A combined process for efficient conversion of heavy oil

Abstract

The application discloses a heavy oil efficient conversion combined technology, and the process comprises the following contents: (1) obtaining unconverted oil after heavy feedstock is subjected to ebullated bed hydroprocessing; the unconverted oil has the following properties: the mass content of n-heptane asphaltene is 1.0%-9.0%, preferably 5%-8%; the sulfur content is 0.3%-1.5%, preferably 0.8%-1.2%; the metal (Ni+V+Fe) content is 30 mg / kg-120 mg / kg, preferably 60 mg / kg-90 mg / kg; (2) the unconverted oil in step (1) enters a solvent extraction unit to obtain a raffinate phase, a solvent and a mixed stream of an extract phase, then the solvent and the extract phase are separated to recover the solvent and obtain the extract phase of a rich alkane component; (3) the raffinate phase in step (2) enters a coking unit to perform green coke reaction, and low-sulfur petroleum coke and a gas-liquid mixed stream are obtained, the gas-liquid mixed stream enters a coking separation unit to obtain coking gas, coking naphtha, coking diesel and coking wax oil. The method can greatly improve the overall conversion rate of residual oil, and simultaneously produce negative electrode coke and high-viscosity-index high-quality heavy lubricating oil raw materials.

Description

Technical Field

[0001] This invention belongs to the field of petrochemical technology, and in particular relates to a combined technology for the efficient conversion of heavy oil. By organically combining fluidized bed hydrotreating, solvent extraction and coking processes, the overall conversion rate of residue oil can be greatly improved, while obtaining high-value specialty chemical products. Background Technology

[0002] The green development concept has driven technological advancements in new energy vehicles, leading to a surge in sales and a significant impact on traditional oil refining structures. This has resulted in insufficient consumption of gasoline and diesel fuels, causing a severe decline in refinery profitability. However, demand for chemical raw materials remains strong, particularly for negative electrode coke materials and high-viscosity-index heavy lubricating oils. The primary reason for this shortage is the lack of suitable high-quality feedstocks. Suitable feedstocks for negative electrode coke materials typically have low sulfur impurity content and a high proportion of aromatics, while suitable feedstocks for high-viscosity-index heavy lubricating oils have low impurity content and a high proportion of isoalkanes. These feedstocks are relatively scarce or have high procurement costs, such as catalytic slurry oil and low-sulfur crude oil. Currently, to improve economic efficiency, refineries tend to purchase cheaper, high-sulfur, low-quality crude oil, directly hindering the production of high-value negative electrode coke materials and high-viscosity-index heavy lubricating oils using existing conventional processing technologies.

[0003] Heavy oil systems are colloidal structures composed of alkanes, aromatics, gums, and asphaltenes. Alkanes are ideal components for lubricating oil fractions, aromatics and gums are ideal for negative electrode coke feedstocks, while asphaltenes are non-ideal components. Currently, there are few reports on technologies for improving asphaltenes conversion while retaining aromatics and gums, and for increasing the production of isoalkanes. Such technologies use high-sulfur, low-quality residual oil as feedstock to simultaneously produce bright oil and negative electrode coke feedstock.

[0004] CN108473889A provides a method for forming lubricant base oils from feedstocks, such as vacuum residue or other feedstocks at 510°C+. The feedstock can be deasphalted and then catalytically and / or solvent-processed 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 use directly from low-quality vacuum residue as feedstock to produce bright oil products.

[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 combined technology for the efficient conversion of heavy oil. By organically combining fluidized bed hydrotreating, solvent extraction, and coking processes, the overall conversion rate of residue oil can be significantly improved, while simultaneously producing high-quality feedstocks such as negative electrode coke and high-viscosity index heavy lubricating oil.

[0007] This invention provides a combined technology for high-efficiency conversion of heavy oil, the process comprising the following: (1) The residue oil feedstock is hydrotreated by fluidized bed to obtain unconverted oil; the unconverted oil is controlled to have the following properties: n-heptane asphaltene content of 1.0%~9.0%, preferably 5%-8%; sulfur content of 0.3%~1.5%, preferably 0.8%-1.2%; metal (Ni+V+Fe) content of 30mg / kg~120mg / kg, preferably 60mg / kg~90mg / kg; (2) The unconverted oil in step (1) enters the solvent extraction unit to obtain a mixture of raffinate, solvent and extract phase. Then the solvent and extract phase are separated to recover the solvent and obtain an extract phase rich in alkane components. (3) In step (2), the residual phase of the raffinate enters the coking unit to carry out the coking reaction, and low-sulfur petroleum coke and gas-liquid mixture are obtained. The gas-liquid mixture enters the coking separation unit to obtain coking gas, coking naphtha, coking diesel and coking wax oil.

