A method for efficient conversion of residual oil

Abstract

This invention discloses a method for efficient conversion of residual oil, the technology including the following: (1) a heavy feedstock is hydrotreated by a fluidized bed to obtain unconverted oil; the unconverted oil is controlled to have the following properties: the mass content of n-heptane asphaltenes is less than 15%, preferably 1%~10%, the sum of the mass content of aromatics and gums is ≥70%, preferably 80%-95%, the ratio of aromatics / gum is ≥5, preferably 610, and the mass content of sulfur is less than 1.5%, preferably 0.3%~1.0%; (2) the unconverted oil in step (1) enters the first solvent extraction unit to obtain the first raffinate phase, the first solvent and The first solvent and the extract phase are separated and the first solvent is recovered to obtain a first extract phase rich in alkane components; (3) The first raffinate phase in step (2) enters the second solvent extraction unit to obtain a second raffinate phase, a second solvent and an extract phase mixture, and then the second solvent and the extract phase are separated and the second solvent is recovered to obtain a second extract phase rich in aromatic components; (4) The second extract phase in step (3) enters the coking unit for coking reaction to obtain coke for negative electrode materials and a gas-liquid mixture, which is returned to the fluidized bed fractionation unit to be mixed and fractionated with the fluidized bed products. The method can realize the production of low-quality residue oil into high-value special products.

Description

Technical Field

[0001] This invention belongs to the field of petrochemical technology, and specifically relates to a method for the efficient conversion of residual oil. It employs fluidized bed hydrogenation technology, solvent extraction separation, and thermal condensation methods to overcome the limitation of insufficient high-quality raw materials in existing technologies for producing high-end products. This enables the production of high-value specialty chemical raw materials from low-quality residual oil, significantly improving the economic efficiency of refineries. Background Technology

[0002] Currently, the refining structure is undergoing transformation and upgrading, resulting in weak demand for petroleum products, while demand for chemical raw materials is strong, especially for high-end carbon materials and specialty oils, where market demand is huge, such as negative electrode coke materials for automotive batteries, high viscosity index heavy lubricating oils, and low-sulfur marine fuel oils. At present, the production of these products generally faces problems such as scarce raw material costs and high production costs.

[0003] Studies have shown that the ideal feedstock for producing negative electrode coke materials is the tri- and tetra-cyclic aromatic hydrocarbon components in petroleum fractions, while the ideal feedstock for high viscosity index heavy lubricating oil is the long-chain isomeric alkane components. The residue oil system, however, is a complex system composed of alkanes, aromatics, gums, and asphaltenes, requiring the use of catalysts to convert non-ideal components into ideal components as much as possible.

[0004] CN108473889A provides a method for forming lubricant base oils from vacuum residue or other feedstocks at 510°C+. The feedstock can be deasphalted, followed by catalytic and / or solvent processing to form lubricant base oils, including bright oils. This catalytic processing can be equivalent to processing in at least two stages. The conversion rates achieved in each stage can be varied to produce bright oils with various properties. However, this patented method has poor adaptability to feedstocks and is difficult to directly produce bright oil products from low-quality vacuum residue.

[0005] CN 103102986 A discloses a combined process of residue oil hydrotreating and delayed coking. Residue oil, coking wax oil, and hydrogen are fed together into a hydrotreating unit. The hydrotreated residue oil is then mixed with separated vacuum gas oil and fed into a delayed coking unit to separate coking products. All coking gas oil is recycled back to the residue oil hydrotreating unit. However, while this patented method can alleviate the problem of high sulfur content in petroleum coke products, it is difficult to produce high-quality coke feedstock specifically for anode materials. Summary of the Invention

[0006] To address the shortcomings of existing technologies, this invention provides a method for the efficient conversion of residual oil. Employing fluidized bed hydrogenation technology, two-stage extraction separation, and thermal condensation, this method converts as many non-ideal components as possible, such as asphaltenes, in low-quality residual oil while retaining polycyclic aromatic hydrocarbons (PAHs). Through two-stage step-by-step extraction separation technology, high-quality feedstocks rich in alkane components and aromatic hydrocarbon-rich feedstocks for producing coke specifically for anode materials are obtained. A coking process is then used to convert the aromatic hydrocarbon-rich components into coke specifically for anode materials, thus realizing the production of high-value specialty chemical feedstocks from low-quality residual oil.

[0007] This invention provides a method for efficient conversion of residual oil, the technology including the following: (1) Unconverted oil is obtained by hydrotreating the residue oil feedstock in a fluidized bed; the unconverted oil is controlled to have the following properties: the mass content of n-heptane asphaltene is less than 15%, preferably 1%~10%, the sum of the mass content of aromatics and gums is ≥70%, preferably 80%-95%, the ratio of aromatics / gum is ≥5, preferably 6-10, and the mass content of sulfur is less than 1.5%, preferably 0.3%~1.0%; (2) The unconverted oil in step (1) enters the first solvent extraction unit to obtain a mixture of the first raffinate phase, the first solvent and the extract phase. Then the first solvent and the extract phase are separated to recover the first solvent and obtain the first extract phase rich in alkane components. (3) In step (2), the first raffinate enters the second solvent extraction unit to obtain a mixture of the second raffinate, the second solvent and the extract phase. Then the second solvent and the extract phase are separated to recover the second solvent and obtain the second extract phase rich in aromatic components. (4) In step (3), the second extraction phase enters the coking unit to carry out the coking reaction, and obtains the special coke for negative electrode material and the gas-liquid mixture. The gas-liquid mixture is returned to the fluidized bed fractionation unit to be mixed and fractionated with the fluidized bed product.

