretreatment catalysts developed by RIPP are NiMo catalysts.

[0036] The first reaction unit is particularly preferably a residual oil liquid phase hydrogenation reactor.

[0037] Preferably, in step (2), the first reaction unit has a reaction temperature of 260 to 500°C, a reaction pressure of 2.0 to 20.0 MPa, a volume ratio of recycled oil to crude oil at the inlet of the first reaction unit of 0.1:1 to 15:1, and a liquid hourly space velocity of 0.1 to 1.5 h -1 The reactor is operated under the condition of 0.05 to 0.05. The liquid hourly space velocity and reaction pressure are selected according to the properties of the material to be treated and the desired conversion and refining depth. The mixed feedstock formed by the deoiled asphalt and aromatics-containing stream can be mixed with hydrogen and then passed through the catalyst bed layer from top to bottom and fed into the top of the reactor of the first reaction unit; or the catalyst is fed into the bottom of the reactor of the first reaction unit and passed through the catalyst bed layer from bottom to top.

[0038] Preferably, in step (2), the mineral-rich precursor material comprises a support and an active component element supported on the support, wherein the support is at least one selected from the group consisting of aluminum hydroxide, alumina, and silica, and the active component element is at least one metal element selected from the group consisting of 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.

[0039] Preferably, in step (2), the mineral-rich precursor material has an ignition loss of 3% by weight or more and a specific surface area of ​​80 m 2 / g or more, and the water absorption is 0.9 g / g or more. Loss on ignition refers to the weight loss of the mineral-rich precursor material after calcination at 600°C for 2 hours compared to the weight before calcination. Water absorption refers to the weight increase of the mineral-rich precursor material after immersion in water at room temperature (e.g., 25°C) for 30 minutes compared to the weight before immersion.

[0040] According to a preferred embodiment, in step (2), the first reaction unit is successively filled with a first mineral-rich precursor material and a second mineral-rich precursor material following the flow direction of the reactants, the second mineral-rich precursor material having an ignition loss equal to or greater than that of the first mineral-rich precursor material.

[0041] According to the above preferred embodiment, more preferably, the first mineral-rich precursor material has a loss on ignition of 3 to 15 wt %, and the second mineral-rich precursor material has a loss on ignition of 15 wt % or more.

[0042] According to the above-mentioned preferred embodiment, more preferably, the first mineral-rich precursor material and the second mineral-rich precursor material are charged in a volume ratio of 5:95 to 95:5.

[0043] The hydrogenation catalyst of the present invention may be a graded combination of different catalysts, preferably the hydrogenation catalyst is capable of catalyzing at least the hydrodemetallization and hydrodesulfurization reactions.

[0044] According to the present invention, the specific type of catalyst capable of catalyzing the hydrodemetallization reaction, hydrodesulfurization reaction, hydrodeasphalting reaction, and hydrodecarburization reaction is not particularly limited, and any catalyst that is capable of catalyzing the above reactions and has been conventionally used in the relevant field may be used.

[0045] The hydrogenation catalyst of the present invention may use, for example, a porous refractory inorganic oxide as a support, and an oxide or sulfide of a metal of Group VIB and / or Group VIII as an active component, and may further optionally contain an auxiliary agent.

[0046] Preferably, after a long period of operation of the first reaction unit, the mineral-rich precursor material is converted into a vanadium-rich material, and the vanadium content in the vanadium-rich material is 10% by weight or more; particularly preferably, the ore-rich precursor material is converted into a vanadium-rich material having a V content of 20% by weight or more, from which high-value V2O5 can be directly refined.

[0047] A preferred implementation of the first reaction unit of the present invention is given below.

[0048] The technology for hydrotreating feedstock contained in the first reaction unit of the present invention is a liquid-phase hydrotreating technology. The reactor or reaction bed contains at least a mineral-rich precursor material and / or a hydrogenation catalyst. The mineral-rich precursor material is mainly composed of two parts: a support with a strong ability to adsorb vanadium-containing organic compounds in oil, and an active component with hydrogenation activity. The support is mainly obtained by extruding, molding, and drying silica, aluminum hydroxide, or a mixture of aluminum hydroxide and alumina. The surface of the support is rich in -OH. This support has a strong adsorption ability for vanadium-containing organic compounds in oil. This support has an ignition loss of 5 wt% or more after calcination at 600°C for 2 hours. The active component mainly contains oxides or sulfides of metals from Group VIB and / or Group VIII, such as W, Mo, Co, and Ni.

[0049] The hydrogenation catalyst for the preferred embodiment described above is generally a heavy residue hydrogenation catalyst. A heavy residue hydrogenation catalyst refers to a composite catalyst capable of hydrodemetallizing, hydrodesulfurizing, and hydrodecarburizing heavy residues. These catalysts generally use a porous refractory inorganic oxide such as alumina as a support, and an oxide or sulfide of a metal from Group VIB and / or Group VIII, such as W, Mo, Co, or Ni, as the active component. Various other additives, such as P, Si, F, and B, are optionally added. Examples include the RDM and RCS series heavy metal and residual oil hydrodemetallization and desulfurization catalysts developed by RIPP. Currently, liquid-phase residual oil hydrogenation technology often uses multiple catalysts in combination. In the present invention, a mineral-rich precursor material, a hydrodemetallization and desulfurization catalyst, and a hydrodesulfurization catalyst are preferably used. These are generally loaded in this order so that the feedstock is successively contacted with the mineral-rich precursor material, the hydrodemetallization and desulfurization catalyst, and the hydrodesulfurization catalyst. Of course, there are techniques for loading mixtures of these catalysts.

[0050] According to a preferred embodiment, in step (41), the second reaction unit is a hydrocracking unit, and the reaction temperature is 360 to 420°C, the reaction pressure is 10.0 to 18.0 MPa, the volume ratio of hydrogen to oil is 600 to 2000, and the liquid hourly space velocity is 1.0 to 3.0 h -1 The operation is carried out under the conditions:

[0051] Preferably, the hydrocracking unit is loaded with at least one hydrotreating catalyst and at least one hydrocracking catalyst.

[0052] Preferably, the hydrocracking unit is a fixed bed hydrocracking unit.

[0053] When the second reaction unit is a hydrocracking unit, preferred embodiments for the second reaction unit of the present invention are provided below.

[0054] In step (41), the second light component is introduced into a second reaction unit for reaction using fixed-bed hydrocracking technology. Taking the conventional industrial fixed-bed hydrocracking of waxy oils as an example, the reactor or reaction bed contains at least two hydrocracking catalysts: a pretreatment catalyst and a hydrocracking catalyst. Because the material obtained from liquid-phase hydrotreating and subsequent fractionation has high metal, sulfur, and nitrogen contents and a high carbon residue, the pretreatment catalyst preferably has strong demetalization activity and good desulfurization and denitrification activities 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 generally use porous refractory inorganic oxides such as alumina or molecular sieves as supports, oxides or sulfides of Group VIB and / or Group VIII metals such as W, Mo, Co, and Ni as active components, and optionally contain various other additives such as elements such as P, Si, F, and B. For example, the RS series pretreatment catalysts and RHC series hydrocracking catalysts developed by RIPP are used. The RS series catalysts are NiW catalysts, and the RHC series catalysts are NiMo molecular sieve catalysts.

[0055] According to another preferred embodiment, in step (41), the second reaction unit is a catalytic cracking unit, and the catalytic cracking unit is a fluid catalytic cracking (FCC) unit.

[0056] According to another preferred embodiment, a second light component fluid catalytic cracking (FCC) technology, preferably that used in the LTAG technology developed by RIPP, is used to produce mainly gasoline fractions and liquefied gases.

[0057] Preferably, the fluid catalytic cracker is operated under conditions where the reaction temperature is 500-600°C, the catalyst to oil ratio is 3-12, and the residence time is 0.6-6 seconds.

[0058] The catalyst to oil ratios of the present invention refer to the weight ratio of catalyst to oil unless otherwise specified.

[0059] According to another preferred embodiment, in step (41), the second reaction unit is a diesel hydrogenation upgrading unit, and the reaction temperature is 330-420°C, 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 The engine is operated under the following conditions:

[0060] Preferably, the diesel hydro-upgrading unit is loaded with at least one diesel hydro-upgrading catalyst.

[0061] The diesel hydroupgrading catalyst may be the RS series pretreatment catalyst and the RHC-100 series diesel hydrocracking catalyst developed by RIPP.