[0008] In the above-mentioned high-efficiency conversion process of heavy oil, in step (1), the residue feedstock enters the first reaction zone of the fluidized bed. The gas-liquid mixture generated in the first reaction zone is separated into a first gas phase and a first liquid phase by an interstage separator. 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 of 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 of the hot low-efficiency fraction and the liquid phase of the cold high-efficiency fraction are mixed and enter the cold low-efficiency fraction. The liquid phase of the cold low-efficiency fraction and the liquid phase of 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 high-efficiency conversion process of heavy oil, the residue oil raw material involved in step (1) can be at least one of the inferior heavy oils such as atmospheric residue oil, vacuum residue oil, and heavy oil.

[0010] In the above-mentioned high-efficiency conversion process of heavy oil, in step (1), the two reaction zones of the fluidized bed hydrogenation unit are each equipped with a fluidized bed reactor. 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 combined process for high-efficiency conversion of heavy 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 high-efficiency heavy oil conversion combined process, 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 high-efficiency heavy oil conversion combined process, step (1) involves loading the two reaction zones of the fluidized bed with a fluidized bed hydrogenation catalyst. The catalyst comprises a support and an active metal, wherein the active metal can be one or more of nickel, cobalt, molybdenum, or tungsten; the support can be one or more of alumina, silica, or alumina-silica. Based on the catalyst mass, the support accounts for 85wt%~96wt%, and the active metal, calculated as oxides, accounts for 4wt%~15wt%. The bulk density of the catalyst is 0.3~0.8 g / cm³. 3 The particle diameter (spherical diameter or strip diameter) is 0.3 to 1.0 mm; the catalyst has a bimodal pore structure, wherein the pore volume of pores with a diameter greater than or equal to 1 nm and less than 100 nm accounts for 50% to 80% of the total pore volume, and the pore volume of pores with a diameter greater than or equal to 100 nm, preferably 100 to 1500 nm, accounts for 30% to 40% of the total pore volume.

[0014] In the above-mentioned combined process for high-efficiency conversion of heavy oil, in step (1), the catalysts in the two reaction zones of the fluidized bed are added and discharged online, and are added and discharged in series or in parallel. The fresh catalyst addition rate is 0.3 to 2.0 kg / ton of raw material, preferably 0.8 to 1.2 kg / ton of raw material.

[0015] In the above-mentioned high-efficiency heavy oil conversion combined process, the property control of the unconverted oil in step (1) of the fluidized bed is achieved by monitoring the temperature rise of the two reaction zones of the fluidized bed and adjusting the catalyst addition mode. In terms of monitoring: the ratio of the temperature rise of the two reaction zones (average bed temperature minus the inlet temperature of the reaction zone) is controlled to be no less than 1.3, preferably above 1.7; in terms of catalyst addition and discharge, when the reaction temperature rise ratio is lower than the set value, it is necessary to add a parallel addition mode of the first reaction (in series mode) or increase the addition rate of the first reaction (in parallel addition mode).

[0016] In the above-mentioned combined process for high-efficiency conversion of heavy oil, in step (1), the average catalyst bed temperature (arithmetic mean) 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 high-efficiency heavy oil conversion combined process, the fractionation part of the fluidized bed unit in step (1) 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 point of hydrotreated naphtha and hydrotreated diesel is 160~210℃, preferably 180~200℃; the atmospheric residue is further fractionated into vacuum to obtain hydrotreated wax oil and vacuum residue, the cut-off point of hydrotreated wax oil and vacuum residue is 450~560℃, preferably 500~540℃.

[0018] In the above-mentioned high-efficiency conversion process of heavy oil, step (2) involves solvent extraction unit using alkane solvent with 3 to 5 carbon atoms; the operating conditions are extraction temperature of 40 to 200°C, preferably 60 to 140°C, extraction pressure of 4.0 to 7.0 MPa, preferably 4.5 to 5.0 MPa, and solvent-to-oil volume ratio of 3.0 to 7.0, preferably 4.0 to 6.0.