[0008] In the above-mentioned method for efficient conversion of residue oil, in step (1), the residue oil feedstock enters the first reaction zone of the fluidized bed. The gas-liquid mixture generated in the first reaction zone passes through an interstage separator to obtain a first gas phase and a first liquid phase. The first liquid phase enters the second reaction zone to react and obtain a gas-liquid mixture, which is then separated in the hot high-efficiency fraction to obtain a second gas phase and a second liquid phase. The first gas phase and the second gas phase are mixed and enter the cold high-efficiency fraction. The gas phase from the cold high-efficiency fraction enters the circulating hydrogen desulfurization tower to remove hydrogen sulfide and then obtains circulating hydrogen, which then enters the first and second reactors. The second liquid phase enters the hot low-efficiency fraction. The gas phase from the hot low-efficiency fraction and the liquid phase from the cold high-efficiency fraction are mixed and enter the cold low-efficiency fraction. The liquid phase from the cold low-efficiency fraction and the liquid phase from the hot low-efficiency fraction are mixed and enter the fractionation system to obtain fluidized bed gas, naphtha, diesel oil, wax oil, and unconverted oil.

[0009] In the above-mentioned method for efficient conversion of residual oil, the residual oil raw material involved in step (1) can be at least one of the inferior heavy oils such as atmospheric residue, vacuum residue, and heavy oil.

[0010] In the above-mentioned method for efficient conversion of residue oil, in step (1), a fluidized bed reactor is set up in each of the two reaction zones of the fluidized bed hydrogenation unit. The fluidized bed used can be a circulating pump fluidized bed reactor or a three-phase separator fluidized bed reactor.

[0011] In the above-mentioned method for efficient conversion of residue oil, the process parameters of the first reaction zone of the fluidized bed in step (1) are as follows: reaction temperature is 370-440℃, preferably 400-430℃; reaction pressure is 15.0-20.0MPa, preferably 16.0-18.0MPa; hydrogen-to-oil volume ratio is 300-1000, preferably 400-600; and liquid hourly space velocity is 0.1-2.0h. -1 Preferably, it is 0.2 to 0.5 h. -1 .

[0012] In the above-mentioned method for efficient conversion of residue oil, the process parameters of the second reaction zone of the fluidized bed in step (1) are as follows: reaction temperature is 370-440℃, preferably 400-430℃; reaction pressure is 15.0-20.0MPa, preferably 16.0-18.0MPa; hydrogen-to-oil volume ratio is 300-1000, preferably 400-600; and liquid hourly space velocity is 0.1-2.0h. -1 Preferably, it is 0.2 to 0.5 h. -1 .

[0013] In the above-mentioned method for efficient conversion of residue oil, the catalyst used in the first reaction zone of the fluidized bed is a spherical catalyst with a diameter of 0.3–0.7 mm. The catalyst comprises a support and an active metal, wherein the active metal can be one or more of nickel, molybdenum, or tungsten; the support can be one or more of alumina, silica, or alumina-silica. Based on catalyst mass, the support accounts for 90 wt%–96 wt%, and the active metal, calculated as oxides, accounts for 3 wt%–10 wt%. The bulk density of the catalyst is 0.4–0.8 g / cm³. 3 The catalyst has a bimodal pore structure, with mesopores having a diameter greater than or equal to 3 nm and less than 100 nm accounting for 50% to 70% of the total pore volume, and macropores having a diameter of 100 to 1000 nm accounting for 30% to 50% of the total pore volume. Among them, the mesopores have a diameter of 15 to 25 nm, and the macropores have a diameter of 700 to 900 nm.

[0014] In the aforementioned method for efficient conversion of residual oil, the catalyst used in the second reaction zone of the fluidized bed is a spherical catalyst with a diameter of 0.3–0.7 mm. The catalyst comprises a support and an active metal, wherein the active metal can be one or more of nickel, molybdenum, or tungsten; the support can be one or more of alumina, silica, or alumina-silica. Based on catalyst mass, the support accounts for 80 wt%–90 wt%, and the active metal, calculated as oxides, accounts for 10 wt%–20 wt%. The bulk density of the catalyst is 0.4–0.8 g / cm³. 3 The catalyst has a bimodal pore structure, in which pores with a diameter greater than or equal to 1 nm and less than 200 nm account for 70% to 80% of the total pore volume, and pores with a diameter of 200 to 1000 nm account for 20% to 30% of the total pore volume. Among them, mesopores have a diameter of 7 to 13 nm and macropores have a diameter of 500 to 700 nm.

[0015] In the above-mentioned method for efficient conversion of residue oil, in step (1), the catalysts in the two reaction zones of the fluidized bed are added in parallel mode, wherein the addition rate of the first reaction catalyst is 0.3 to 1.5 kg / ton of raw material, preferably 0.5 to 0.9 kg / ton of raw material; and the addition rate of the second reaction catalyst is 0.3 to 1.5 kg / ton of raw material, preferably 0.3 to 0.7 kg / ton of raw material.

[0016] In the above method for efficient conversion of residue oil, in step (1), the average catalyst bed temperature in the two reaction zones of the fluidized bed is controlled at a temperature gradient of 2~6℃, and the temperature of the second reaction zone is higher than that of the second reaction zone.

[0017] In the above-mentioned method for efficient conversion of residue oil, the ratio of temperature rise (average bed temperature minus reaction zone inlet temperature) of the two reaction zones in the fluidized bed in step (1) is controlled to be ≥1.5, preferably ≥1.7.

