Method and system for producing ethylene and propylene from heavy oil

Abstract

The invention relates to a method and system for producing ethylene and propylene from heavy oil, and the method comprises the following steps: introducing heavy hydrocarbon oil into a supercritical extraction separation unit for supercritical extraction separation to respectively obtain light deasphalted oil, heavy deasphalted oil and deoiled asphalt; carrying out hydrotreating on the deoiled asphalt to obtain hydrogenated deoiled asphalt; introducing the heavy deasphalted oil and the hydrogenated deoiled asphalt into a first reactor to be in contact with a first catalytic conversion catalyst for a first catalytic conversion reaction; introducing the light deasphalted oil into a second reactor to contact with a second catalytic conversion catalyst for a second catalytic conversion reaction; wherein the first catalytic conversion catalyst and the second catalytic conversion catalyst contain MPZ hierarchical pore molecular sieves. According to the method disclosed by the invention, polycyclic aromatic hydrocarbon which is not easy to react can be saturated, the cracking property of the deoiled asphalt is improved, the generation of coke is reduced, and the yields of ethylene and propylene are improved.

Description

Technical Field

[0001] This invention relates to the field of petrochemicals, specifically to a method and system for producing ethylene and propylene from heavy oil. Background Technology

[0002] Ethylene and propylene are important basic chemical raw materials. Currently, the production of ethylene and propylene mainly relies on tubular steam cracking, using feedstocks including light petroleum hydrocarbons such as ethane, propane, butane, naphtha, light diesel oil, and condensate. Naphtha dominates the steam cracking feedstock, accounting for 50-55%, ethane accounts for 25-30%, while propane, butane, light diesel oil, and condensate oil each account for less than 10%.

[0003] As crude oil becomes increasingly heavier, the supply of light petroleum hydrocarbons can no longer meet the demand for steam cracking. Therefore, the technology of using heavy petroleum hydrocarbons to produce low-carbon olefins such as ethylene and propylene is receiving increasing attention.

[0004] CN1393510A discloses a method for catalytic conversion of heavy petroleum hydrocarbons to increase the production of ethylene and propylene. The method involves contacting and reacting the petroleum hydrocarbon feedstock with a catalyst containing a five-membered ring high-silica molecular sieve in a riser or fluidized bed reactor. Water vapor is injected in stages as the oil-agent mixture moves upward along the reactor. This method can also alleviate the hydrothermal deactivation of the catalyst to some extent.

[0005] CN102443423A discloses a method for catalytic conversion of petroleum hydrocarbons to produce high yields of ethylene, propylene, and light aromatics. Preheated petroleum hydrocarbon feedstock and a pyrolysis catalyst are introduced into the bottom of a fluidized bed reactor. The mixture of the petroleum hydrocarbon feedstock and the pyrolysis catalyst passes sequentially from bottom to top through a pre-lifting zone, a first reaction zone, a quenching zone, and a second reaction zone of the fluidized bed reactor under the action of a lifting medium. The reaction occurs in the first reaction zone, the quenching zone, and the second reaction zone under catalytic pyrolysis conditions. By controlling the reaction temperature in each reaction zone, and by controlling the reaction temperature in the first reaction zone to be 20-150°C higher than that in the second reaction zone, higher yields of ethylene, propylene, and light aromatics can be obtained.

[0006] CN109666506A discloses a method for catalytic cracking of hydrogenated oil. The method includes contacting non-hydrogenated oil with a catalytic cracking catalyst from the bottom of a riser reactor to carry out a first catalytic cracking reaction, and contacting hydrogenated oil with a first reaction product and a semi-raw catalyst to carry out a second catalytic cracking reaction. This method can regulate the hydrogen transfer activity in the reaction system and improve the yield of low-carbon olefins by feeding hydrogenated oil and non-hydrogenated oil in layers and adjusting the feed ratio of hydrogenated oil and non-hydrogenated oil.

[0007] However, in order to balance heavy oil conversion and increased ethylene and propylene production, existing technologies typically require relatively mild reaction conditions to suppress coking in the unit, resulting in lower ethylene and propylene yields; or heavy oil is hydrogenated before catalytic cracking conversion, but this involves high energy consumption and demanding conditions for catalytic cracking conversion. Summary of the Invention

[0008] The purpose of this invention is to increase the yield of ethylene and propylene in heavy oil catalytic conversion products while reducing the formation of coke.

[0009] To achieve the above objectives, the present invention provides a method for producing ethylene and propylene from heavy oil, the method comprising: Heavy hydrocarbon oil is introduced into a supercritical extraction separation unit for supercritical extraction separation to obtain light deasphalted oil, heavy deasphalted oil and deoiled asphalt, respectively. The deoiled bitumen is subjected to hydrogenation treatment to obtain hydrogenated deoiled bitumen; The heavy deasphalted oil and the hydrodeoiled asphalt are introduced into the first reactor and contacted with the first catalytic conversion catalyst to carry out the first catalytic conversion reaction, thereby obtaining the first oil-agent mixture. The light deasphalted oil is introduced into the second reactor and contacted with the second catalytic conversion catalyst to carry out the second catalytic conversion reaction, thereby obtaining the second oil-agent mixture. The first oil mixture and the second oil mixture are introduced into a dense-phase bed reactor for a third catalytic conversion reaction. The reaction streams are separated to obtain reaction oil and gas and a catalyst to be generated. The reaction oil and gas are further separated to obtain dry gas, liquefied petroleum gas, gasoline, diesel, and slurry oil. The first catalytic conversion catalyst and the second catalytic conversion catalyst contain MPZ hierarchical porous molecular sieves.