[0062] According to a preferred embodiment, in step (42), the second heavy component is introduced into a delayed coking unit for reaction to produce at least one product selected from coker gasoline, coker diesel, coker wax oil, and low-sulfur petroleum coke. The delayed coking unit is operated under conditions of a reaction temperature of 440 to 520°C and a residence time of 0.1 to 4 hours.

[0063] According to another preferred embodiment, in step (42), the second heavy component has a sulfur content of 1.8 wt% or less, and the second heavy component is introduced into a delayed coking unit for reaction to provide a low-sulfur petroleum coke. More preferably, the low-sulfur petroleum coke has a sulfur content of 3 wt% or less.

[0064] Preferably, in step (42), the second heavy component is used as a low sulfur marine fuel oil component, and the conditions are controlled so that the sulfur content of the low sulfur marine fuel oil component is 0.5 wt% or less.

[0065] According to the present invention, the specific operation of the solvent deasphalting treatment is not particularly limited, and conventional solvent deasphalting treatment can be used. The operating parameters of the solvent deasphalting method are exemplified in the examples of the present invention, but these should not be construed as limitations of the present invention by those skilled in the art.

[0066] The process of the present invention is suitable for the hydroconversion of atmospheric and vacuum residues, especially lean residual oils with high metal contents (Ni+V>150 μg / g, especially Ni+V>200 μg / g), high carbon residue contents (carbon residue weight fraction>17%, especially carbon residue weight fraction>20%) and high contents of fused ring substances.

[0067] As mentioned above, a second aspect of the present invention provides a system for processing an aromatics-rich distillate oil, the system comprising: a third reaction unit for hydrogen saturation and fractionation on the aromatics-rich distillate oil to provide a first light component and a first heavy component; a hydrogen dissolving unit in fluid communication with the third reaction unit for mixing the deoiled asphalt and the aromatics-containing stream containing the first heavy component from the third reaction unit with hydrogen therein; a first reaction unit in fluid communication with the hydrogen dissolving unit, which is a liquid-phase hydrogenation reaction unit used to carry out a hydrogenation reaction of the mixed material from the hydrogen dissolving unit therein; a separation unit in fluid communication with the first reaction unit for fractionating the liquid-phase product from the first reaction unit therein; and a second light component obtained in the separation unit for fractionating the second light component therein. a second reaction unit in fluid communication with the separation unit for reacting the second heavy component obtained from the separation unit in a delayed coking unit, the delayed coking unit being at least one product selected from the group consisting of coker gasoline, coker diesel, coker wax oil, and low sulfur petroleum coke; and an outlet in fluid communication with the separation unit for discharging the second heavy fraction obtained from the separation unit from the system as a low sulfur marine fuel oil fraction.

[0068] Preferably, the delayed coking unit is in fluid communication with the hydrogen dissolution unit and the coker gas oil and / or coker gas oil obtained in the delayed coking unit is recycled back to the first reaction unit.

[0069] Preferably, the system further comprises a solvent deasphalting unit in fluid communication with the hydrogen dissolution unit, the solvent deasphalting unit being used for solvent deasphalting the heavy oil feedstock therein and for introducing the deasphalted asphalt obtained after solvent deasphalting into the hydrogen dissolution unit.

[0070] According to a preferred embodiment, in the system of the present invention, the second reaction unit is a hydrocracking unit.

[0071] 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.

[0072] According to another preferred embodiment, in the system of the present invention, the second reaction unit is a diesel hydro-upgrading unit.

[0073] The present invention also provides a first variant of the method, which includes: (11) introducing heavy crude oil into a solvent deasphalting unit for solvent deasphalting treatment to provide deoiled asphalt and deasphalted oil; (12) introducing the deasphalted oil into a fourth hydrogenation unit for a hydrogenation reaction, and introducing a liquid-phase effluent obtained in the fourth hydrogenation unit into a DCC unit for reaction to provide propylene, LCO, HCO, and a slurry oil, wherein the fourth hydrogenation unit is a fixed-bed hydrogenation unit; and (1) using an aromatics-rich distillate containing LCO and / or HCO from the DCC unit as the aromatics-rich distillate in step (1).

[0074] In this first variant, preferably, the process of the present invention further comprises recycling the coker diesel and / or coker gas oil obtained in step (42) to a third reaction unit for hydrogen saturation.

[0075] Preferably, in step (12), the fourth reaction unit has a reaction temperature of 280 to 400°C, a reaction pressure of 6.0 to 14.0 MPa, a volume ratio of hydrogen to oil of 600 to 1200, and a liquid hourly space velocity of 0.3 to 2.0 h -1 The engine is operated under the following conditions:

[0076] Preferably, in step (12), at least two hydrogenation catalysts are loaded into the fourth reaction unit. More preferably, in step (12), the hydrogenation catalyst is a catalyst capable of catalyzing at least one reaction selected from the group consisting of hydrodemetallization, hydrodesulfurization, and hydrodecarburization. The hydrogenation catalyst is generally supported on a porous refractory inorganic oxide such as alumina. Particularly preferably, in step (12), the hydrogenation catalyst comprises alumina as a support and a Group VIB and / or Group VIII metal element as an active component element, and optionally also comprises at least one auxiliary element selected from P, Si, F, and B. In the hydrogenation catalyst, the Group VIB and Group VIII metal element may be, for example, W, Mo, Co, Ni, etc. In the hydrogenation catalyst, the active component may be an oxide and / or sulfide of the above-mentioned active component element.

[0077] A preferred embodiment of the fourth reaction unit of the present invention is shown below: The third hydrogenation unit for deasphalted oil (DAO) in the presence of hydrogen is generally as follows: the DAO hydrotreatment technology is a fixed-bed hydrotreatment technology. Taking current industrial fixed-bed heavy oil and residual oil hydrogenation technology as an example, the reactor or reaction bed layer contains at least two hydrogenation catalysts. The heavy oil and residual oil hydrogenation catalyst refers to a composite catalyst that has functions such as hydrodemetalization, hydrodesulfurization, hydrodenitrification, and hydrodecarburization for both heavy oil and residual oil. These catalysts generally use a porous refractory inorganic oxide such as alumina as a support, and an oxide or sulfide of a Group VIB and / or Group VIII metal, such as W, Mo, Co, or Ni, as the active component, with various other additives such as elements such as P, Si, F, and B optionally added. For example, the RDM and RCS series heavy metal and residual oil hydrodemetalization and desulfurization catalysts developed by RIPP are used. Current fixed-bed residual oil hydrogenation technology often uses multiple catalysts in combination. The hydrodemetalization, hydrodesulfurization, and hydrodenitrification catalysts are typically loaded in a sequential order, with crude oil contacting the hydrodemetalization, hydrodesulfurization, and hydrodenitrification catalysts sequentially. Depending on the circumstances, one or two catalysts may be absent. For example, only the hydrodemetalization and hydrodesulfurization catalysts may be loaded, but the hydrodenitrogenation catalyst may not. Of course, some technologies load these catalysts as a mixture.

[0078] The method of processing aromatic-rich distillate oils according to the present invention is described in further detail below with reference to FIGS.

[0079] As shown in FIG. 1 , aromatic-rich distillate oil 20 is fed to a third reaction unit 21 for hydrogen saturation and subsequently fractionated to provide a first light component and a first heavy component 22. Heavy oil feedstock 1 is fed to a solvent deasphalting unit 2 for solvent deasphalting to provide a deoiled asphalt 4 and a deasphalted oil 3. The deoiled asphalt 4 and an aromatics-containing stream containing the first heavy component 22 are mixed to form a mixed feedstock 6, which is mixed with hydrogen in a hydrogen dissolution unit 23. The resulting mixed material is fed to a first reaction unit 7 for a hydrogenation reaction. Here, the aromatics-containing stream preferably also contains external aromatic hydrocarbons 5, and the first reaction unit contains a mineral-rich precursor material and a hydrogenation catalyst capable of catalyzing at least one reaction selected from hydrodemetallization, hydrodesulfurization, hydrodeasphalting, and hydrodecarbonization. The first reaction unit is a liquid-phase hydrogenation reaction unit. The liquid-phase product from the first reaction unit 7 is fed to a separation unit 19 for fractionation to provide a second light component 8 and a second heavy component 9. The second light component and the second heavy component have cut points between 240°C and 450°C. The second light component 8 is fed to a second reaction unit 10 for reaction to provide at least one product selected from a gasoline component 13, a BTX feedstock component 12, and a diesel component 14. The second reaction unit is at least one selected from a hydrocracking unit, a catalytic cracking unit, and a diesel hydroupgrading unit. The second heavy component 9 is fed to a delayed coking unit 11 for reaction to provide at least one product selected from the group consisting of coker gasoline 15, coker diesel 16, coker waxy oil 17, and low-sulfur petroleum coke 18. Alternatively, the second heavy component 9 is used as a low-sulfur marine fuel oil component.