[0018] In the above-mentioned method for efficient conversion of residue oil, step (1) of the fluidized bed unit fractionation is divided into atmospheric fractionation and vacuum fractionation, wherein the atmospheric and vacuum fractionation points are 300~400℃, preferably 330~360℃; the atmospheric fraction can obtain hydrotreated naphtha, hydrotreated diesel and atmospheric residue oil, the cut-off points of hydrotreated naphtha and hydrotreated diesel are 160~210℃, preferably 180~200℃; the atmospheric residue oil is further fractionated into vacuum to obtain hydrotreated wax oil and vacuum residue oil, the cut-off points of hydrotreated wax oil and vacuum residue oil are 450~560℃, preferably 500~540℃.

[0019] In the above-mentioned method for efficient conversion of residue oil, step (2) involves the solvent used in the solvent extraction unit being an alkane with 3 to 4 carbon atoms, wherein the mass fraction of alkane with 3 carbon atoms in the solvent is not less than 80%; the operating conditions are an extraction temperature of 40 to 100°C, preferably 60 to 90°C, an extraction pressure of 4.0 to 7.0 MPa, preferably 4.5 to 5.0 MPa, and a solvent-to-oil volume ratio of 3.0 to 7.0, preferably 4.0 to 6.0.

[0020] In the above-mentioned method for efficient conversion of residue oil, the extraction rate in the solvent extraction unit of step (2) is related to the selection of solvent. Changes in solvent composition will cause changes in the relationship between the yield and properties of the extraction phase. In this invention, the carbon residue content in the first extraction phase is used as the control condition, and the carbon residue in the extraction phase is controlled to be no higher than 2.0%. Under this premise, the yield of the extraction phase is increased as much as possible.

[0021] In the above-mentioned method for efficient conversion of residual oil, the extract phase in step (2) can be processed by the "old three-set" process of hydrorefining-isomeric pour point depressant, furfural extraction-ketone benzene dewaxing-clay refining to generate bright oil; it can also be used as a low-sulfur marine fuel blending component or a feedstock for hydrocracking and catalytic cracking.

[0022] In the above method for efficient conversion of residue oil, step (2) solvent extraction is carried out in an extraction tower, which can be a packed tower or a tray tower, preferably a packed tower.

[0023] In the above method for efficient conversion of residue oil, in step (2) solvent extraction unit, solvent is recovered and reused. Solvent recovery is carried out under supercritical conditions, and the recovery conditions are set according to the properties of the selected solvent.

[0024] In the above-mentioned method for efficient conversion of residue oil, step (3) involves the solvent used in the solvent extraction unit being an alkane with 4 to 6 carbon atoms, wherein the mass fraction of alkane with 4 carbon atoms in the solvent is not less than 50%; the operating conditions are an extraction temperature of 100 to 160°C, preferably 120 to 140°C, an extraction pressure of 3.0 to 7.0 MPa, preferably 3.5 to 4.5 MPa, and a solvent-to-oil volume ratio of 3.0 to 7.0, preferably 4.0 to 6.0.

[0025] In the above-mentioned method for efficient conversion of residue oil, the extraction rate in the solvent extraction unit of step (3) is related to the selection of solvent. Changes in solvent composition will cause changes in the relationship between the yield and properties of the extract phase. In this invention, the C7 asphaltenes content in the second extract phase is used as the control condition, and the C7 asphaltenes content in the extract phase is controlled to be no higher than 1.0%. Under this premise, the yield of the extract phase is increased as much as possible.

[0026] In the above-mentioned method for efficient conversion of residue oil, the second extract phase in step (3) can be partially returned to the fluidized bed reaction zone, preferably returned to the second fluidized bed reaction zone, with a return ratio of 0-10% of the fresh feed of the device, preferably 2-5%.

[0027] In the above-mentioned method for efficient conversion of residual oil, the second raffinate phase in step (3) can be used as a raw material for hydrogen production or asphalt production.

[0028] In the above method for efficient conversion of residue oil, step (3) solvent extraction is carried out in an extraction tower, which can be a packed tower or a tray tower, preferably a packed tower.

[0029] In the above-mentioned method for efficient conversion of residual oil, the coking unit in step (4) includes a heating furnace, a coking tower, and a fractionation tower, and at least one heating furnace, two coking towers, and one fractionation tower. The coking reaction includes two processes: continuous coking and intermittent decoking. The outlet temperature of the heating furnace is 480~520℃, preferably 490℃~510℃; the pressure of the coking tower is 0.1MPa~0.5MPa, preferably 0.15~0.25MPa; and the recycle ratio is 0~1.0, preferably 0~0.3, by mass.

[0030] In the above-mentioned method for efficient conversion of residue oil, the coking-generated oil and gas in step (4) can also enter the fluidized bed fractionation unit, eliminating the need for the coking fractionation unit. In this mode, the circulating oil can be provided by the wax oil from the fractionation system.

[0031] In the above-mentioned method for efficient conversion of residual oil, the coking-generated oil and gas in step (4) can be separated by a separate fractionation system.