[0010] Optionally, the supercritical extraction separation includes the following steps: feeding heavy hydrocarbon oil into a first extraction tower for first separation to obtain a deasphalted oil solution and a deoiled asphalt solution; feeding the deasphalted oil solution into a second extraction tower for second separation to obtain a light deasphalted oil solution and a heavy deasphalted oil solution; removing the extraction solvent from the light deasphalted oil solution, the heavy deasphalted oil solution, and the deoiled asphalt solution respectively; wherein the extraction solvent is selected from one or more of C4 alkanes, C5 alkanes, and C6 alkanes; the operating conditions of the first extraction tower include: an extraction temperature 5-40 °C higher than the critical temperature of the extraction solvent, and an internal pressure of 6-20 MPa; the operating conditions of the second extraction tower include: an extraction temperature 5-40 °C higher than the critical temperature of the extraction solvent, and an internal pressure of 6-15 MPa. MPa; Optionally, the content of gum in the heavy hydrocarbon oil is 5.0-20.0% by weight, and the content of asphaltene is 0.5-5.0% by weight; The heavy hydrocarbon oil is selected from one or a mixture of several of vacuum gas oil, atmospheric residue, vacuum residue, hydrotreated vacuum gas oil, and hydrotreated atmospheric residue.

[0011] Optionally, the asphaltene content in the lightly deasphalted oil is 0.01-0.2% by weight, preferably 0.01-0.1% by weight; the carbon residue content is 0.1-2.0% by weight, preferably 0.1-2.0% by weight; optionally, the asphaltene content in the heavily deasphalted oil is 0.05-0.5% by weight, preferably 0.05-0.2% by weight; and / or the carbon residue content in the heavily deasphalted oil is 0.5-4.0% by weight, preferably 0.5-2.0% by weight.

[0012] Optionally, the conditions for the first catalytic conversion reaction include: a reaction temperature of 580-640 ℃, preferably 600-620 ℃; a catalyst-to-oil ratio of 5-25, preferably 10-20; an oil-gas residence time of 0.5-1.5 s, preferably 0.5-1.0 s; and a reaction pressure of 0.15-0.30 MPa, preferably 0.15-0.20 MPa. The conditions for the second catalytic conversion reaction include: a reaction temperature of 600-660 ℃, preferably 620-640 ℃; a catalyst-to-oil ratio of 5-50, preferably 10-40; an oil-gas residence time of 0.5-1.5 s, preferably 0.5-1.0 s; and a reaction pressure of 0.15-0.30 MPa, preferably 0.15-0.20 MPa. The conditions for the third catalytic conversion reaction include: a reaction temperature of 560-640 ℃, preferably 580-620 ℃. ℃; weight hourly space velocity is 1-30 h -1 Preferably 3-20 h -1 The reaction pressure is 0.15-0.30 MPa, preferably 0.15-0.20 MPa.

[0013] Optionally, the first catalytic conversion catalyst and the second catalytic conversion catalyst each independently comprise a molecular sieve mixture, a heat-resistant inorganic oxide, and clay; based on the weight of each of the first catalytic conversion catalyst and the second catalytic conversion catalyst, each of the first catalytic conversion catalyst and the second catalytic conversion catalyst independently contains 1-60% by weight of the molecular sieve mixture, 5-99% by weight of the heat-resistant inorganic oxide, and 0-70% by weight of clay, and the sum of the weight contents of each component is 100%; based on the total weight of the molecular sieve mixture, the molecular sieve mixture contains 80-100% by weight of the MPZ hierarchical porous molecular sieve.

[0014] Optionally, in the surface XPS elemental analysis of the MPZ hierarchical porous molecular sieve, n1 / n2 is less than or equal to 0.08, where n1 represents the number of moles of phosphorus in the MPZ hierarchical porous molecular sieve, and n2 represents the total number of moles of silicon and aluminum in the MPZ hierarchical porous molecular sieve; preferably, in the NH3-TPD spectrum after hydrothermal aging at 800 ℃, 100% water vapor conditions, for 17 h, the area of ​​strong acid central peaks with desorption temperatures above 200 ℃ accounts for more than or equal to 45% of the total acid central peak area.

[0015] Optionally, the method for hydrotreating the deasphalted oil includes the steps of: introducing the deasphalted oil into a fixed-bed reactor and contacting it with a hydrotreating catalyst; the hydrotreating conditions for the deasphalted oil include: a reaction temperature of 300-460℃, a reaction pressure of 6-25 MPa, and a liquid hourly space velocity of 0.10-1.0 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100-1500; optionally, based on the total weight of the hydrogenation catalyst, the hydrogenation catalyst comprises 0-30% by mass of a metal active component and a support, wherein the metal active component is selected from one or more of nickel, tungsten, cobalt and molybdenum, preferably selected from one or more of nickel-cobalt combination, nickel-tungsten-cobalt combination, nickel-molybdenum combination and cobalt-molybdenum combination.

[0016] Optionally, the method further includes: regenerating the spent catalyst by coking to obtain a regenerated catalyst; dividing the regenerated catalyst into at least two streams and returning them to the first reactor and the second reactor for recycling, respectively; optionally, the conditions for coking regeneration include: a temperature of 650-750 ℃ ​​and a pressure of 0.15-0.30 MPa.

[0017] Optionally, the method further includes: returning a portion of the slurry to the first reactor for catalytic conversion; the distance between the slurry's inlet location and the outlet location of the first reactor is 1 / 3 to 1 / 2 of the height of the first reactor.