[0019] In the above-mentioned high-efficiency conversion process of heavy 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 extract phase. In this invention, the residual carbon content in the alkane-rich extract phase is used as the control condition, and the residual carbon in the extract phase is controlled to be ≤2.0%. Under this premise, the yield of the extract phase is increased as much as possible.

[0020] In the above-mentioned high-efficiency conversion process of heavy 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.

[0021] In the above-mentioned high-efficiency conversion combination process of heavy oil, the extract phase in step (2) can also be used as a raw material for catalytic cracking, hydrocracking, etc. When processing it, the residual carbon in the extract phase is controlled to be less than 8.0%. When the raffinate is used to produce low-sulfur coke, it needs to be blended with low-sulfur components such as catalytic oil slurry.

[0022] In the above-mentioned combined process for high-efficiency conversion of heavy oil, the solvent extraction in step (2) is carried out in an extraction tower, which can be a packed tower or a tray tower, with a packed tower being preferred.

[0023] In the above-mentioned high-efficiency conversion process of heavy oil, the solvent is recycled and reused in the solvent extraction unit in step (2). The solvent recycling is carried out under supercritical conditions, and the specific recycling conditions are set according to the properties of the selected solvent.

[0024] In the above-mentioned high-efficiency heavy oil conversion combined process, the coking unit in step (3) includes a heater, a coking tower, and a fractionation tower, including at least one heater, two coking towers, and one fractionation tower. The coking reaction includes two processes: continuous coking and intermittent coking removal. The outlet temperature of the heater is 480~520℃, preferably 490℃~510℃; the pressure of the coking tower is 0.1MPa~0.5MPa, preferably 0.15~0.25MPa; the recycle ratio is 0.05~1.0, preferably 0.1~0.3, by mass.

[0025] In the above-mentioned high-efficiency heavy oil conversion process, in step (3), petroleum coke is generated in a coking tower, and the coking-generated oil and gas are separated in a fractionation tower to obtain coking dry gas, coking naphtha, coking diesel oil, and coking wax oil. The cut-off points of coking naphtha and coking diesel oil are 160~210℃, preferably 180~200℃; the cut-off points of coking diesel oil and coking wax oil are 300~400℃, preferably 330~360℃.

[0026] In the above-mentioned high-efficiency conversion process of heavy oil, the coking wax oil in step (3) is also completely sent out of the unit, or it can be partially or completely returned to the first and second reaction zones of the fluidized bed. The specific settings need to be determined based on the processing raw materials, operating conditions and catalyst activity.

[0027] Compared with existing technologies, the high-efficiency heavy oil conversion combined process of the present invention has the following advantages: 1. This invention employs a combined process of fluidized bed hydrogenation, solvent extraction, and coking, which can significantly improve the overall conversion rate of residual oil and simultaneously obtain high-value specialty chemical products.

[0028] 2. In the fluidized bed hydrogenation unit, based on the fluidized bed thermal cracking and catalytic hydrogenation mechanism, a large-pore, low-cracking functional catalyst is employed. Simultaneously, through precise catalyst activity control methods, the conversion of non-ideal asphaltene macromolecules is improved, while effectively retaining high-quality coking components (polycyclic aromatic hydrocarbons) and high-quality bright oil components (long-chain isoalkanes). This invention uses inferior residue oil feedstock to generate high-value-added bright oil feedstock and coke specifically for negative electrode materials, successfully overcoming the limitation of insufficient high-quality feedstock in existing technologies for producing high-end products.

[0029] 3. The present invention returns coking wax oil to the fluidized bed unit for conversion, which has two advantages: (1) Low viscosity coking wax oil will reduce the mass transfer resistance of the system and increase the diffusion rate of macromolecular asphaltene; (2) The aromatic content in coking wax oil is relatively high, which plays a role in enhancing the stability of the system in the residue oil system, avoiding the aggregation of macromolecular asphaltene, and improving the asphaltene removal efficiency. Attached Figure Description

[0030] Figure 1 This is a schematic diagram of a combined process for the efficient conversion of heavy oil.

[0031] 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-Fractioning system; 20-Fluidized bed gas; 21-Stone 22-Breath oil; 23-Diesel oil; 24-Wax oil; 25-Unconverted oil; 26-Extraction tower; 27-Raffinate phase; 28-Mixed extractant and solvent stream; 29-Coking tower; 30-Petroleum coke; 31-Coking gas-liquid stream; 32-Coking fractionation system; 33-Coking gas; 34-Coking naphtha; 35-Coking diesel oil; 36-Coking wax oil; 37-Extraction solvent recovery tower; 38-Extraction phase; 39-Circulating hydrogen desulfurization tower; 40-Circulating hydrogen; 41-New hydrogen. Detailed Implementation

[0032] 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.