[0032] Compared with existing technologies, the high-efficiency heavy oil conversion combined process of the present invention has the following advantages: 1. This invention solves the problem of how to produce high-value-added oil products and negative electrode petroleum coke materials from low-value, inferior residue oil feedstock. Based on the thermal cracking and hydrogenation reaction laws of residue oil under the fluidized bed system, and the adsorption competition reaction mechanism of different molecules in residue oil, catalysts with different metal active phases and pore structures are used in two reactors. At the same time, a catalyst activity control strategy is proposed to convert non-ideal asphaltenes and other components in inferior residue oil as much as possible, while reducing the saturation ratio of aromatic components. The large-pore, low-cracking functional catalyst unique to the fluidized bed is adopted to retain the high-quality bright oil components, long-chain isoalkanes, in the residue oil system.

[0033] 2. The dual-stage gradient extraction technology in this invention achieves the separation of aromatic-poor components, aromatic-rich components, and asphaltenes by utilizing the difference in solubility of different hydrocarbon compositions in the solvent of the fluidized bed hydrogenated heavy oil, thereby obtaining high-quality raw materials for bright oil and special coke for negative electrode materials.

[0034] 3. The aromatic components are partially returned to the second reaction zone of the fluidized bed to achieve directional desulfurization in the second reaction, further obtaining a high aromatic hydrocarbon and low impurity fraction.

[0035] 4. The present invention proposes a common fractionation system for mixing the gas-liquid mixture generated by coking with the oil generated by fluidized bed hydrogenation. This reduces equipment investment, and the properties of the product after separation of the coking gas-liquid components are similar to those of the fluidized bed product, which does not adversely affect the subsequent processing and utilization of the fluidized bed product. Attached Figure Description

[0036] Figure 1 This is a schematic diagram of a combined process for the efficient conversion of heavy oil.

[0037] Wherein, 1-residue feedstock; 2-first reaction zone of fluidized bed; 3-interstage separator; 4-first liquid phase; 5-first gas phase; 6-second reaction zone of fluidized bed; 7-hot high-efficiency fraction; 8-second liquid phase; 9-second gas phase; 10-hot low-efficiency fraction; 11-hot low-efficiency fraction liquid phase; 12-hot low-efficiency fraction gas phase; 13-cold high-efficiency fraction; 14-cold high-efficiency fraction liquid phase; 15-cold high-efficiency fraction gas phase; 16-cold low-efficiency fraction; 17-cold low-efficiency fraction liquid phase; 18-cold low-efficiency fraction gas phase; 19-fractionation system; 20-fluidized bed gas; 21-naphtha; 22-diesel; 23-wax oil; 24-Unconverted oil; 25-First extraction tower; 26-First raffinate phase; 27-Mixed feed of first extract phase and solvent; 28-Second extraction tower; 29-Second raffinate phase; 30-Mixed feed of second extract phase and solvent; 31-Second solvent recovery tower; 32-Second solvent; 33-Second extract phase; 34-Coking tower; 35-Petroleum coke; 36-Coking gas-liquid mixture; 37-First extraction solvent recovery tower; 38-First extract phase; 39-First extraction solvent; 40-Circulating hydrogen desulfurization tower; 41-Circulating hydrogen; 42-New hydrogen. Detailed Implementation

[0038] The technical features of the present invention will be further described in the following steps with reference to the accompanying drawings and embodiments, but these embodiments are not intended to limit the present invention.

[0039] Unless otherwise specified, all percentages in the methods of this invention refer to mass percentages.

[0040] In this invention, embodiments 1, 2, and 3 adopt the following... Figure 1 The process flow diagram is shown.

[0041] Figure 1In the process, residual oil feedstock 1 is mixed with hydrogen 41 and recycled hydrogen 40 and enters the first reaction zone 2 of the fluidized bed. The gas-liquid components generated by the reaction are separated in the interstage separator 3 to obtain the first liquid phase 4 and the first gas phase 5. The first liquid phase and recycled hydrogen are mixed and then enter the second reactor 6 of the fluidized bed. The gas-liquid components generated by the reaction are separated in the thermal high-efficiency separator 7 to obtain the second liquid phase 8 and the second gas phase 9. The second liquid phase 8 enters the thermal low-efficiency separator 10 to obtain the thermal low-efficiency liquid phase 11 and the thermal low-efficiency gas phase 12. The first gas phase 5 and the second gas phase 9 are mixed and enter the cold high-efficiency separator 13 to obtain the cold high-efficiency liquid phase 14 and the cold high-efficiency gas phase 15. The cold high-efficiency gas phase 15 passes through the recycled hydrogen desulfurization tower 39 to remove impurities such as hydrogen sulfide to obtain recycled hydrogen 40. The hot low-efficiency gas phase 12 and the cold high-efficiency liquid phase 14 are mixed and enter the cold low-efficiency separator 16 to obtain the cold low-efficiency liquid phase 17 and the cold low-efficiency gas phase 18. The hot low-efficiency liquid phase 11 and the cold high-efficiency gas phase 18 are mixed and then enter the cold low-efficiency separator 16 to obtain the cold low-efficiency liquid phase 17 and the cold low-efficiency gas phase 18. The low-volume liquid phase 17 is fed into the fractionation system 19 to separate fluidized bed gas 20, fluidized bed naphtha 21, fluidized bed diesel 22, fluidized bed wax oil 23, and unconverted oil 24. The unconverted oil 24 is fed into the first extraction tower 25, where it is reacted with the first extraction solvent 39 to obtain raffinate phase 26 and a mixture of extract phase and solvent 27. 27 is fed into the first extraction solvent recovery tower 37 to separate the first extraction solvent 39 and the first extract phase 38. The first raffinate phase 26 is reacted with the second solvent 32 to obtain the second raffinate phase 29 and a mixture of the second extract phase and solvent 30. 30 is fed into the second solvent recovery tower 31 to obtain the second solvent 32 and the second extract phase 33. The second extract phase 33 is fed into the coking tower 34 to undergo a coking reaction to obtain petroleum coke 35 and a coking gas-liquid mixture 36. 36 is fed into the fractionation system 19 for separation.