[0080] As shown in FIG. 2 , a heavy oil feedstock 1 is supplied to a solvent deasphalting unit 2 for solvent deasphalting to provide a deoiled asphalt 4 and a deasphalted oil 3. The deasphalted oil 3 is supplied to a fourth reaction unit 24 for a hydrogenation reaction, and the liquid-phase effluent obtained in the fourth reaction unit 24 is supplied to a DCC unit 25 for a reaction to provide propylene 26, LCO 27, HCO 28, and / or slurry oil 29. An aromatic-rich distillate oil 20 containing LCO 27 and / or HCO 28 from the DCC unit is supplied to a third reaction unit 21 for hydrogen saturation and subsequently fractionated to provide a first heavy component 22 and a first light component. A mixed feedstock 6 formed from the deoiled asphalt 4 and an aromatic-containing stream containing the first heavy component 22 is supplied to a first reaction unit 7 for a hydrogenation reaction, and the aromatic-containing stream also preferably contains aromatic hydrocarbons 5 from outside. The first reaction unit 7 contains a mineral-rich precursor material and a hydrogenation catalyst capable of catalyzing at least one reaction selected from hydrodemetallization, hydrodesulfurization, hydrodeasphalting, and hydrodecarbonization. The liquid-phase product from the first reaction unit 7 is fed to a separation unit 19 for fractionation to provide a second light component 8 and a second heavy component 9. The second light component 8 is fed to a second reaction unit 10 for reaction to provide at least one product selected from the group consisting of a gasoline component 13, a BTX feedstock component 12, and a diesel component 14, or the second light component 8 is recycled back to the DCC unit 25. The second heavy component 9 is fed to a delayed coking unit 11 for reaction to provide at least one product selected from the group consisting of coker gasoline 15, coker diesel 16, coker waxy oil 17, and low-sulfur petroleum coke 18, or the second heavy component 9 is used as a low-sulfur marine fuel oil component.

[0081] The technology of the present invention allows for the efficient conversion of heavy oil to create systems and methods for the flexible production of gasoline and BTX feedstocks, as well as low sulfur marine fuel and low sulfur petroleum coke.

[0082] Compared with the prior art, the present invention preferably applies an effective combination of methods such as residual oil hydrogenation, hydrocracking, or catalytic cracking, so as to convert low-value DOA into low-sulfur marine fuel components and low-sulfur petroleum coke feedstock, and meet the requirements of environmental protection, thereby realizing high efficiency, environmental protection, and comprehensive utilization of heavy oil resources.

[0083] In addition, the technology provided by the present invention allows DOA to be efficiently converted in a residual liquid phase hydrotreating reactor to produce gasoline fractions and BTX feedstock, which can provide a feedstock for producing low sulfur marine fuel and low sulfur coke products.

[0084] The present invention will now be described in detail with reference to the following examples. Unless otherwise specified, the following examples were carried out using the process flow shown in Figure 1. Unless specifically stated, the following examples have the following general characteristics: The results in Tables I-3 and II-4 in the following Examples are average values ​​obtained from sampling tests every 25 hours over a 100-hour continuous run of the equipment, unless otherwise specified.

[0085] The saturation experiments of aromatics-rich fraction partial hydrogenation were carried out in a medium-scale fixed-bed diesel hydrotreater, with a total reactor volume of 200 mL. In the following examples, the hydrogenation catalysts and materials used for partial hydrogenation of aromatics-rich fraction oil were RS-2100 series hydrogenation catalysts developed by RIPP.

[0086] The liquid-phase stream obtained by partial hydrosaturation was fractionated to obtain a first light component and a first heavy component with a cutoff point of 180°C, where the first heavy component and DOA formed a mixed feedstock. The mixed feedstock was subjected to a hydrotreating test in a medium-scale heavy oil liquid-phase hydrotreating unit, with a total reactor volume of 200 mL. In the following examples, the materials used in the hydrotreating catalyst and the first reaction unit were the RG-30B protected catalyst developed by RIPP, and mineral-rich precursor material 1, mineral-rich precursor material 2, RDM-33B residue demetallization desulfurization transition catalyst, and RCS-31 desulfurization catalyst developed by the Petrochemical Engineering Science Research Institute. The hydrotreating catalyst, mineral-rich precursor material 1, mineral-rich precursor material 2, hydrodemetallization and desulfurization catalyst, and hydrodesulfurization catalyst were sequentially loaded according to the flow direction of the reactants. In the first reaction unit, the loading ratio between the catalysts was as follows: RG-30B: mineral-rich precursor material 1: mineral-rich precursor material 2: RDM-33B: RCS-31 = 6:30:30:14:20 (V / V).

[0087] The second reaction unit was a fixed-bed hydrocracker, and the catalysts used were the RS-2100 refinery catalyst and the RHC-131 hydrocracking catalyst developed by RIPP. The fixed-bed hydrocracker unit had a reaction temperature of 370°C in the refinery section, 385°C in the cracking section, a reaction pressure of 10 MPa, and a liquid hourly space velocity of 2.0 h . -1 and a volume ratio of hydrogen to oil of 1200:1.

[0088] [Example A] Preparation of Mineral-Rich Precursor Material 1: 2000 g of RPB110 pseudoboehmite (manufactured by the CHANGLING DIVISION of SINOPEC CATALYST CO., LTD.) was used. 1000 g of the pseudoboehmite was treated at 550 °C for 2 hours to produce approximately 700 g of alumina. Approximately 700 g of alumina and another 1000 g of pseudoboehmite were thoroughly mixed, and then 40 g of sesbania powder and 20 g of citric acid were added. 2200 g of deionized water was then added. The mixture was then kneaded, extruded into strips, and dried at 300 °C for 3 hours to obtain approximately 1730 g of support. This support was saturated with 2100 mL of a solution containing Mo and Ni. The Mo content in the solution was 5.5 wt. % calculated as MoO3, and the Ni content was 1.5 wt. % calculated as NiO. After 30 minutes of impregnation and then treatment at 180°C for 4 hours, mineral-rich precursor material 1 was obtained, the properties of which are shown in Table I-6.

[0089] Preparation of Mineral-Rich Precursor Material 2: 2000 g of RPB220 pseudoboehmite (manufactured by the CHANGLING DIVISION of SINOPEC CATALYST CO., LTD.) was used, and 30 g of sesbania powder and 30 g of citric acid were added. 2400 g of deionized water was then added. The mixture was then kneaded, extruded into strips, and dried at 120°C for 5 hours to obtain approximately 2040 g of support. This support was then saturated with 2200 mL of a solution containing Mo and Ni. The Mo content, calculated as MoO3, was 7.5 wt. % and the Ni content, calculated as NiO, was 1.7 wt. After 30 minutes of impregnation, the solution was treated at 200°C for 3 hours to obtain Mineral-Rich Precursor Material 2, the properties of which are listed in Table I-6.

[0090] Preparation of Mineral-Rich Precursor Material 3: 2000 g of commercially available silica was used, and 30 g of sesbania powder and 30 g of sodium hydroxide were added. 2400 g of deionized water was then added. The mixture was then kneaded, extruded into a strip, and dried at 120°C for 5 hours to obtain a support. 2200 mL of a solution containing Mo and Ni was added to the support for saturation impregnation. The Mo content in the solution was 4.5 wt. % calculated as MoO3, and the Ni content was 1.0 wt. % calculated as NiO. After 30 minutes of impregnation, the solution was treated at 200°C for 3 hours to obtain Mineral-Rich Precursor Material 3, the properties of which are listed in Table I-6.

[0091] [Example I-1] In this example, LCO from the RLG plant of Shanghai Petrochemical was used as the aromatics-rich distillate oil, and LCO hydrogenation was carried out at a reaction temperature of 290°C, a reaction pressure of 4 MPa, and a liquid hourly space velocity of 1 h -1 and a volume ratio of hydrogen to oil of 800:1.