[0018] A second aspect of the present invention provides a system for the catalytic conversion method described in the first aspect of the present invention, the system comprising a supercritical extraction unit, a hydrotreating unit, a catalytic cracking unit, and a product separation unit; the supercritical extraction and separation unit comprising a first extraction tower, a second extraction tower, a solvent recovery tower, a first stripping tower, a second stripping tower, and a third stripping tower; the hydrotreating unit comprising one or more reactors in series, the reactor preferably being a fixed-bed reactor; the catalytic cracking unit comprising at least a first reactor and a second reactor, the first reactor and the second reactor being selected from one or more combinations of bubbling bed reactors, dense-phase bed reactors, turbulent bed reactors, fast-bed reactors, conveying bed reactors, and riser reactors; preferably, the first reactor is a riser reactor, and the second reactor is a riser reactor; the first reactor and the second reactor are arranged in parallel and connected in series with the dense-phase bed reactor, respectively.

[0019] Through the above technical solution, the present invention has the following beneficial effects: (1) By supercritical extraction separation of heavy hydrocarbon oil, some of the gum and most of the asphaltenes in the heavy hydrocarbon oil can be removed, while most of the residual carbon and metals are removed, thus reducing coking. (2) The deoiled bitumen separated by the supercritical extraction separation unit is introduced into the hydrotreating unit for hydrotreating, which saturates the polycyclic aromatic hydrocarbons that are not easily reacted, improves the cracking performance of the deoiled bitumen, and can also reduce the hydrogen consumption and energy consumption of the hydrotreating unit. (3) Light deasphalted oil, heavy deasphalted oil and hydrodeasphalted oil are catalytically converted in different reactors. Creating a suitable reaction environment based on the properties of the catalytic conversion feedstock is beneficial to increasing the yield of ethylene and propylene and reducing the yield of coke.

[0020] Other features and advantages of the present invention will be described in detail in the following detailed description section. Attached Figure Description

[0021] The accompanying drawings are provided to further illustrate the invention and form part of the specification. They are used together with the following detailed description to explain the invention, but do not constitute a limitation thereof. In the drawings: Figure 1 This is a flowchart illustrating a specific embodiment of the catalytic conversion method of the present invention.

[0022] Figure 2 This is a flowchart illustrating a catalytic conversion method according to another specific embodiment of the present invention.

[0023] Figure 3 This is a schematic flowchart of a supercritical extraction separation method for heavy hydrocarbon oil according to a specific embodiment of the present invention.

[0024] Explanation of reference numerals in the attached figures: Detailed Implementation

[0025] The following provides a detailed description of specific embodiments of the present invention. It should be understood that the specific embodiments described herein are for illustrative and explanatory purposes only and are not intended to limit the scope of the invention.

[0026] In this invention, unless otherwise specified, pressure refers to absolute pressure.

[0027] A first aspect of the present invention provides a method for producing ethylene and propylene from heavy oil, the method comprising: Heavy hydrocarbon oil is introduced into a supercritical extraction separation unit for supercritical extraction separation to obtain light deasphalted oil, heavy deasphalted oil and deoiled asphalt, respectively. The deoiled bitumen is subjected to hydrogenation treatment to obtain hydrogenated deoiled bitumen; The heavy deasphalted oil and the hydrodeoiled asphalt are introduced into the first reactor and contacted with the first catalytic conversion catalyst to carry out the first catalytic conversion reaction, thereby obtaining the first oil-agent mixture. The light deasphalted oil is introduced into the second reactor and contacted with the second catalytic conversion catalyst to carry out the second catalytic conversion reaction, thereby obtaining the second oil-agent mixture. The first oil mixture and the second oil mixture are introduced into a dense-phase bed reactor for a third catalytic conversion reaction. The reaction streams are separated to obtain reaction oil and gas and a catalyst to be generated. The reaction oil and gas are further separated to obtain dry gas, liquefied petroleum gas, gasoline, diesel, and slurry oil. The first catalytic conversion catalyst and the second catalytic conversion catalyst contain MPZ hierarchical porous molecular sieves.

[0028] Through the above technical solution, this invention removes some of the gum and most of the asphaltenes from heavy hydrocarbon oil by supercritical extraction separation, while also removing most of the residual carbon and metals, thus reducing coking. The de-oiled bitumen separated by the supercritical extraction separation unit is introduced into the hydrotreating unit for hydrotreating, which saturates the non-reactive polycyclic aromatic hydrocarbons, improves the cracking performance of the de-oiled bitumen, and also reduces the hydrogen consumption and energy consumption of the hydrotreating unit.

[0029] Meanwhile, light deasphalted oil, heavy deasphalted oil, and hydrodeasphalted asphalt undergo catalytic conversion reactions in different reactors. Creating a suitable reaction environment based on the properties of the catalytic conversion feedstock is beneficial for increasing the yield of ethylene and propylene and reducing the yield of coke.

[0030] In this invention, the reaction oil and gas are introduced into the product separation unit for separation. The resulting dry gas is rich in ethylene, and the liquefied gas is rich in propylene. The coke yield is calculated by the concentration of carbon monoxide and carbon dioxide in the regenerated flue gas and the main air flow rate.

[0031] In some embodiments of the present invention, the supercritical extraction separation includes the following steps: The heavy hydrocarbon oil is fed into the first extraction tower for the first separation to obtain a deasphalted oil solution and a deoiled asphalt solution. The deasphalted oil solution is fed into a second extraction tower for a second separation to obtain a light deasphalted oil solution and a heavy deasphalted oil solution. The extraction solvent is removed from the light deasphalted oil solution, the heavy deasphalted oil solution, and the deoiled asphalt solution, respectively.