[0033] Unless otherwise specified, all percentages (%) in the methods of this invention refer to mass percentages.

[0034] In this invention, the embodiments adopt the following... Figure 1 The process flow diagram is shown.

[0035] 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 4 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 hot high-efficiency separator 7 to obtain the second liquid phase 8 and the second gas phase 9. The second liquid phase 8 enters the hot low-efficiency separator 10 to obtain the hot low-efficiency liquid phase 11 and the hot 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 high-efficiency gas phase 18. Low-grade gas phase 18, hot low-grade liquid phase 11, and cold low-grade liquid phase 17 are fed into 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. Unconverted oil 24 is fed into extraction tower 25, where it is reacted with extraction solvent 38 to obtain raffinate phase 26 and a mixture of extract phase and solvent 27. The mixture of extract phase and solvent 27 is fed into extraction solvent recovery tower 36 to separate extraction solvent 38 and extract phase 37. Raffinate phase 26 is fed into coking tower 28 to obtain petroleum coke 29 and coking gas-liquid stream 30. Coking gas-liquid stream 30 is fed into coking fractionation system 31 to obtain coking dry gas 32, coked naphtha 33, coked diesel 34, and coking wax oil 35. Coking wax oil 35 is recycled to the first reaction zone 2 of the fluidized bed.

[0036] Table 1 shows the properties of the raw materials used in the examples and comparative examples.

[0037] Table 1 Properties of Residue Oil Feedstock project vacuum residue <![CDATA[Density (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

[0038] Each of the first and second fluidized bed hydrogenation reaction zones is equipped with a fluidized bed hydrogenation reactor, which is a fluidized bed reactor with a circulating pump.

[0039] Operating conditions in the first fluidized bed reaction zone: reaction temperature 425℃, reaction pressure 18.5MPa, 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 17.5MPa, hydrogen-to-oil volume ratio 400, liquid hourly space velocity 0.28h⁻¹ -1 ; The catalyst used in the examples was a strip-shaped catalyst with an average length of 3.6 mm. The catalyst support was alumina, and the active metals were molybdenum and nickel, with a molybdenum oxide content of 12%, a nickel oxide content of 2.5%, and a catalyst bulk density of 0.5 g / cm³.3 The catalyst has a bimodal pore structure, with pores of 3-50 mm accounting for 60% of the total pore volume and pores of 100-1000 nm accounting for 35% of the total pore volume.

[0040] In the fluidized bed reactor, the catalyst addition and removal primarily employs a series addition mode. Fresh catalyst is first added to the second reaction zone, then the equilibrium catalyst in the second reaction zone is removed and added to the first fluidized bed reaction zone, and then the equilibrium catalyst in the first fluidized bed reaction zone is removed again. Simultaneously, a parallel mode is periodically adopted within each cycle, where fresh catalyst is directly added to the first reaction zone.

[0041] Compared to the total feed, the catalyst addition rate is 1.1 kg / ton of raw material, and the proportion of fresh catalyst added directly to the primary reactor accounts for 15% of the total addition.

[0042] The 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.

[0043] 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 0.2.

[0044] A portion of the coking wax oil is returned to the first fluidized bed reactor, with a blending ratio of 5% (compared to fresh feed). Example 2

[0045] The first and second fluidized bed hydrogenation reaction zones are each equipped with one fluidized bed hydrogenation reactor. The reactors are STRONG fluidized bed reactors with built-in three-phase separators developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd.

[0046] Operating conditions in the first fluidized bed reaction zone: reaction temperature 428℃, 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 430℃, reaction pressure 17.5MPa, hydrogen-to-oil volume ratio 500, liquid hourly space velocity 0.36h⁻¹ -1 ; The first and second fluidized bed reactors utilize FEM-10 and FES-31 catalysts, respectively, developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd. The catalysts in the fluidized beds are added in parallel, with FEM-10 catalyst added to the first reactor and FES-31 catalyst added to the second reactor. The catalyst addition rate for the first reactor is 1.1 kg / ton of fresh residue feedstock, and the catalyst addition rate for the second reactor is 0.9 kg / ton of fresh residue feedstock.