[0092] The properties of LCO and first heavy component 1 are shown in Table I-1.

[0093] The DOA from the vacuum residue was blended with the first heavy component 1 in a weight ratio of 1:10, and the properties of the blended feedstock are shown in Table I-2.

[0094] First, the mixed feedstock of DOA and the first heavy component 1 was mixed with hydrogen in a hydrocracking unit (the volume ratio of the hydrogen supply to the mixed feedstock formed by deoiled asphalt and the first heavy component was 100, the operating temperature of the hydrocracking unit was 320°C, and the pressure was 10 MPa), and the resulting mixed material was fed to the first reactor. The first reactor had a reaction temperature of 360°C, a reaction pressure of 10 MPa, and a liquid hourly space velocity of 0.6 h . -1 The first reaction unit was operated under conditions where the volume ratio of recycled oil to crude oil at the inlet was 0.5:1. After hydrogenation, the properties of the mixed feedstock were shown in Table I-3.

[0095] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0096] The second light component at a temperature below 350°C was tested in the second reaction unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0097] [Example I-2] HCO from a catalytic cracking unit of Shanghai Petrochemical was used as the aromatic-rich distillate in this example, where HCO hydrogenation was carried out at a reaction temperature of 330°C, a reaction pressure of 6 MPa, and a liquid hourly space velocity of 1 h -1 and a volume ratio of hydrogen to oil of 800:1.

[0098] The properties of HCO and first heavy component 2 are shown in Table I-1.

[0099] The DOA from the vacuum residue was blended with the first heavy component 2 in a weight ratio of 5:10, and the properties of the blended feedstock are shown in Table I-2.

[0100] First, the first heavy component 2, a mixed feedstock of DOA and hydrogenated HCO, was mixed with hydrogen in a hydrocracking unit (the volume ratio of the hydrogen supply to the mixed feedstock formed by deoiled asphalt and the first heavy component 2 was 100, the operating temperature of the hydrocracking unit was 320°C, and the pressure was 10 MPa), and the resulting mixed material was supplied to the first reactor. The first reactor had a reaction temperature of 380°C, a reaction pressure of 10 MPa, and a liquid hourly space velocity of 0.6 h . -1 The volume ratio of recycled oil to crude oil at the inlet of the first reaction unit was 0.5:1. After hydrogenation, the properties of the mixed feedstock were shown in Table I-3.

[0101] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0102] The second light component at a temperature below 350°C was tested in the second reaction unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0103] [Example I-3] In this example, the same LCO as in Example I-1 was used as the aromatic-rich distillate oil, and the reaction temperature was 320°C, the reaction pressure was 6 MPa, and the liquid hourly space velocity was 1 h -1 LCO hydrogenation was carried out under conditions where the volume ratio of hydrogen to oil was 800:1.

[0104] The properties of LCO and first heavy component 3 are shown in Table I-1. The properties of LCO and first heavy component 3 are shown in Table I-1.

[0105] DOA was derived from vacuum residue and mixed with first heavy component 3 in a weight ratio of 10:10, and the properties of the mixed feedstock are shown in Table I-2.

[0106] First, the mixed feedstock of DOA and the first heavy component 3 was mixed with hydrogen in a hydrocracking unit (the volume ratio of the hydrogen supply to the mixed feedstock formed by deoiled asphalt and the first heavy component 3 was 100, the operating temperature of the hydrocracking unit was 320°C, and the pressure was 8 MPa), and the resulting mixed material was fed to the first reactor. The first reactor had a reaction temperature of 370°C, a reaction pressure of 8 MPa, and a liquid hourly space velocity of 0.6 h . -1 The volume ratio of recycled oil to crude oil at the inlet of the first reaction unit was 0.5:1. After hydrogenation, the properties of the mixed feedstock were shown in Table I-3.

[0107] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0108] The second heavy component was subjected to a coking reaction at a reaction temperature of 500°C for 0.5 hours to obtain a petroleum coke (yield 32% by weight) having a sulfur content of 2.7% by weight.

[0109] The second light component at a temperature below 350°C was tested in the second reaction unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0110] [Example I-4] The aromatic-rich distillate oil used in Example I-4 was coal tar from a coal tar unit in China. The coal tar was hydrogenated at a reaction temperature of 300°C, a reaction pressure of 10 MPa, and a liquid hourly space velocity of 0.8 h -1 The reaction was carried out under conditions where the volume ratio of hydrogen to oil was 800:1.

[0111] The properties of the coal tar and the first heavy component 4 are shown in Table I-1.

[0112] The DOA from the vacuum residue was blended with the first heavy component 4 in a weight ratio of 15:10, and the properties of the blended feedstock are shown in Table I-2.

[0113] First, the mixed feedstock of DOA and the first heavy component 4 was mixed with hydrogen in a hydrocracking unit (the volume ratio of the hydrogen supply to the mixed material formed by the deoiled asphalt and the first heavy component 4 was 100, the operating temperature of the hydrocracking unit was 320°C, and the pressure was 12 MPa), and the resulting mixed material was supplied to the first reaction unit. The first reaction unit had a reaction temperature of 350°C, a reaction pressure of 12 MPa, and a liquid hourly space velocity of 0.6 h . -1 The volume ratio of recycled oil to crude oil at the inlet of the first reaction unit was 2:1. After hydrogenation, the properties of the mixed feedstock were shown in Table I-3.

[0114] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0115] The second light component at a temperature below 350°C was tested in the second reaction unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0116] [Example I-5] The same method as in Example I-3 was used, except for the following conditions: In this example, the hydrotreating first reaction unit was operated at 395°C.

[0117] Other conditions were the same as in Example I-3.

[0118] After hydrogenation, the properties of the mixed feedstock are shown in Table I-3.

[0119] The main physicochemical properties of the obtained second heavy fraction at temperatures above 350°C are shown in Table I-3.

[0120] [Example I-6] The feedstock, catalyst loading and operating conditions of the heavy oil liquid phase hydrotreater for Example I-6 were the same as those of Example I-1, except for the following: The same mixed feedstock as in Example I-1 was subjected to liquid phase heavy oil hydrogenation, and the reaction temperature was increased by 3°C every 30 days. The operation was stopped after 360 days of the hydrogenation test run.

[0121] The mineral-rich precursor material 1 and mineral-rich precursor material 2 initially charged into the reactor became V-rich material 1 and V-rich material 2, which had V contents of 76 wt% and 71 wt%, respectively, after roasting analysis, representing V contents more than 10 times higher than that of natural ore and therefore were high-quality materials for preparing high-value V2O5.

[0122] [Example I-7] In a small-scale catalytic cracking fixed fluidized bed test apparatus, a catalytic cracking test was carried out on the second light component from Example I-3 at temperatures below 350°C. The catalyst used was catalytic cracking catalyst MLC-500 manufactured by Sinopec Catalyst Co., Ltd., Changing Division. The fluidized catalyst unit was operated under the following conditions: reaction temperature: 540°C, catalyst to oil ratio: 6, and residence time: 2 seconds.

[0123] As a result, a product gasoline was obtained in a yield of 42% by weight, and its RON octane number was 92.

[0124] [Example I-8] In this example, the procedure was the same as that in Example I-1, except that the obtained second heavy component was fed to a delayed coking unit for reaction to obtain coker gasoline, coker diesel and coker wax oil.

[0125] The delayed coking unit was operated under conditions of a reaction temperature of 510° C. and a residence time of 0.6 hours.

[0126] The sulfur content of the coker diesel was 0.26 wt%, the condensation point was -11°C, and the cetane number was 48.

[0127] The sulfur content of the coker wax oil was 1.12 wt % and the condensation point was 32°C.

[0128] Coker gasoline was obtained with a yield of 14.7%, a sulfur content of 0.10 wt.%, and an MON of 61.8.

[0129] The coker diesel and coker wax oil were recycled to the third reaction unit and mixed with LCO 1 for hydrotreating, and the reaction conditions were the same as in Example I-1.

[0130] The properties of the blend of coker diesel, coker wax oil and LCO and the properties of the first heavy component 8 are shown in Table I-1.

[0131] The DOA from the vacuum residue was blended with the first heavy component 8 in a weight ratio of 1:10, and the properties of the blended feedstock are shown in Table I-2.