[0042] Table 1 lists the properties of the residue feedstocks in the examples and comparative examples.

[0043] Table 1 Properties of Residue Oil Feedstock project vacuum residue <![CDATA[Density (at 20°C), kg / m 3 > 1032.7 Viscosity (150℃), mPa·s 175 Residual charcoal, wt% 22.3 Sulfur content, wt% 5.2 Nitrogen content, wt% 0.39 Metal (Ni), mg / kg 87 Metal (V), mg / kg 157 Saturated fraction, wt% 7.7 Aromatic components, wt% 60.97 Gel, wt% 16.11 Asphalt, wt% 15.22 Example 1

[0044] The first and second fluidized bed hydrogenation reaction zones are each equipped with one fluidized bed hydrogenation reactor. The fluidized bed hydrogenation reactor is the STRONG fluidized bed reactor with a built-in three-phase separator developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd. Operating conditions in the first fluidized bed reaction zone: reaction temperature 430℃, reaction pressure 18.5MPa, hydrogen-to-oil volume ratio 400, liquid hourly space velocity 0.40h⁻¹ -1 ; Operating conditions in the second fluidized bed reaction zone: reaction temperature 432℃, reaction pressure 17.5MPa, hydrogen-to-oil volume ratio 400, liquid hourly space velocity 0.38h⁻¹ -1 ; The first reaction zone of the fluidized bed uses a microsphere catalyst with alumina as the support and molybdenum and nickel as active metals. Based on the catalyst mass, the support accounts for 95 wt%, and the active metals, calculated as oxides, account for 4 wt%; the bulk density is 0.6 g / cm³. 3 Mesopores with a diameter greater than or equal to 3 nm and less than 100 nm account for 60% of the total pore volume, while macropores with a diameter of 100~1000 nm account for 40% of the total pore volume. Among them, mesopores have a diameter of approximately 20 nm, and macropores have a diameter of approximately 800 nm.

[0045] The second reaction zone of the fluidized bed uses a microsphere catalyst with alumina as the support and molybdenum and nickel as the active metals. Based on the catalyst mass, the support accounts for 85 wt%, and the active metal, calculated as oxides, accounts for 15 wt%; the bulk density is 0.7 g / cm³. 3 Mesopores with a diameter greater than or equal to 3 nm and less than 100 nm account for 75% of the total pore volume, while macropores with a diameter of 100~1000 nm account for 25% of the total pore volume. Among them, mesopores have a diameter of approximately 10 nm, and macropores have a diameter of approximately 600 nm.

[0046] The catalyst addition and removal adopts a parallel addition and removal mode, with the first reaction catalyst addition rate being 0.8 kg / ton of fresh residue feedstock and the second reaction catalyst addition rate being 0.5 kg / ton of fresh residue feedstock.

[0047] The first extraction zone uses a packed tower filled with corrugated plate structured packing; propane is selected as the solvent; the operating conditions are as follows: extraction temperature is 65℃, extraction pressure is 4.3MPa, and solvent-to-oil volume ratio is 5.0.

[0048] The second extraction zone uses a packed tower filled with corrugated plate structured packing; the solvent selected is n-butane; the operating conditions are as follows: extraction temperature is 135℃, extraction pressure is 4.0MPa, and solvent-to-oil volume ratio is 5.0.

[0049] The second extract phase is partially returned to the second reaction zone of the fluidized bed, accounting for 4% of the fresh feed to the unit.

[0050] The coking reaction unit is equipped with one heating furnace and two coking towers. The outlet temperature of the heating furnace is 500℃, the pressure of the coking tower is 0.20 MPa, and the circulation ratio is 0.15. Example 2

[0051] The process flow and reactor are the same as in Example 1, the difference lies in the process conditions.

[0052] Operating conditions in the first fluidized bed reaction zone: reaction temperature 425℃, reaction pressure 19MPa, hydrogen-to-oil volume ratio 400, liquid hourly space velocity 0.30h⁻¹ -1 ; Operating conditions in the second fluidized bed reaction zone: reaction temperature 428℃, reaction pressure 18MPa, hydrogen-to-oil volume ratio 500, liquid hourly space velocity 0.28h⁻¹ -1 ; The first reaction zone of the fluidized bed uses a microsphere catalyst with alumina as the support and molybdenum and nickel as active metals. Based on the catalyst mass, the support accounts for 95 wt%, and the active metals, calculated as oxides, account for 4 wt%; the bulk density is 0.6 g / cm³. 3 Mesopores with a diameter greater than or equal to 3 nm and less than 100 nm account for 60% of the total pore volume, while macropores with a diameter of 100~1000 nm account for 40% of the total pore volume. Among them, mesopores have a diameter of approximately 20 nm, and macropores have a diameter of approximately 800 nm.

[0053] The second reaction zone of the fluidized bed uses a microsphere catalyst with alumina as the support and molybdenum and nickel as the active metals. Based on the catalyst mass, the support accounts for 85 wt%, and the active metal, calculated as oxides, accounts for 15 wt%; the bulk density is 0.7 g / cm³. 3 Mesopores with a diameter greater than or equal to 3 nm and less than 100 nm account for 75% of the total pore volume, while macropores with a diameter of 100~1000 nm account for 25% of the total pore volume. Among them, mesopores have a diameter of approximately 10 nm, and macropores have a diameter of approximately 600 nm.