[0047] The extraction zone uses a packed tower filled with corrugated plate structured packing; isobutane is selected as the solvent; the operating conditions are as follows: extraction temperature is 130℃, extraction pressure is 4.0MPa, and solvent-to-oil volume ratio is 5.0.

[0048] The coking reaction unit is equipped with one heating furnace and two coking towers. The outlet temperature of the heating furnace is 490℃, the pressure of the coking towers is 0.2MPa, and the circulation ratio is 0.15.

[0049] A portion of the coking wax oil is returned to the first fluidized bed reactor, with a blending ratio of 5% (compared to fresh feed). Example 3

[0050] The first and second fluidized bed hydrogenation reaction zones are each equipped with one fluidized bed hydrogenation reactor. The reactors are STRONG fluidized bed reactors with built-in three-phase separators developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd.

[0051] Operating conditions in the first fluidized bed reaction zone: reaction temperature 423℃, reaction pressure 18.5MPa, 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 426℃, reaction pressure 17.5MPa, hydrogen-to-oil volume ratio 500, liquid hourly space velocity 0.28h⁻¹ -1 ; The first and second fluidized bed reactors utilize FEM-10 and FES-31 catalysts, respectively, developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd. The catalysts in the fluidized beds are added in parallel, with FEM-10 catalyst added to the first reactor and FES-31 catalyst added to the second reactor. The catalyst addition rate for the first reactor is 1.2 kg / ton of fresh residue feedstock, and the catalyst addition rate for the second reactor is 0.8 kg / ton of fresh residue feedstock.

[0052] The 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 145℃, extraction pressure is 4.0MPa, and solvent-to-oil volume ratio is 5.0.

[0053] 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 0.18MPa, and the circulation ratio is 0.15.

[0054] A portion of the coking wax oil is returned to the second fluidized bed reactor, with a blending ratio of 5% (compared to fresh feed). Example 4

[0055] The difference from Example 3 lies in the extraction unit conditions and the coking unit conditions.

[0056] The extract phase is used as feedstock for catalytic cracking, and the raffinate phase is blended with catalytic oil slurry to produce low-sulfur petroleum coke.

[0057] The first and second fluidized bed hydrogenation reaction zones are each equipped with one fluidized bed hydrogenation reactor. The reactors are STRONG fluidized bed reactors with built-in three-phase separators developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd.

[0058] Operating conditions in the first fluidized bed reaction zone: reaction temperature 423℃, reaction pressure 18.5MPa, 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 426℃, reaction pressure 17.5MPa, hydrogen-to-oil volume ratio 500, liquid hourly space velocity 0.28h⁻¹ -1 ; The first and second fluidized bed reactors utilize FEM-10 and FES-31 catalysts, respectively, developed by Sinopec (Dalian) Petrochemical Research Institute Co., Ltd. The catalysts in the fluidized beds are added in parallel, with FEM-10 catalyst added to the first reactor and FES-31 catalyst added to the second reactor. The catalyst addition rate for the first reactor is 1.2 kg / ton of fresh residue feedstock, and the catalyst addition rate for the second reactor is 0.8 kg / ton of fresh residue feedstock.

[0059] The 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 136℃, extraction pressure is 4.0MPa, and solvent-to-oil volume ratio is 5.0.

[0060] The coking unit is equipped with one heater and two coking towers. The heater outlet temperature is 495℃, the coking tower pressure is 0.18MPa, and the circulation ratio is 0.15. The feed to the coking unit is a mixture of raffinate and catalytic oil slurry, with the catalytic oil slurry accounting for 25%.

[0061] A portion of the coking wax oil is returned to the first fluidized bed reactor, with a blending ratio of 5% (compared to fresh feed). Example 5

[0062] The difference from Example 1 is that the coking wax oil is not returned to the fluidized bed unit. Example 6

[0063] The difference from Example 3 is that the coking wax oil is not returned to the fluidized bed unit.