[0132] First, the mixture of DOA and the first heavy component 8 was mixed with hydrogen in a hydrocracking unit (the volume ratio of the hydrogen supply to the mixed feedstock formed by the deoiled asphalt and the first heavy component 8 was 100, the operating temperature of the hydrocracking unit was 320°C, and the pressure was 8 MPa), and the resulting mixture was supplied to the first reaction unit. The first reaction unit had a reaction temperature of 360°C, a reaction pressure of 8 MPa, and a liquid hourly space velocity of 0.3 h .-1 The volume ratio of recycled oil to crude oil at the inlet of the first reaction unit was 0.5:1. After hydrogenation, the properties of the mixed feedstock were shown in Table I-3.

[0133] The liquid phase product obtained from the first reaction unit was fractionated, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0134] The second light component at a temperature below 350°C was tested in the second reaction unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0135] [Example I-9] The second light component below 350°C obtained in Example I-1 was tested using a diesel hydrocracker to obtain a diesel component.

[0136] The operating conditions were a reaction temperature of 360°C, a reaction pressure of 10 MPa, a hydrogen to oil volume ratio of 1000, and a liquid volumetric space velocity of 1.0 h -1 This included being

[0137] As a result, the sulfur content of the diesel component was 5 ppm, the condensation point was -32°C, and the cetane number was 53.

[0138] [Example I-10] In this example, the procedure was similar to that of Example I-1, except that the catalyst loading in the first reaction unit was as follows: The hydrogenation catalyst, mineral-rich precursor material 1, hydrodemetallization and desulfurization catalyst, and hydrodesulfurization catalyst were sequentially loaded according to the flow direction of the reactants. In the first reaction unit, the loading ratio of the catalysts was as follows: RG-30B:mineral-rich precursor material 1:RDM-33B:RCS-31=6:60:14:20 (V / V).

[0139] After hydrogenation, the properties of the mixed feedstock are shown in Table I-3.

[0140] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0141] The second light component at a temperature below 350°C was tested in the second reaction unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0142] [Example I-11] In this example, the procedure was similar to that of Example I-1, except that the catalyst loading in the first reaction unit was as follows: According to the flow direction of the reactants, the hydrogenation catalyst, mineral-rich precursor material 2, mineral-rich precursor material 1, hydrodemetallization and desulfurization catalyst, and hydrodesulfurization catalyst were sequentially loaded. In the first reaction unit, the catalyst loading ratio was as follows: RG-30B: mineral-rich precursor material 2: mineral-rich precursor material 1: RDM-33B: RCS-31 = 6:30:30:14:20 (V / V).

[0143] After hydrogenation, the properties of the mixed feedstock are shown in Table I-3.

[0144] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0145] The second light component at a temperature below 350°C was tested in the second reaction unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0146] [Example I-12] In this example, the procedure was similar to that of Example I-1, except that the catalyst loading in the first reaction unit was as follows: According to the flow direction of the reactants, the hydroprotection catalyst, the hydrodemetallization and desulfurization catalyst, and the hydrodesulfurization catalyst were sequentially loaded. In the first reaction unit, the loading ratio between the catalysts was as follows: RG-30B:RDM-33B:RCS-31=15:35:50 (V / V).

[0147] After hydrogenation, the properties of the mixed feedstock are shown in Table I-3.

[0148] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0149] The second light component at a temperature below 350°C was tested in the second reaction unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0150] [Example I-13] In this example, the procedure was similar to that of Example I-1, except that the catalyst loading in the first reaction unit was as follows: According to the flow direction of the reactants, the hydrogenation catalyst, the mineral-rich precursor material 3, the hydrodemetallization and desulfurization catalyst, and the hydrodesulfurization catalyst were sequentially loaded. In the first reaction unit, the loading ratio between the catalysts was as follows: RG-30B:mineral-rich precursor material 3:RDM-33B:RCS-31=10:40:20:30 (V / V).

[0151] After hydrogenation, the properties of the mixed feedstock are shown in Table I-3.

[0152] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0153] The second light component at a temperature below 350°C was tested in the second reaction unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0154] [Comparative example I-1] The catalyst and equipment were the same as in Example I-1, except for the following: In the comparative example, aromatic-rich distillate oil QY (aromatic content 20 wt%) was directly mixed with DOA without passing through the partial hydrogenation unit. DOA and QY were mixed in a weight ratio of 1:10, and the mixed feedstock properties are shown in Table I-2.

[0155] Similar to Example I-1, the mixed feedstock in this comparative example was first mixed with hydrogen in a hydrocracking unit, and the resulting mixture was fed to a first reaction unit where it was hydrotreated. Product properties are shown in Table I-3.

[0156] The liquid phase product obtained from the hydrotreatment in the first reaction unit was fractionated, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0157] The second light component at a temperature below 350°C was tested in a fixed-bed hydrocracking unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0158] [Comparative example I-2] The catalyst and equipment were the same as in Example I-1, except for the following: In the comparative example, the aromatic-rich distillate oil QY was directly mixed with DOA without passing through the partial hydrosaturation unit. DOA and QY were mixed in a weight ratio of 2:10, and the mixed feedstock properties are shown in Table I-2.

[0159] Similar to Example I-1, the mixed feedstock in this comparative example was first mixed with hydrogen in a hydrocracking unit, and the resulting mixture was fed to a first reaction unit where it was hydrotreated. The properties of the water product are shown in Table I-3.

[0160] The liquid phase product obtained from the hydrotreatment in the first reaction unit was fractionated, and the properties of the second heavy component at temperatures above 350°C are shown in Table I-4.

[0161] The second light component at a temperature below 350°C was tested in a fixed-bed hydrocracking unit to obtain hydrocracking products, the properties of which are shown in Table I-5.

[0162] [Comparative example I-3] The catalyst and equipment were the same as in Example I-1, except for the following: In a comparative example, aromatic-rich distillate oil QY was directly mixed with DOA without passing through a partial hydrosaturation unit. DOA and QY were mixed in a weight ratio of 3:10. The mixed feedstock contained a large amount of solids (100°C), so the following experiment could not be carried out.

[0163] [Table 1]

[0164] [Table 2]

[0165] [Table 3]

[0166] [Table 4]

[0167] [Table 5]

[0168] [Table 6]

[0169] [Table 7]

[0170] [Example B] Solvent deasphalting was carried out using vacuum residue as the feedstock, and the solvent was a hydrocarbon mixture containing butane (butane content 70 wt%). The solvent deasphalting was carried out at 120°C with a solvent:vacuum residue ratio of 3:1 (by weight), resulting in a 70 wt% yield of DAO and a 30 wt% yield of DOA.

[0171] The properties of the obtained DAO and DOA are shown in Table II-1.

[0172] The properties of the obtained DAO and DOA are shown in Table II-1.

[0173] [Example II-1] The DAO and DOA used in this example were from Example II-B.

[0174] The properties of the liquid phase product obtained from the DAO subjected to the hydrogenation reaction in the fourth reaction unit are shown in Table II-1. The liquid product was fed to the DCC unit for reaction to obtain LCO1 and HCO1.

[0175] LCO1 was subjected to hydrosaturation in the third reaction unit and then fractionated to obtain a first light component 1 and a first heavy component 1.

[0176] The hydrogenation in the third reaction unit was carried out at a reaction temperature of 290°C, a reaction pressure of 4 MPa, and a liquid hourly space velocity of 1 h -1 The operation was carried out under conditions of 100:1 and a volume ratio of hydrogen to oil of 800:1. The properties of LCO1 and first heavy component 1 are shown in Table II-2.

[0177] DOA and First Heavy Component 1 were mixed in a weight ratio of 1:10, and the properties of the mixed feedstock are shown in Table II-3.

[0178] DOA and the first heavy component 1 were mixed with hydrogen in a hydrocracking unit to obtain a mixed material (the hydrogen content of which is shown in Table II-3). The mixed material was subjected to a reaction temperature of 360°C, a reaction pressure of 10 MPa, and a liquid hourly space velocity of 0.3 h -1The first reaction unit was operated under the conditions that the volume ratio of recycled oil to crude oil at the inlet of the first reaction unit was 0.5:1. After hydrogenation, the properties of the mixed feedstock were shown in Table II-4.

[0179] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0180] The second light fraction was tested in the second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0181] [Example II-2] The DAO and DOA used in this example were from Example II-B.