[0032] The extraction solvent is selected from one or more of C4 alkanes, C5 alkanes, and C6 alkanes.

[0033] In some embodiments of the present invention, the operating conditions of the first extraction tower include: the extraction temperature is 5-40 °C higher than the critical temperature of the extraction solvent, and the internal pressure is 6-20 MPa.

[0034] In some embodiments of the present invention, the operating conditions of the second extraction tower include: the extraction temperature is 5-40 °C higher than the critical temperature of the extraction solvent, and the internal pressure is 6-15 MPa.

[0035] In some embodiments of the present invention, the heavy hydrocarbon oil is separated by the above-mentioned supercritical extraction, and the yield of the light deasphalted oil 102 is 50-80%, the yield of the heavy deasphalted oil 103 is 10-40%, and the yield of the deoiled bitumen 104 is 10-50%. Compared with the heavy hydrocarbon oil, the obtained light deasphalted oil 102 has an asphalt removal rate of greater than 90% and a residual carbon removal rate of greater than 80%; the heavy deasphalted oil 103 has an asphalt removal rate of greater than 80% and a residual carbon removal rate of greater than 60%.

[0036] In some embodiments of the present invention, the content of gum in the heavy hydrocarbon oil is 5.0-20.0% by weight, and the content of asphaltene is 0.5-5.0% by weight; the heavy hydrocarbon oil is selected from one or a mixture of several of vacuum gas oil, atmospheric residue, vacuum residue, hydrotreated vacuum gas oil, and hydrotreated atmospheric residue.

[0037] In some embodiments of the present invention, the initial boiling point of the light deasphalted oil is any temperature between 280-320 °C; the final boiling point is any temperature between 640-680 °C.

[0038] In some embodiments of the present invention, the content of asphaltene in the light deasphalted oil is 0.01-0.2% by weight, preferably 0.01-0.1% by weight; and the content of carbon residue is 0.1-2.0% by weight, preferably 0.1-1.0% by weight.

[0039] In some embodiments of the present invention, the initial boiling point of the deasphalted oil is any temperature between 300-340°C, and the final boiling point is any temperature between 680-720°C.

[0040] In some embodiments of the present invention, the asphalt content in the deasphalted oil is 0.05-0.5% by weight, preferably 0.05-0.2% by weight.

[0041] In some embodiments of the present invention, the residual carbon content in the deasphalted oil is 0.5-4.0% by weight, preferably 0.5-2.0% by weight.

[0042] In some embodiments of the present invention, the method for hydrotreating the deasphalted oil includes the step of introducing the deasphalted oil into a fixed-bed reactor and contacting it with a hydrotreating catalyst.

[0043] In this invention, the deoiled bitumen is introduced into one or more reactors in series in the hydrotreating unit to react with the hydrotreating catalyst, and the reactor is preferably a fixed-bed reactor.

[0044] In some embodiments of the present invention, the hydrotreating conditions for the deasphalted oil include: a reaction temperature of 300-460°C, a reaction pressure of 6-25 MPa, and a liquid hourly space velocity of 0.10-1.0 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100-1500.

[0045] In some embodiments of the present invention, the hydrogenation catalyst comprises 0-30% by mass of a metal active component and a support, based on the total weight of the hydrogenation catalyst. The metal active component is selected from one or more of nickel, tungsten, cobalt and molybdenum, preferably from one or more of nickel-cobalt combination, nickel-tungsten-cobalt combination, nickel-molybdenum combination and cobalt-molybdenum combination.

[0046] In some embodiments of the present invention, the first catalytic conversion catalyst and the second catalytic conversion catalyst each independently comprise a molecular sieve mixture, a heat-resistant inorganic oxide, and clay; based on the weight of each of the first catalytic conversion catalyst and the second catalytic conversion catalyst, each of the first catalytic conversion catalyst and the second catalytic conversion catalyst independently contains 1-60% by weight of the molecular sieve mixture, 5-99% by weight of the heat-resistant inorganic oxide, and 0-70% by weight of clay, and the sum of the weight contents of each component is 100%; based on the total weight of the molecular sieve mixture, the molecular sieve mixture contains 80-100% by weight of the MPZ hierarchical porous molecular sieve.

[0047] In this invention, the MPZ hierarchical porous molecular sieve is a phosphorus- and metal-containing hierarchical porous ZSM-5 molecular sieve, in which phosphorus is fully coordinated with the framework aluminum, which can fully protect the framework aluminum.

[0048] In some embodiments of the present invention, in the surface XPS elemental analysis of the MPZ multi-level porous molecular sieve, n1 / n2 is less than or equal to 0.08, where n1 represents the number of moles of phosphorus in the MPZ multi-level porous molecular sieve, and n2 represents the total number of moles of silicon and aluminum in the MPZ multi-level porous molecular sieve.

[0049] In some embodiments of the present invention, the MPZ hierarchical porous molecular sieve exhibits good hydrothermal stability. Specifically, after hydrothermal aging at 800°C and 100% steam for 17 hours, the NH3-TPD spectrum of the MPZ hierarchical porous molecular sieve shows that the area of ​​strong acid centers with desorption temperatures above 200°C accounts for greater than or equal to 45% of the total area of ​​acid centers, preferably between 45% and 60%, indicating a high retention rate of strong acid centers. This molecular sieve can increase the production of liquefied petroleum gas, improve the yield of low-carbon olefins, and produce more high-value-added products. The preparation method of the MPZ hierarchical porous molecular sieve can be found in CN113526519A and CN14715911A.