[0064] Table 2. Product yield and properties in different embodiments and comparative examples. project Example 1 Example 2 Example 3 Example 4 Example 5 Example 6 Unconverted oil properties in fluidized bed n-Heptane asphaltenes content, % 7 6 6.6 6.5 7.2 6.8 Sulfur content, % 1.05 0.95 1.04 1.10 1.13 1.14 Metal (Ni+V+Fe) content, mg / kg 80 70 77 75 85 80 Extraction phase yield (relative to fresh feed), % 5.3 5.0 4.9 19.5 4.7 4.5 Extraction phase residue, % 1.95 2.01 1.98 7.98 2.05 2.06 Bright oil yield relative to fresh feed) % 4.22 4.13 4.01 3.76 3.69 Petroleum coke yield (relative to fresh feed), wt% 7.12 6.89 7.53 3.15 7.29 7.65 Properties of petroleum coke Sulfur content, wt% 2.02 1.82 2.01 2.64 2.17 2.19 Volatile matter, wt% 2.6 2.6 2.5 3.7 2.6 2.5 Ash content, wt% 0.22 0.23 0.21 0.23 0.23 0.22 <![CDATA[Specific capacity, mA·h·g -1 > 335 336 335 341 332 335 Graphitization degree, % 92.3 92.2 92.9 94.6 91.8 92.8 First Coulomb efficiency, % 93.2 92.6 92.0 93.7 92.7 92.0

Claims

1. A combined technology for high-efficiency conversion of heavy oil, characterized in that: The process includes the following: (1) The residue oil feedstock is hydrotreated by fluidized bed to obtain unconverted oil; the unconverted oil is controlled to have the following properties: n-heptane asphaltene content of 1.0%~9.0%, preferably 5%-8%; sulfur content of 0.3%~1.5%, preferably 0.8%-1.2%; metal (Ni+V+Fe) content of 30mg / kg~120mg / kg, preferably 60mg / kg~90mg / kg; (2) The unconverted oil in step (1) enters the solvent extraction unit to obtain a mixture of raffinate, solvent and extract phase. Then the solvent and extract phase are separated to recover the solvent and obtain an extract phase rich in alkane components. (3) In step (2), the residual phase of the raffinate enters the coking unit to carry out the coking reaction, and low-sulfur petroleum coke and gas-liquid mixture are obtained. The gas-liquid mixture enters the coking separation unit to obtain coking gas, coking naphtha, coking diesel and coking wax oil.

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 1, characterized in that: The residue oil raw material involved in step (1) is at least one of the inferior heavy oils such as atmospheric residue oil, vacuum residue oil, and heavy oil.

4. The method according to claim 1, 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: Step (1) The two reaction zones of the fluidized bed are filled with fluidized bed hydrogenation catalyst, which includes a support and an active metal, wherein the active metal is one or more of nickel, cobalt, molybdenum or tungsten; and the support is one or more of alumina, silicon oxide or alumina-silicon oxide.

8. The method according to claim 7, characterized in that: Based on the mass of the hydrogenation catalyst, the support accounts for 85wt%~96wt%, and the active metal, calculated as oxide, accounts for 4wt%~15wt%.

9. The method according to claim 2, characterized in that: Step (1) The catalyst in the two reaction zones of the fluidized bed is added and drained online, either in series or in parallel. The fresh catalyst addition rate is 0.3 to 2.0 kg / ton of raw material, preferably 0.8 to 1.2 kg / ton of raw material.

10. 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℃.

11. The method according to claim 1, characterized in that: Step (2) involves using a solvent in the solvent extraction unit that is an alkane solvent with 3 to 5 carbon atoms; the operating conditions are an extraction temperature of 40 to 200°C, preferably 60 to 140°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.

12. The method according to claim 1, characterized in that: Step (2) The carbon residue in the extract phase is controlled to be less than 2.0%.

13. The method according to claim 1, characterized in that: The coking unit in step (3) includes a heating furnace, a coking tower and a fractionation tower, including at least one heating furnace, two coking towers and one fractionation tower.

14. The method according to claim 13, characterized in that: 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; the circulation ratio is 0.05~1.0, preferably 0.1~0.3, by mass.

15. The method according to claim 1, characterized in that: In step (3), petroleum coke is generated in the coking tower, and the coking-generated oil and gas are separated in the fractionation tower to obtain coking dry gas, coking naphtha, coking diesel oil and coking wax oil.

16. The method according to claim 15, characterized in that: The cutting point of coking naphtha and coking diesel is 160~210℃, preferably 180~200℃; the cutting point of coking diesel and coking wax oil is 300~400℃, preferably 330~360℃.

17. The method according to claim 1, characterized in that: Step (3) All coking wax oil leaves the device, or is partially or completely returned to the first and second reaction zones of the fluidized bed.

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