[0182] The properties of the liquid phase product obtained from the DAO subjected to the hydrogenation reaction in the fourth reaction unit are shown in Table II-1. The liquid product was fed to the DCC unit for reaction to obtain LCO2 and HCO2.

[0183] The HCO2 was subjected to hydrogenation saturation in the third reaction unit and then fractionated to provide a first light component 2 and a first heavy component 2. The hydrogenation in the third reaction unit was carried out at a reaction temperature of 330°C, a reaction pressure of 6 MPa, and a liquid hourly space velocity of 1 h -1 The operation was carried out under the conditions of 100:1 and a volume ratio of hydrogen to oil of 800:1. The properties of HCO2 and the first heavy component 2 are shown in Table II-2.

[0184] DOA and first heavy component 2 were mixed in a weight ratio of 5:10, and the properties of the mixed feedstock are shown in Table II-3.

[0185] The DOA and the first heavy component 2 were mixed with hydrogen in a hydrocracking unit to obtain a mixed material (the hydrogen content of which is shown in Table II-3). The mixed material was subjected to a reaction temperature of 380°C, a reaction pressure of 8 MPa, and a liquid hourly space velocity of 0.3 h -1The first reaction unit was operated under the conditions that the volume ratio of recycled oil to crude oil at the inlet of the first reaction unit was 0.5:1. After hydrogenation, the properties of the mixed feedstock were shown in Table II-4.

[0186] The liquid phase product obtained by the treatment in the first reaction unit was fractionated, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0187] The second light fraction was tested in the second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0188] [Example II-3] The DAO and DOA used in the examples were from Example II-B.

[0189] The properties of the liquid phase product obtained from the DAO hydrogenated in the fourth reaction unit are shown in Table II-1. The liquid phase product was fed to a DCC unit (under the same operating conditions as in Example II-1) for reaction to obtain LCO1 and HCO1.

[0190] LCO1 was hydrogenated and saturated in the third reaction unit, and then fractionated to obtain a first light component 3 and a first heavy component 3. The hydrogenation in the third reaction unit was carried out at a reaction temperature of 320°C, a reaction pressure of 6 MPa, and a liquid hourly space velocity of 1 h -1 The operation was carried out under conditions of 100:1 and a volume ratio of hydrogen to oil of 800:1. The properties of LCO1 and first heavy component 3 are shown in Table II-2.

[0191] DOA and first heavy component 3 were blended in a weight ratio of 10:10, and the properties of the blended feedstock are shown in Table II-3.

[0192] The DOA and the first heavy component 3 were mixed with hydrogen in a hydrocracking unit to obtain a mixed material (the hydrogen content of which is shown in Table II-3). For the mixed material, the reaction temperature was 370°C, the reaction pressure was 8 MPa, and the liquid hourly space velocity was 0.3 h -1The first reaction unit was operated under the conditions that the volume ratio of recycled oil to crude oil at the inlet of the first reaction unit was 0.5:1. After hydrogenation, the properties of the mixed feedstock were shown in Table II-4.

[0193] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0194] The second heavy component was subjected to a coking reaction at a reaction temperature of 500°C for 0.5 hours to obtain a petroleum coke (yield 31% by weight) having a sulfur content of 2.7% by weight.

[0195] The second light fraction was tested in the second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0196] [Example II-4] The DAO and DOA used in this example were from Example II-B.

[0197] The properties of the liquid-phase product obtained from the DAO subjected to the hydrogenation reaction in the fourth reaction unit are shown in Table II-1. The liquid-phase product was fed to a DCC unit (under the same operating conditions as in Example II-1) for reaction to obtain LCO1 and HCO1.

[0198] The aromatic-rich distillate oil used in this example was coal tar (properties shown in Table II-1) derived from a Chinese coal tar unit and LCO1. LCO1 and coal tar were used in a weight ratio of 1:1. The aromatic-rich distillate oil was subjected to hydrogen saturation in the third reaction unit and then fractionated to obtain a first light component 4 and a first heavy component 4. The hydrogenation in the third reaction unit was carried out at a reaction temperature of 300°C, a reaction pressure of 10 MPa, and a liquid hourly space velocity of 0.8 h . -1 The aromatics-rich distillate oil and the first heavy component 4 were operated under the conditions of 100:1 and 800:1 volume ratio of hydrogen to oil. The properties of the aromatics-rich distillate oil and the first heavy component 4 are shown in Table II-2.

[0199] DOA and first heavy component 4 were blended in a weight ratio of 15:10, and the properties of the blended feedstock are shown in Table II3.

[0200] The DOA and the first heavy component 4 were mixed with hydrogen in a hydrocracking unit to obtain a mixed material (the hydrogen content of which is shown in Table II-3). For the mixed material, the reaction temperature was 350°C, the reaction pressure was 12 MPa, and the liquid hourly space velocity was 0.3 h -1 The first reaction unit was operated under the conditions that the volume ratio of recycled oil to crude oil at the inlet of the first reaction unit was 0.5:1. After hydrogenation, the properties of the mixed feed were shown in Table II-4.

[0201] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0202] The second light fraction was tested in the second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0203] [Example II-5] The same method as in Example II-3 was used, except for the following: In this example, the hydrotreating first reaction unit was operated at 395°C.

[0204] Other conditions were the same as in Example II-3.

[0205] After hydrogenation, the properties of the mixed feedstock are shown in Table II-4.

[0206] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0207] [Example II-6] The catalyst loading and hydrotreating conditions were the same as in Example II-4.

[0208] The same mixed feedstock as in Example II-4 was subjected to hydrotreatment in the first reaction unit, and then the reaction temperature was increased by 3°C every 30 days. The operation was stopped after 360 days of hydrogenation test operation.

[0209] The mineral-rich precursor material 1 and mineral-rich precursor material 2 initially charged into the reactor became V-rich material 1 and V-rich material 2 after the reaction, which had V contents of 69 wt % and 60 wt %, respectively, after roasting analysis, and therefore were high-quality materials for preparing valuable V2O5.

[0210] [Example II-7] A catalytic cracking test was carried out on the second light component from Example II-3 at temperatures below 350°C in a small-scale fixed fluidized bed catalytic cracking test apparatus. The catalyst used was catalytic cracking catalyst MLC-500 manufactured by SINOPEC CATALYST CO., LTD. CHANGLING DIVISION. The fluidized catalyst unit was operated under the following conditions: reaction temperature: 540°C, catalyst to oil ratio: 6, and residence time: 3 seconds.

[0211] As a result, a product gasoline was obtained in a yield of 40% by weight, and its RON octane number was 93.

[0212] [Example II-8] In this example, the procedure was the same as that of Example II-1, except that the obtained second heavy component was fed to a delayed coking unit for reaction to provide coker gasoline, coker diesel and coker waxy oil.

[0213] The delayed coking unit was operated under conditions of a reaction temperature of 510°C and a residence time of 0.6 hours.

[0214] The sulfur content of the coker diesel was 0.26 wt%, the condensation point was -11°C, and the cetane number was 48.

[0215] The sulfur content of the coker wax oil was 1.12 wt % and the condensation point was 32°C.

[0216] Coker gasoline was obtained with a yield of 14.7%, a sulfur content of 0.10 wt.%, and an MON of 61.8.

[0217] The coker diesel and coker wax oil were recycled to the third reaction unit and mixed with LCO1 for hydrogen saturation, and then fractionated to obtain a first light component 8 and a first heavy component 8 with a cut point of 180°C, under the same reaction conditions as in Example II-1. The properties of the mixed oil of coker diesel, coker wax oil, LCO1 and the properties of the first heavy component 8 are shown in Table II-2.

[0218] DOA from Example II-B and first heavy component 8 were mixed in a mass ratio of 1:10, and the properties of the mixed feedstock were shown in Table II-3.

[0219] DOA and the first heavy component 8 were mixed with hydrogen in a hydrocracking unit to obtain a mixed material (the hydrogen content of which is shown in Table II-3). The first reaction unit was operated at a reaction temperature of 360°C, a reaction pressure of 8 MPa, and a liquid hourly space velocity of 0.3 h -1 and a volume ratio of hydrogen to oil of 800: 1. After hydrogenation, the properties of the mixed feedstock are shown in Table II-4.

[0220] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0221] The second light fraction was tested in the second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0222] [Example II-9] The second light component obtained from Example II-1 was tested in a diesel hydro-upgrader at a temperature below 350°C to obtain a diesel component.