[0054] The catalyst addition and removal adopts a parallel addition and removal mode, with the first reaction catalyst addition rate being 0.7 kg / ton of fresh residue feedstock and the second reaction catalyst addition rate being 0.5 kg / ton of fresh residue feedstock.

[0055] The first extraction zone uses a packed tower filled with corrugated plate structured packing; propane is selected as the solvent; the operating conditions are as follows: extraction temperature is 65℃, extraction pressure is 4.3MPa, and solvent-to-oil volume ratio is 5.0.

[0056] The second extraction zone uses a packed tower filled with corrugated plate structured packing; the solvent selected is n-pentane; the operating conditions are as follows: extraction temperature is 175℃, extraction pressure is 3.7MPa, and solvent-to-oil volume ratio is 5.0.

[0057] The second extract phase is partially returned to the second reaction zone of the fluidized bed, accounting for 4% of the fresh feed to the unit.

[0058] The coking reaction unit is equipped with one heating furnace and two coking towers. The outlet temperature of the heating furnace is 495℃, the pressure of the coking tower is 0.18 MPa, and the circulation ratio is 0.15. Example 3

[0059] The process flow and reactor are the same as in Example 1, the difference lies in the process conditions and catalyst performance.

[0060] Operating conditions in the first fluidized bed reaction zone: reaction temperature 430℃, reaction pressure 18.5MPa, hydrogen-to-oil volume ratio 400, liquid hourly space velocity 0.40h⁻¹ -1 ; Operating conditions in the second fluidized bed reaction zone: reaction temperature 432℃, reaction pressure 17.5MPa, hydrogen-to-oil volume ratio 400, liquid hourly space velocity 0.38h⁻¹ -1 ; The first reaction zone of the fluidized bed uses a microsphere catalyst with alumina as the support and molybdenum and nickel as active metals. Based on the catalyst mass, the support accounts for 94.5 wt%, and the active metals, calculated as oxides, account for 5.5 wt%; the bulk density is 0.57 g / cm³. 3 Mesopores with a diameter greater than or equal to 3 nm and less than 100 nm account for 55% of the total pore volume, while macropores with a diameter of 100~1000 nm account for 45% of the total pore volume. Among them, mesopores have a diameter of approximately 20 nm, and macropores have a diameter of approximately 850 nm.

[0061] The second reaction zone of the fluidized bed uses a microsphere catalyst with alumina as the support and molybdenum and nickel as the active metals. Based on the catalyst mass, the support accounts for 85 wt%, and the active metal, calculated as oxides, accounts for 15 wt%; the bulk density is 0.7 g / cm³. 3 Mesopores with a diameter greater than or equal to 3 nm and less than 100 nm account for 75% of the total pore volume, while macropores with a diameter of 100~1000 nm account for 25% of the total pore volume. Among them, mesopores have a diameter of 8 nm and macropores have a diameter of 600 nm.

[0062] The catalyst addition and removal adopts a parallel addition and removal mode, with the first reaction catalyst addition rate being 0.9 kg / ton of fresh residue feedstock and the second reaction catalyst addition rate being 0.5 kg / ton of fresh residue feedstock.

[0063] The first extraction zone uses a packed tower filled with corrugated plate structured packing; propane is selected as the solvent; the operating conditions are as follows: extraction temperature is 65℃, extraction pressure is 4.3MPa, and solvent-to-oil volume ratio is 5.0.

[0064] The second extraction zone uses a packed tower filled with corrugated plate structured packing; the solvent selected is n-butane; the operating conditions are as follows: extraction temperature is 135℃, extraction pressure is 4.0MPa, and solvent-to-oil volume ratio is 5.0.

[0065] The second extract phase is partially returned to the second reaction zone of the fluidized bed, accounting for 4% of the fresh feed to the unit.

[0066] The coking reaction unit is equipped with one heating furnace and two coking towers. The outlet temperature of the heating furnace is 500℃, the pressure of the coking tower is 0.20 MPa, and the circulation ratio is 0.15. Example 4

[0067] Compared with Example 1, the difference is that the second extract phase is not returned to the second reaction zone of the fluidized bed, but is entirely fed into the coking unit. Example 5

[0068] Compared with Example 2, the difference is that the second extract phase is not returned to the second reaction zone of the fluidized bed, but is entirely fed into the coking unit.

[0069] Comparative Example 1 Compared to Example 1, the process flow is that the unconverted oil 24 directly enters the coking tower 34 for coking reaction, without undergoing the two-stage solvent extraction process. Everything else is the same as in Example 1.

[0070] Table 2. Product yield and properties in different embodiments and comparative examples. project Example 1 Example 2 Example 3 Example 4 Example 5 Comparative Example 1 Unconverted oil properties in fluidized bed n-Heptane asphaltenes content, % 6.6 8.5 7.0 6.9 8.8 6.6 The sum of aromatic components and gum content, % 72.5 71.0 74.3 71.0 69.8 72.5 Aromatic component / colloid ratio 6.6 5.5 5.0 6.1 5.1 6.6 Sulfur content, % 0.71 0.80 0.73 0.73 0.84 0.71 Extraction phase yield (relative to fresh feed), % 5.1 5.2 4.9 4.8 4.9 Extraction phase residue, % 1.99 1.98 2.01 2.10 2.11 Bright oil yield (relative to fresh feed), % 4.08 4.28 4.01 3.81 4.02 Petroleum coke yield (relative to fresh feed), wt% 6.05 7.61 6.63 6.14 7.71 11.20 Properties of petroleum coke Sulfur content, wt% 1.11 1.25 1.14 1.18 1.31 1.46 Volatile matter, wt% 2.7 2.9 2.5 2.7 2.9 4.5 Ash content, wt% 0.21 0.22 0.21 0.22 0.22 0.49 <![CDATA[Specific capacity, mA·h·g -1 > 357 356 355 355 352 324 Graphitization degree, % 94.6 94.4 94.4 94.3 94.2 89.6 First Coulomb efficiency, % 95.3 95.1 95.2 95.2 95.0 92.1