[0050] The first reactor and the second reactor are selected from one or more combined reactor types, including bubbling bed reactor, dense phase bed reactor, turbulent bed reactor, fast bed reactor, conveying bed reactor and riser reactor; preferably, the first reactor and the second reactor are riser reactors.

[0051] In some embodiments of the present invention, the conditions for the first catalytic conversion reaction include: a reaction temperature of 580-640℃, preferably 600-620℃; an oil-to-catalyst ratio of 5-25, preferably 10-20; an oil-gas residence time of 0.5-1.5 s, preferably 0.5-1.0 s; and a reaction pressure of 0.15-0.30 MPa, preferably 0.15-0.20 MPa.

[0052] In some embodiments of the present invention, the conditions for the second catalytic conversion reaction include: a reaction temperature of 600-660℃, preferably 620-640℃; an oil-to-catalyst ratio of 5-50, preferably 10-40; an oil-gas residence time of 0.5-1.5 s, preferably 0.5-1.0 s; and a reaction pressure of 0.15-0.30 MPa, preferably 0.15-0.20 MPa.

[0053] In some embodiments of the present invention, the conditions for the third catalytic conversion reaction include: a reaction temperature of 560-640°C, preferably 580-620°C; and a weight hourly space velocity of 1-30 h⁻¹. -1 Preferably 3-20 h -1 The reaction pressure is 0.15-0.30 MPa, preferably 0.15-0.20 MPa.

[0054] In some embodiments of the present invention, the method further includes: returning a portion of the slurry to the first reactor for catalytic conversion reaction; the distance between the slurry's inlet position and the outlet position of the first reactor is 1 / 3 to 1 / 2 of the height of the first reactor.

[0055] In some embodiments of the present invention, the method further includes: regenerating the spent catalyst by coking to obtain a regenerated catalyst; dividing the regenerated catalyst into at least two streams, which are then returned to the first reactor and the second reactor for recycling, respectively; optionally, the conditions for coking regeneration include: a temperature of 650-750 °C and a pressure of 0.15-0.30 MPa.

[0056] The present invention also provides a system for the above-described catalytic conversion method, the system comprising a supercritical extraction unit, a hydrogenation unit, a catalytic cracking unit, and a product separation unit.

[0057] The supercritical extraction separation unit includes a first extraction tower, a second extraction tower, a solvent recovery tower, a first stripping tower, a second stripping tower, and a third stripping tower. Heavy hydrocarbon oil is introduced into the first extraction tower to obtain a deasphalted oil solution and a deoiled asphalt solution. The deasphalted oil solution is introduced into the second extraction tower to obtain a light deasphalted oil mixture and a heavy deasphalted oil solution. The light deasphalted oil mixture is introduced into the solvent recovery tower to remove entrained extraction solvent, yielding a light deasphalted oil solution. The light deasphalted oil solution, heavy deasphalted oil solution, and deoiled asphalt solution are respectively introduced into the first stripping tower, the second stripping tower, and the third stripping tower to remove entrained extraction solvent, yielding light deasphalted oil, heavy deasphalted oil, and deoiled asphalt, respectively. The extraction solvent can be recycled.

[0058] The hydrogenation treatment unit includes one or more reactors in series, preferably a fixed-bed reactor.

[0059] The catalytic cracking unit includes at least a first reactor and a second reactor. The first reactor and the second reactor are selected from one or more combinations of bubbling bed reactor, dense phase bed reactor, turbulent bed reactor, fast bed reactor, conveying bed reactor, and riser reactor. The first reactor is a riser reactor and the second reactor is a riser reactor. The first reactor and the second reactor are connected in series. The first reactor and the second reactor are arranged in parallel and are connected in series with the dense phase bed reactor respectively.

[0060] Figure 1 This illustration schematically shows a particularly preferred embodiment of the invention, such as... Figure 1 As shown, heavy hydrocarbon oil 101 is introduced into supercritical extraction separation unit 1 for supercritical extraction separation to obtain light deasphalted oil 102, heavy deasphalted oil 103, and deoiled asphalt 104. Deoiled asphalt 104 is introduced into hydrotreating unit 2 to react with a hydrotreating catalyst to obtain hydrotreated deoiled asphalt 201.

[0061] Hydrogenated deasphalted bitumen 201 and heavy deasphalted oil 103 are mixed and preheated to 250-350°C. The mixture is then sprayed into the lower part of the first reactor 3-1 through a nozzle. It comes into contact with the first regeneration catalyst 301 from the regenerator 3-6, which is introduced into the bottom of the first reactor 3-1 through a pipeline, to carry out the first catalytic conversion reaction and obtain the first oil-agent mixture.

[0062] After the light deasphalted oil 102 is preheated to 200-300℃, it is sprayed into the lower part of the second reactor 3-2 through a nozzle, and comes into contact with the second regeneration catalyst 302 from the regenerator 3-6 introduced into the bottom of the second reactor 3-2 through a pipeline to carry out the second catalytic conversion reaction, and the second oil-agent mixture is obtained.

[0063] The oil-solid mixture produced by the first reactor 3-1 and the second reactor 3-2 is introduced into the dense phase bed reactor 3-3 for the third catalytic conversion reaction. The reacted oil-solid mixture is then sent to a separation unit for gas-solid separation to obtain a spent catalyst and reacted oil-gas. The spent catalyst obtained after separation is introduced into the stripper 3-4 for stripping. The stripped spent catalyst 303 is then introduced into the regenerator 3-6 through a spent catalyst inclined tube for regeneration to obtain a regenerated catalyst. The regenerated catalyst is divided into at least two streams and introduced into the first reactor 3-1 and the second reactor 3-2 respectively for recycling. The oil-gas mixture obtained after separation by the separation unit is introduced into the product separation unit 4 for further separation to obtain products such as dry gas 401, liquefied petroleum gas 402, gasoline 403, diesel 404, and oil slurry 405.