[0223] The diesel hydrogenation quality improvement equipment was designed with the following conditions: reaction temperature 360°C, reaction pressure 12 MPa, volume ratio of hydrogen to oil 1000, and liquid volumetric space velocity 1.0 h -1 The engine was operated under the following conditions.

[0224] As a result, the sulfur content of the diesel component was 5 ppm, the condensation point was -33°C, and the cetane number was 53.

[0225] [Example II-10] In this example, the procedure was similar to that of Example II-1, except that the catalyst loading in the first reaction unit was as follows: The hydrogenation catalyst, mineral-rich precursor material 1, hydrodemetallization and desulfurization catalyst, and hydrodesulfurization catalyst were sequentially loaded according to the flow direction of the reactants. In the first reaction unit, the loading ratio of the catalysts was as follows: RG-30B:mineral-rich precursor material 1:RDM-33B:RCS-31=6:60:14:20 (V / V).

[0226] After hydrogenation, the properties of the mixed feedstock are shown in Table II-4.

[0227] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0228] The second light fraction was tested in the second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0229] [Example II-11] In this example, the procedure was similar to that of Example II-1, except that the catalyst loading in the first reaction unit was as follows: According to the flow direction of the reactants, the hydrogenation catalyst, mineral-rich precursor material 2, mineral-rich precursor material 1, hydrodemetallization and desulfurization catalyst, and hydrodesulfurization catalyst were sequentially loaded. In the first reaction unit, the catalyst loading ratio was as follows: RG-30B: mineral-rich precursor material 2: mineral-rich precursor material 1: RDM-33B: RCS-31 = 6:30:30:14:20 (V / V).

[0230] After hydrogenation, the properties of the mixed feedstock are shown in Table II-4.

[0231] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0232] The second light fraction was tested in the second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0233] [Example II-12] In this example, the procedure was similar to that of Example II-1, except that the catalyst loading in the first reaction unit was as follows: The catalyst loading ratios were RG-30B:RDM-33B:RCS-31 = 15:40:45 (V / V).

[0234] After hydrogenation, the properties of the mixed feedstock are shown in Table II-4.

[0235] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0236] The second light fraction was tested in the second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0237] [Example II-13] In this example, the procedure was similar to that of Example II-1, except that the catalyst loading in the first reaction unit was as follows: The hydrogenation catalyst, mineral-rich precursor material 3, hydrodemetallization and desulfurization catalyst, and hydrodesulfurization catalyst were sequentially loaded according to the flow direction of the reactants. The loading ratio between the catalysts was RG-30B:mineral-rich precursor material 3:RDM-33B:RCS-31 = 10:40:25:35 (V / V).

[0238] After hydrogenation, the properties of the mixed feedstock are shown in Table II-4.

[0239] The liquid phase product obtained by the treatment in the first reaction unit was fractionally distilled, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0240] The second light fraction was tested in a second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0241] [Comparative Example II-1] The catalyst and equipment were the same as in Example II-1, except for the following: In the comparative example, aromatic-rich distillate oil QY (aromatic content 20 wt%) was directly mixed with DOA without passing through the partial hydrogenation unit. DOA and QY were mixed in a weight ratio of 1:10, and the properties of the mixed feedstock are shown in Table II-3.

[0242] This mixed material was mixed with hydrogen in a hydrocracking unit, and the resulting mixed material (the hydrogen content of which is shown in Table II-3) was hydrotreated in the first reaction unit, and the product properties were shown in Table II-4.

[0243] The liquid phase product obtained by hydrotreating in the first reaction unit was fractionated, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0244] The second light fraction was tested in the second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0245] [Comparative Example II-2] The catalyst and equipment were the same as in Example II-1, except for the following: In the comparative example, the aromatic-rich distillate oil QY was directly mixed with DOA without passing through the partial hydrogenation unit. DOA and QY were mixed in a weight ratio of 2:10, and the properties of the mixed feedstock are shown in Table II-3.

[0246] This mixed material was mixed with hydrogen in a hydrocracking unit, and the resulting mixed material (the hydrogen content of which is shown in Table II-3) was hydrotreated in the first reaction unit, and the product properties were shown in Table II-4.

[0247] The liquid phase product obtained by hydrotreating in the first reaction unit was fractionated, and the properties of the second heavy component at temperatures above 350°C are shown in Table II-5.

[0248] The second light fraction was tested in the second reaction unit at a temperature below 350°C to obtain hydrocracking products, the properties of which are shown in Table II-6.

[0249] [Comparative Example II-3] The catalyst and equipment were the same as in Example II-1, except for the following: In Comparative Example II-3, aromatic-rich distillate oil QY was directly mixed with DOA without passing through a partial hydrogenation unit. DOA and QY were mixed in a weight ratio of 3:10. The mixed feedstock contained a large amount of solids (100°C), so the following experiment could not be carried out.

[0250] [Table 8]

[0251] [Table 9]

[0252] [Table 10]

[0253] [Table 11]

[0254] [Table 12]

[0255] [Table 13]

[0256] [Table 14]

[0257] [Table 15]

[0258] From these results, it can be seen that the technology of the present invention enables a high quality raw material for producing low sulfur marine fuel or low sulfur coke products from DOA.

[0259] Furthermore, the technology of the present invention can provide high quality gasoline products that meet the National V standard.

[0260] Although the preferred embodiments of the present invention have been described in detail above, the present invention is not limited thereto. Within the technical spirit of the present invention, many simple modifications can be made to the technical solutions of the present invention, including combining various technical features in any other suitable manner, and these simple modifications and combinations should also be considered as the disclosure of the present invention, and all fall within the scope of the present invention. [Brief explanation of the drawings]

[0261] [Figure 1]FIG. 1 is a flow chart for processing aromatic-rich distillate oils according to a preferred embodiment of the present invention. [Figure 2] FIG. 2 is a flow chart for processing an aromatic-rich distillate oil according to a first embodiment of the present invention.

Claims

1. (1) introducing the aromatics-rich distillate oil into a third reaction unit for hydrogen saturation, and then fractionating to provide a first light component and a first heavy component, wherein the first light component and the first heavy component have a cut point of 100 to 250°C, and the aromatics content in the first heavy component is 20 wt% or more; (2) introducing the deoiled asphalt and the aromatics-containing stream containing the first heavy component into a hydrogen dissolution unit and mixing with hydrogen to form a mixed material, and introducing the mixed material into a first reaction unit for hydrogenation reaction; wherein the deoiled asphalt and the aromatics-containing stream are used in an amount ratio such that the mixed feedstock formed by the deoiled asphalt and the aromatics-containing stream is in a liquid state at a temperature of 400°C or less; (3) fractionating the liquid phase product from the first reaction unit to provide a second light component and a second heavy component; wherein the cut point for the second light component and the second heavy component is 240-450°C; (41) introducing the second light component into a second reaction unit for reaction to provide at least one product selected from a gasoline component, a diesel component, and a BTX feedstock component; wherein the second reaction unit is at least one selected from the group consisting of a hydrocracking unit, a catalytic cracking unit, and a diesel hydroupgrading unit; and (42) introducing the second heavy component into a delayed coking unit for reaction to provide at least one product selected from the group consisting of coker gasoline, coker diesel, coker waxy oil, and low sulfur petroleum coke; or using the second heavy component as a component of low sulfur marine fuel oil; A method for processing an aromatics-rich distillate oil, characterized in that the first heavy component in step (1) is used as the first heavy component contained in the aromatics-containing stream in step (2).

2. 2. The method of claim 1, wherein in step (2), the first reaction unit comprises a mineral-rich precursor material and / or a hydrogenation catalyst, the first reaction unit is a liquid-phase hydrogenation reaction unit, and the mineral-rich precursor material is a material capable of adsorbing at least one metal selected from V, Ni, Fe, Ca, and Mg.

3. In the step (2), the deoiled asphalt and the aromatics-containing stream are mixed to form a mixed feedstock having a viscosity of 400 mm at 100°C. ２ 10. The method of claim 1, wherein the hydroxyl group is used in a ratio of 0.1 to 0.

5.

4. In the step (2), the aromatic-containing stream further comprises aromatic hydrocarbons and / or aromatic oils, wherein the aromatic oils are at least one selected from the group consisting of LCO, HCO, FGO, ethylene tar, coal tar, coker diesel, and coker wax oil; the aromatic hydrocarbons are benzene, toluene, xylene, naphthalene, at least one C １－６ 2. The method of claim 1, wherein the alkyl group is at least one selected from the group consisting of naphthalene substituted with an alkyl group, and tricyclic or higher aromatic hydrocarbons.