Claims

1. A method for efficient conversion of residue oil, characterized in that: The technology includes the following: (1) Unconverted oil is obtained by hydrotreating the residue oil feedstock in a fluidized bed; the unconverted oil is controlled to have the following properties: the mass content of n-heptane asphaltene is less than 15%, preferably 1%~10%, the sum of the mass content of aromatics and gums is ≥70%, preferably 80%-95%, the ratio of aromatics / gum is ≥5, preferably 6-10, and the mass content of sulfur is less than 1.5%, preferably 0.3%~1.0%; (2) The unconverted oil in step (1) enters the first solvent extraction unit to obtain a mixture of the first raffinate phase, the first solvent and the extract phase. Then the first solvent and the extract phase are separated to recover the first solvent and obtain the first extract phase rich in alkane components. (3) In step (2), the first raffinate enters the second solvent extraction unit to obtain a mixture of the second raffinate, the second solvent and the extract phase. Then the second solvent and the extract phase are separated to recover the second solvent and obtain the second extract phase rich in aromatic components. (4) In step (3), the second extraction phase enters the coking unit to carry out the coking reaction, and obtains the special coke for negative electrode material and the gas-liquid mixture. The gas-liquid mixture is returned to the fluidized bed fractionation unit to be mixed and fractionated with the fluidized bed product.

2. The method according to claim 1, characterized in that: In step (1), the residue oil feedstock enters the first reaction zone of the fluidized bed. The gas-liquid mixture generated in the first reaction zone passes through an interstage separator to obtain the first gas phase and the first liquid phase. The first liquid phase enters the second reaction zone to react and obtain a gas-liquid mixture, which is then separated in the hot high-efficiency fraction to obtain the second gas phase and the second liquid phase. The first gas phase and the second gas phase are mixed and enter the cold high-efficiency fraction. The gas phase from the cold high-efficiency fraction enters the circulating hydrogen desulfurization tower to remove hydrogen sulfide and then obtains circulating hydrogen, which then enters the first and second reactors. The second liquid phase enters the hot low-efficiency fraction. The gas phase from the hot low-efficiency fraction and the liquid phase from the cold high-efficiency fraction are mixed and enter the cold low-efficiency fraction. The liquid phase from the cold low-efficiency fraction and the liquid phase from the hot low-efficiency fraction are mixed and enter the fractionation system to obtain fluidized bed gas, naphtha, diesel oil, wax oil, and unconverted oil.

3. The method according to claim 2, characterized in that: The residue oil raw material involved in step (1) is at least one of atmospheric residue oil, vacuum residue oil, and heavy oil.

4. The method according to claim 2, characterized in that: In step (1), a fluidized bed reactor is set up in each of the two reaction zones of the fluidized bed hydrogenation unit. The fluidized bed used is either a circulating pump fluidized bed reactor or a three-phase separator fluidized bed reactor.

5. The method according to claim 2, characterized in that: The process parameters for the first reaction zone of the fluidized bed in step (1) are as follows: reaction temperature is 370–440℃, preferably 400–430℃; reaction pressure is 15.0–20.0 MPa, preferably 16.0–18.0 MPa; hydrogen-to-oil volume ratio is 300–1000, preferably 400–600; and liquid hourly space velocity is 0.1–2.0 h⁻¹. -1 Preferably, it is 0.2 to 0.5 h. -1 .

6. The method according to claim 2, characterized in that: The process parameters for the second reaction zone of the fluidized bed in step (1) are as follows: reaction temperature is 370–440℃, preferably 400–430℃; reaction pressure is 15.0–20.0 MPa, preferably 16.0–18.0 MPa; hydrogen-to-oil volume ratio is 300–1000, preferably 400–600; and liquid hourly space velocity is 0.1–2.0 h⁻¹. -1 Preferably, it is 0.2 to 0.5 h. -1 .

7. The method according to claim 2, characterized in that: The catalyst used in the first reaction zone of the fluidized bed is a spherical catalyst with a diameter of 0.3–0.7 mm. The catalyst comprises a support and an active metal, wherein the active metal can be one or more of nickel, molybdenum, or tungsten; the support can be one or more of alumina, silica, or alumina-silica; based on catalyst mass, the support accounts for 90 wt%–96 wt%, and the active metal, calculated as oxides, accounts for 3 wt%–10 wt%. The bulk density of the catalyst is 0.4–0.8 g / cm³. 3 The catalyst has a bimodal pore structure, with mesopores having a diameter greater than or equal to 3 nm and less than 100 nm accounting for 50% to 70% of the total pore volume, and macropores having a diameter of 100 to 1000 nm accounting for 30% to 50% of the total pore volume. Among them, the mesopores have a diameter of 15 to 25 nm, and the macropores have a diameter of 700 to 900 nm.