[0064] In one specific embodiment, the method provided by the present invention can also be achieved through, for example... Figure 2 The system described is implemented in conjunction with Figure 1 Compared to the method of the illustrated embodiment, this embodiment further includes introducing a portion of the oil slurry 405 obtained from the product separation unit 4 into the first reactor 3-1 of the catalytic cracking unit 3, where it comes into contact with the existing stream in the first reactor 3-1, which is introduced through a pipeline to the bottom of the first reactor 3-1. The reaction conditions may include: a reaction temperature of 580-640°C, preferably 600-620°C; an oil-to-catalyst ratio of 5-25, preferably 10-20; an oil-gas residence time of 0.5-1.5 s, preferably 0.5-1.0 s; and a reaction pressure of 0.15-0.30 MPa, preferably 0.15-0.20 MPa.

[0065] Figure 3 This illustration schematically depicts a preferred embodiment of the supercritical extraction and separation of heavy hydrocarbon oils according to the present invention, with reference to... Figure 3 The supercritical extraction separation unit 1 includes a first extraction tower 1-1, a second extraction tower 1-2, a solvent recovery tower 1-3, a first stripping tower 1-4, a second stripping tower 1-5, and a third stripping tower 1-6.

[0066] Heavy hydrocarbon oil 101 is introduced into the first extraction tower 1-1 and subjected to first separation under supercritical extraction conditions, where it is miscible with the extraction solvent, yielding deasphalted oil solution 105 and deoiled asphalt solution 106. Deasphalted oil solution 105 is then introduced into the second extraction tower 1-2 and subjected to second separation under supercritical extraction conditions, where it is miscible with the extraction solvent, yielding a light deasphalted oil mixture 107 and a heavy deasphalted oil solution 108. The light deasphalted oil mixture 107 is then introduced into the solvent recovery tower 1-3 to remove entrained extraction solvent, yielding a light deasphalted oil solution 109.

[0067] The light deasphalted oil solution 109, heavy deasphalted oil solution 108, and deoiled asphalt solution 106 are introduced into the first stripping tower 1-4, the second stripping tower 1-5, and the third stripping tower 1-6, respectively, to remove the entrained extraction solvent, thereby obtaining light deasphalted oil 102, heavy deasphalted oil 103, and deoiled asphalt 104, respectively. The extraction solvent 110 recovered in the first stripping tower 1-4, the second stripping tower 1-5, and the third stripping tower 1-6 can be recycled.

[0068] The present invention will be further described in detail below through embodiments, but the invention is not limited thereto. All raw materials used in the embodiments are commercially available.

[0069] In the embodiments and comparative examples of this invention, the gaseous products were tested using the RIPP 77-90 method for petrochemical analysis, the coke content was determined using the RIPP 107-90 method for petrochemical analysis, and the composition of the organic liquid products was determined using the SH / T0558-1993 method.

[0070] In the following embodiments, the product yield is calculated according to the following formula:

[0071] The RIPP petrochemical analysis method used in this invention is selected from "Petrochemical Analysis Methods (RIPP Test Methods)", edited by Yang Cuiding et al., Science Press, 1990.

[0072] In the method of this embodiment, the first reactor and the second reactor are riser reactors, and the first reactor and the second reactor are arranged in parallel. The outlet of the first reactor is connected to the dense phase bed reactor, and the outlet of the second reactor extends to the settler and is connected to the dense phase bed reactor.

[0073] The hydrogenation catalysts used in the examples and comparative examples include hydrogenation protection catalyst A, hydrogenation demetallization catalyst B, and hydrogenation desulfurization and decarbonization catalyst C. Based on the total weight of the hydrogenation catalysts, the content of hydrogenation protection catalyst A is 5% by weight, the content of hydrogenation demetallization catalyst B is 35% by weight, and the content of hydrogenation desulfurization and decarbonization catalyst C is 60% by weight. The specific composition and properties of the hydrogenation catalysts are shown in Table 1.

[0074] The catalytic conversion catalysts used in the examples and comparative examples were MPZ molecular sieve catalyst and Eplene-10 catalyst, both produced by Sinopec Catalyst Branch. Their composition and properties are shown in Table 2. Based on the total amount of MPZ molecular sieve catalyst, the MPZ catalyst contained 35 wt% MPZ molecular sieve (silicon-to-alumina ratio of 22), 40 wt% kaolin, and 25 wt% alumina binder. The Eplene-10 catalyst contained 35 wt% ZSM-5 molecular sieve (silicon-to-alumina ratio of 22). Both the MPZ molecular sieve catalyst and the Eplene-10 catalyst were aged for 17 h at 800℃ in a 100% steam atmosphere before use. After hydrothermal aging, in the NH3-TPD spectra, the proportions of the strong acid center peak area with a desorption temperature above 200℃ to the total acid center peak area were 55.7% and 42.9%, respectively.

[0075] The raw materials used in the examples and comparative examples were heavy hydrocarbon oils, the properties of which are shown in Table 3.

[0076] The extraction solvent used in the supercritical extraction units of the examples and comparative examples was a mixture of butanes, the composition of which is shown in Table 4.

[0077] Table 1. Composition and properties of hydrogenation catalysts

[0078] Table 2 Composition and properties of catalytic conversion catalysts

[0079] Table 3 Composition and properties of heavy hydrocarbon oils

[0080] Table 4 Composition of Extraction Solvents

[0081] Example 1 The experiment was conducted in Figure 1 The medium-sized device and in the flowchart shown Figure 3 The process was carried out on a medium-sized unit of the supercritical separation unit shown.