5. 2. The method of claim 1, wherein the aromatics content in the aromatic-rich distillate oil is 20 wt.% or more.

6. 10. The method of claim 1, wherein in step (2), the deoiled asphalt is obtained by subjecting a heavy oil feedstock to a solvent deasphalting process in a solvent deasphalting unit.

7. 2. The method of claim 1, wherein in step (2), the aromatics-containing stream is an aromatics-rich distillate oil, and the weight ratio of the amount of the deoiled asphalt to the amount of the aromatics-containing stream is 1:10 to 50:

10.

8. 10. The method of claim 1, further comprising recycling the coker diesel and / or coker waxy oil obtained in step (42) to the third reaction unit of step (1) for hydrogen saturation.

9. In step (1), the third reaction unit is at least one of a fixed bed reactor, a moving bed reactor, and an ebullated bed reactor; and / or The third reaction unit has a reaction temperature of 200 to 420°C, a reaction pressure of 2 to 18 MPa, and a liquid hourly space velocity of 0.3 to 10 h －１ and the volume ratio of hydrogen to oil is 50 to 5000.

10. In step (2), the first reaction unit has a reaction temperature of 260 to 500°C, a reaction pressure of 2.0 to 20.0 MPa, a volume ratio of recycled oil to feedstock crude oil at the inlet of the first reaction unit of 0.1:1 to 15:1, and a liquid hourly space velocity of 0.1 to 1.5 h －１ 2. The process of claim 1, wherein the process is operated under conditions where:

11. In the step (2), the mineral-rich precursor material has an ignition loss of 3% by weight or more and a specific surface area of ​​80 m ２ 3. The method according to claim 2, wherein the water absorption is 0.9 g / g or more and the water absorption is 0.9 g / g or more.

12. 12. The method of claim 11, wherein in step (2), the mineral-rich precursor material comprises a support and an active component element supported on the support, wherein the support is at least one selected from the group consisting of aluminum hydroxide, alumina, and silica, and the active component element is at least one metal element selected from the group consisting of Group VIB and Group VIII.

13. 13. The method of claim 12, wherein in step (2), a first mineral-rich precursor material and a second mineral-rich precursor material are successively charged into the first reaction unit according to a flow direction of reactants, and the second mineral-rich precursor material has an ignition loss equal to or greater than that of the first mineral-rich precursor material.

14. In step (41), the second reaction unit is a hydrocracking unit, the reaction temperature is 360 to 420°C, the reaction pressure is 10.0 to 18.0 MPa, the volume ratio of hydrogen to oil is 600 to 2000, and the liquid hourly space velocity is 1.0 to 3.0 h －１ and / or the hydrocracking unit is loaded with at least one hydrotreating catalyst and at least one hydrocracking catalyst.

15. In step (41), the second reaction unit is a catalytic cracking unit, and the catalytic cracking unit is a fluid catalytic cracking unit; 10. The method of claim 1, wherein the fluid catalytic cracking unit is operated under conditions of reaction temperature of 500-600°C, catalyst to oil ratio of 3-12, and residence time of 0.6-6 seconds.

16. In step (41), the second reaction unit is a diesel hydrogenation upgrading unit, the reaction temperature is 330-420°C, 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 and / or the diesel hydroupgrading unit is loaded with at least one diesel hydroupgrading catalyst.

17. In step (42), the second heavy component is introduced into the delayed coking unit for reaction to provide at least one product selected from coker gasoline, coker diesel, coker wax oil, and low-sulfur petroleum coke, the delayed coking unit being operated under conditions of a reaction temperature of 440 to 520°C and a residence time of 0.1 to 4 hours; 10. The method of claim 1, wherein in step (42), the second heavy component has a sulfur content of 1.8 wt.% or less, and the second heavy component is introduced into the delayed coking unit for reaction to provide the low sulfur petroleum coke.

18. 2. The method of claim 1, wherein in step (42) the second heavy component is used as a low sulfur marine fuel oil component and conditions are controlled such that the low sulfur marine fuel oil component has a sulfur content of 0.5 wt.% or less.

19. (11) introducing the heavy oil feedstock into a solvent deasphalting unit for solvent deasphalting to provide deoiled asphalt and deasphalted oil; (12) introducing the deasphalted oil into a fourth reaction unit for hydrogenation reaction, and introducing the liquid phase effluent obtained in the fourth reaction unit into a DCC unit for reaction to provide propylene, LCO, HCO and slurry oil, wherein the fourth reaction unit is a fixed-bed reaction unit; and 2. The method of claim 1, further comprising using the aromatic-rich distillate containing LCO and / or HCO from the DCC unit as the aromatic-rich distillate in step (1).

20. 20. The method of claim 19, further comprising recycling the coker diesel and / or coker waxy oil obtained in step (42) to the third reaction unit for hydrogen saturation.

21. In step (12), the fourth reaction unit has a reaction temperature of 280 to 400°C, a reaction pressure of 6.0 to 14.0 MPa, a volume ratio of hydrogen to oil of 600 to 1200, and a liquid hourly space velocity of 0.3 to 2.0 h －１ and / or operated under conditions in which 20. The method of claim 19, wherein in step (12), the fourth reaction unit is charged with at least two hydrogenation catalysts; at least one of the hydrogenation catalysts is a catalyst capable of catalyzing at least one reaction selected from the group consisting of hydrodemetallization, hydrodesulfurization, and hydrodecarburization; and at least one of the hydrogenation catalysts comprises alumina as a support, a metal element from Group VIB and / or Group VIII as an active component element, and optionally at least one auxiliary element selected from P, Si, F, and B.

22. a third reaction unit for hydrogen saturation and fractionation on the aromatic-rich distillate oil, providing a first light component and a first heavy component; a hydrogen dissolution unit in fluid communication with the third reaction unit for mixing therein the deoiled asphalt and the aromatics-containing stream comprising the first heavy component from the third reaction unit with hydrogen; a first reaction unit in fluid communication with the hydrogen dissolving unit, the first reaction unit being a liquid-phase hydrogenation reaction unit used for carrying out a hydrogenation reaction of the mixed material from the hydrogen dissolving unit therein; a separation unit in fluid communication with the first reaction unit for fractionating the liquid-phase product from the first reaction unit therein to provide a second light component and a second heavy component; a second reaction unit in fluid communication with the separation unit for reacting therein the second light component obtained in the separation unit, the second reaction unit being at least one selected from the group consisting of a hydrocracking unit, a catalytic cracking unit, and a diesel hydroupgrading unit; a delayed coking unit in fluid communication with the separation unit for reacting therein the second heavy component obtained from the separation unit to provide at least one product selected from the group consisting of coker gasoline, coker diesel, coker waxy oil, and low sulfur petroleum coke; and an outlet in fluid communication with the separation unit for discharging the second heavy component obtained from the separation unit from the system as a low sulfur marine fuel oil fraction.

23. 23. The system of claim 22, wherein the delayed coking unit is in fluid communication with the hydrogen dissolution unit, and coker diesel and / or coker waxy oil obtained in the delayed coking unit is recycled to the first reaction unit.

24. 23. The system of claim 22, further comprising a solvent deasphalting unit in fluid communication with the hydrogen dissolution unit, the solvent deasphalting unit being used for solvent deasphalting a heavy oil feedstock therein to provide deoiled asphalt and deasphalted oil, and for introducing the deoiled asphalt obtained after the solvent deasphalting into the hydrogen dissolution unit.

25. a solvent deasphalting unit used to subject heavy oil feedstock to solvent deasphalting treatment therein to provide deoiled asphalt and deasphalted oil; a fourth reaction unit in fluid communication with the solvent deasphalting unit, the fourth reaction unit being a fixed-bed reaction unit for hydrogenating the deasphalted oil from the solvent deasphalting unit therein; a DCC unit in fluid communication with the fourth reaction unit for reacting therein the liquid-phase effluent obtained from the fourth reaction unit, to provide propylene, LCO, HCO, and a slurry oil; 23. The system of claim 22, wherein the DCC unit is in fluid communication with the third reaction unit for transporting the aromatic-rich distillate comprising LCO and / or HCO from the DCC unit to the third reaction unit for use as an aromatic-rich distillate.

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