8. The method according to claim 2, characterized in that: The catalyst used in the second reaction zone of the fluidized bed is a spherical catalyst with a diameter of 0.3–0.7 mm. The catalyst comprises a support and an active metal, wherein the active metal can be one or more of nickel, molybdenum, or tungsten; the support can be one or more of alumina, silica, or alumina-silica; based on catalyst mass, the support accounts for 80 wt%–90 wt%, and the active metal, calculated as oxides, accounts for 10 wt%–20 wt%. The bulk density of the catalyst is 0.4–0.8 g / cm³. 3 The catalyst has a bimodal pore structure, in which pores with a diameter greater than or equal to 1 nm and less than 200 nm account for 70% to 80% of the total pore volume, and pores with a diameter of 200 to 1000 nm account for 20% to 30% of the total pore volume. Among them, mesopores have a diameter of 7 to 13 nm and macropores have a diameter of 500 to 700 nm.

9. The method according to claim 2, characterized in that: Step (1) The catalysts in the two reaction zones of the fluidized bed are added in parallel. The catalyst addition rate of the first reaction zone is 0.3 to 1.5 kg / ton of raw material, preferably 0.5 to 0.9 kg / ton of raw material; the catalyst addition rate of the second reaction zone is 0.3 to 1.5 kg / ton of raw material, preferably 0.3 to 0.7 kg / ton of raw material.

10. The method according to claim 2, characterized in that: Step (1) The average catalyst bed temperature in the two reaction zones of the fluidized bed is controlled at a temperature gradient of 2~6℃, and the temperature of the second reaction zone is higher than that of the second reaction zone.

11. The method according to claim 2, characterized in that: Step (1) The temperature rise ratio of the two reaction zones in the fluidized bed is controlled to be ≥1.5, preferably ≥1.

7.

12. The method according to claim 1, characterized in that: Step (1) The fractionation section of the fluidized bed unit is divided into atmospheric fractionation and vacuum fractionation, wherein the atmospheric and vacuum fractionation points are 300~400℃, preferably 330~360℃; the atmospheric fraction can obtain hydrotreated naphtha, hydrotreated diesel and atmospheric residue, the cut-off points of hydrotreated naphtha and hydrotreated diesel are 160~210℃, preferably 180~200℃; the atmospheric residue is further fractionated into vacuum to obtain hydrotreated wax oil and vacuum residue, the cut-off points of hydrotreated wax oil and vacuum residue are 450~560℃, preferably 500~540℃.

13. The method according to claim 1, characterized in that: Step (2) involves using a solvent in the solvent extraction unit that is an alkane with 3 to 4 carbon atoms, wherein the mass fraction of alkane with 3 carbon atoms in the solvent is not less than 80%; the operating conditions are an extraction temperature of 40 to 100°C, preferably 60 to 90°C, an extraction pressure of 4.0 to 7.0 MPa, preferably 4.5 to 5.0 MPa, and a solvent-to-oil volume ratio of 3.0 to 7.0, preferably 4.0 to 6.

0.

14. The method according to claim 1, characterized in that: Step (2) uses the carbon residue content in the first extraction phase as the control condition, and the carbon residue in the extraction phase is controlled to be no higher than 2.0%.

15. The method according to claim 1, characterized in that: In step (2), the extract phase is processed by hydrorefining-isomerization dewaxing, furfural extraction-ketone-benzene dewaxing-clay refining process or used to generate low-sulfur marine fuel blending components or as feedstock for hydrocracking and catalytic cracking.

16. The method according to claim 1, characterized in that: Step (2) Solvent extraction is carried out in an extraction tower, which can be a packed tower or a tray tower, preferably a packed tower.

17. The method according to claim 1, characterized in that: In step (2), the solvent is recycled and reused in the solvent extraction unit. The solvent recycling is carried out under supercritical conditions.

18. The method according to claim 1, characterized in that: Step (3) involves using an alkane with 4 to 6 carbon atoms as the solvent in the solvent extraction unit, wherein the mass fraction of alkane with 4 carbon atoms in the solvent is not less than 50%; the operating conditions are an extraction temperature of 100 to 160°C, preferably 120 to 140°C, an extraction pressure of 3.0 to 7.0 MPa, preferably 3.5 to 4.5 MPa, and a solvent-to-oil volume ratio of 3.0 to 7.0, preferably 4.0 to 6.

0.

19. The method according to claim 1, characterized in that: Step (3) uses the C7 asphaltene content in the second extraction phase as the control condition, and the C7 asphaltene content in the extraction phase is controlled to be no higher than 1.0%.

20. The method according to claim 1, characterized in that: Step (3) The second extract phase is returned to the fluidized bed reaction zone, preferably to the second fluidized bed reaction zone, and the return ratio is 0-10% of the fresh feed of the device, preferably 2-5%.

21. The method according to claim 1, characterized in that: In step (3), the second raffinate phase is used as a raw material for hydrogen production or asphalt production.

22. The method according to claim 1, characterized in that: The coking unit in step (4) includes a heating furnace, a coking tower, and a fractionation tower, and at least one heating furnace, two coking towers, and one fractionation tower. The coking reaction includes two processes: a continuous coking process and an intermittent coking removal process; the outlet temperature of the heating furnace is 480~520℃, preferably 490℃~510℃; the pressure of the coking tower is 0.1MPa~0.5MPa, preferably 0.15~0.25MPa; and the recycle ratio is 0~1.0, preferably 0~0.3, by mass.

23. The method according to claim 1, characterized in that: In step (4), the coking-generated oil and gas enter the fluidized bed fractionation unit.

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