[0082] Heavy hydrocarbon oil (properties and composition shown in Table 3) is introduced from the middle of the first extraction tower, while the extraction solvent is fed from the top of the first extraction tower. Under supercritical extraction conditions, it undergoes a first separation with mixed butane (properties and composition shown in Table 4) to obtain a deasphalted oil solution and a deoiled asphalt solution. The deasphalted oil solution is then introduced into a second extraction tower for a second separation under supercritical extraction conditions with mixed butane to obtain a light deasphalted oil mixture and a heavy deasphalted oil solution. The light deasphalted oil mixture is then introduced into a solvent recovery tower to remove entrained extraction solvent, yielding a light deasphalted oil solution.

[0083] The total extractant mixture of butane and heavy hydrocarbon oil entering the first extraction tower had a weight ratio of 0.5:1. The operating conditions for the first extraction tower were: temperature 162.5 ℃ and pressure 15.3 MPa; the operating conditions for the second extraction tower were: temperature 157.9 ℃ and pressure 10.8 MPa. The properties of the obtained light deasphalted oil and heavy deasphalted oil are shown in Table 5.

[0084] Deoiled bitumen was introduced into a fixed-bed reactor and reacted with a hydrogenation catalyst packed in the reactor. The properties of the resulting hydrogenated deoiled bitumen are shown in Table 5. The hydrogenation reaction conditions included a reaction temperature of 421.4 °C, a reaction pressure of 18.5 MPa, and a volume hourly space velocity of 0.5 h⁻¹. -1 The hydrogen-to-oil volume ratio is 200.

[0085] Hydrogenated deasphalted asphalt and heavy deasphalted oil are mixed and sprayed into the lower part of the first reactor through a nozzle. They come into contact with the first catalytic conversion catalyst from the regenerator to undergo the first catalytic conversion reaction, yielding a first oil-fuel mixture. The heavy deasphalted oil is then introduced into the second reactor to come into contact with the second catalytic conversion catalyst to undergo the second catalytic conversion reaction, yielding a second oil-fuel mixture. The first and second oil-fuel mixtures are then introduced into a dense-phase bed reactor for a third catalytic conversion reaction. The resulting oil-fuel mixture is separated by a separation device. The resulting spent catalyst is introduced into a stripper for stripping. The stripped spent catalyst is then introduced into the regenerator through a spent catalyst delivery pipe for regeneration. The resulting regenerated catalyst is divided into at least two streams and introduced into the first and second reactors respectively for recycling. The resulting oil-gas mixture is introduced into a product separation unit to obtain dry gas, liquefied petroleum gas (LPG), gasoline, diesel, and slurry oil. The coke yield of the catalytic cracking unit is calculated based on the main air flow rate in the regenerator and the concentrations of CO and CO2 in the regenerated flue gas. The reaction conditions and results are shown in Table 6.

[0086] Table 5. Composition and properties of lightly deasphalted oil, heavily deasphalted oil, and hydrodeasphalted oil.

[0087] Example 2 Example 2, the method for producing ethylene and propylene from heavy oil, is basically similar to the catalytic conversion method in Example 1, except that: The reaction temperature in the first reactor was 605 ℃, and the agent-to-oil ratio was 15.1; the reaction temperature in the second reactor was 630 ℃, and the agent-to-oil ratio was 16.5. The reaction conditions and results are shown in Table 6.

[0088] Example 3 Example 3, the method for producing ethylene and propylene from heavy oil, is basically similar to the catalytic conversion method in Example 1, except that: The reaction temperature in the first reactor was 590 ℃, and the agent-to-oil ratio was 14.9; the reaction temperature in the second reactor was 650 ℃. The reaction conditions and results are shown in Table 6.

[0089] Comparative Example 1 The method in Comparative Example 1 is basically similar to that in Example 1, except that: heavy hydrocarbon oil (properties shown in Table 3) is introduced into the first reactor to react with the catalytic conversion catalyst Eplene-10 (properties shown in Table 2). The oil-catalyst mixture obtained from the first reactor is introduced into a dense-phase bed reactor for further reaction. The reaction stream is then introduced into a cyclone separator for separation. The separated spent catalyst is stripped and then introduced into a regenerator for regeneration. The regenerated catalyst is returned to the first reactor for recycling. The reaction conditions and results are shown in Table 6.

[0090] Comparative Example 2 The method in Comparative Example 2 is basically similar to that in Example 1, except that: heavy hydrocarbon oil is introduced into a fixed-bed reactor to contact with a hydrogenation catalyst for hydrogenation treatment. The hydrogenation reaction conditions include: a reaction temperature of 420.1°C, a reaction pressure of 18.5 MPa, and a volume hourly space velocity of 0.5 h⁻¹. -1 The hydrogen-to-oil volume ratio was 350. The hydrotreating product was introduced into the first reactor and reacted with the catalytic conversion catalyst Eplene-10. The oil-to-catalyst mixture obtained from the first reactor was introduced into a dense-phase bed reactor for further reaction. The reaction stream was separated by a cyclone separator, and the separated spent catalyst was stripped and then introduced into a regenerator for regeneration. The regenerated catalyst was returned to the first reactor for recycling. The reaction conditions and results are shown in Table 6.

[0091] Table 6 Reaction conditions and results of Examples 1-3 and Comparative Examples 1-2

[0092] As can be seen from the data in Table 6, compared with Comparative Example 1 and Comparative Example 2, the method provided by the present invention can reduce the hydrogen-to-oil ratio of the hydrotreating unit, thereby reducing hydrogen consumption, while increasing the yield of ethylene and propylene and reducing the yield of coke.

[0093] Example 4 The method in Example 4 is basically similar to the catalytic conversion method in Example 1, except that: exist Figure 2 The medium-sized unit shown in the flowchart was tested, which also included introducing the oil slurry obtained from the reaction oil-gas separation into the first reactor for reprocessing. The distance between the oil slurry introduction point and the outlet of the riser reactor was 1 / 3 of the reactor height, and the oil slurry accounted for 2.5% by weight of the total raw material in the first reactor. The reaction conditions and results are shown in Table 7.

[0094] Comparative Example 3 The method in Comparative Example 3 is similar to that in Comparative Example 2, except that the reaction temperature for the hydrogenation of heavy hydrocarbon oil is 419.5 °C, and the catalyst fed into the first reactor is an MPZ molecular sieve catalyst. exist Figure 2 The medium-sized unit shown in the flowchart was tested, which also included introducing the oil slurry obtained from the reaction oil-gas separation into the first reactor for reprocessing. The distance between the oil slurry introduction point and the outlet of the riser reactor was 1 / 3 of the reactor height, and the oil slurry accounted for 2.5% by weight of the total raw material in the first reactor. The reaction conditions and results are shown in Table 7.

[0095] Table 7 Reaction conditions and results of Example 4 and Comparative Example 3

[0096] As shown in Table 7, by refining the oil slurry, the yield of ethylene and propylene can be further increased, thus improving the utilization rate of raw materials.

[0097] The preferred embodiments of the present invention have been described in detail above. However, the present invention is not limited to the specific details in the above embodiments. Within the scope of the technical concept of the present invention, various simple modifications can be made to the technical solution of the present invention, and these simple modifications all fall within the protection scope of the present invention.

[0098] It should also be noted that the various specific technical features described in the above embodiments can be combined in any suitable manner without contradiction. To avoid unnecessary repetition, the present invention will not describe the various possible combinations separately.

[0099] Furthermore, various different embodiments of the present invention can be combined in any way, as long as they do not violate the spirit of the present invention, they should also be regarded as the content disclosed by the present invention.

Claims

1. A method for producing ethylene and propylene from heavy oil, characterized in that, The method includes: Heavy hydrocarbon oil is introduced into a supercritical extraction separation unit for supercritical extraction separation to obtain light deasphalted oil, heavy deasphalted oil and deoiled asphalt, respectively. The deoiled bitumen is subjected to hydrogenation treatment to obtain hydrogenated deoiled bitumen; The heavy deasphalted oil and the hydrodeoiled asphalt are introduced into the first reactor and contacted with the first catalytic conversion catalyst to carry out the first catalytic conversion reaction, thereby obtaining the first oil-agent mixture. The light deasphalted oil is introduced into the second reactor and contacted with the second catalytic conversion catalyst to carry out the second catalytic conversion reaction, thereby obtaining the second oil-agent mixture. The first oil mixture and the second oil mixture are introduced into a dense-phase bed reactor for a third catalytic conversion reaction. The reaction streams are separated to obtain reaction oil and gas and a catalyst to be generated. The reaction oil and gas are further separated to obtain dry gas, liquefied petroleum gas, gasoline, diesel, and slurry oil. The first catalytic conversion catalyst and the second catalytic conversion catalyst contain MPZ hierarchical porous molecular sieves.

2. The method according to claim 1, wherein, The supercritical extraction separation includes the following steps: The heavy hydrocarbon oil is fed into the first extraction tower for the first separation to obtain a deasphalted oil solution and a deoiled asphalt solution. The deasphalted oil solution is fed into a second extraction tower for a second separation to obtain a light deasphalted oil solution and a heavy deasphalted oil solution. The extraction solvent is removed from the light deasphalted oil solution, the heavy deasphalted oil solution, and the deoiled asphalt solution, respectively. The extraction solvent is selected from one or more of C4 alkanes, C5 alkanes, and C6 alkanes; The operating conditions of the first extraction tower include: the extraction temperature is 5-40 °C higher than the critical temperature of the extraction solvent, and the internal pressure is 6-20 MPa; The operating conditions of the second extraction tower include: the extraction temperature is 5-40 °C higher than the critical temperature of the extraction solvent, and the internal pressure is 6-15 MPa; Optionally, the content of gum in the heavy hydrocarbon oil is 5.0-20.0% by weight, and the content of asphaltene is 0.5-5.0% by weight; the heavy hydrocarbon oil is selected from one or a mixture of several of vacuum gas oil, atmospheric residue, vacuum residue, hydrotreated vacuum gas oil, and hydrotreated atmospheric residue.

3. The method according to claim 1, wherein, The asphalt content in the lightly deasphalted oil is 0.01-0.2% by weight, preferably 0.01-0.1% by weight; the carbon residue content is 0.1-2.0% by weight, preferably 0.1-1.0% by weight. Optionally, the asphalt content in the re-deasphalted oil is 0.05-0.5% by weight, preferably 0.05-0.2% by weight; and / or The residual carbon content in the deasphalted oil is 0.5-4.0% by weight, preferably 0.5-2.0% by weight.

4. The method according to claim 1, wherein, The conditions for the first catalytic conversion reaction include: a reaction temperature of 580-640 ℃, preferably 600-620 ℃; a catalyst-to-oil ratio of 5-25, preferably 10-20; an oil-gas residence time of 0.5-1.5 s, preferably 0.5-1.0 s; and a reaction pressure of 0.15-0.30 MPa, preferably 0.15-0.20 MPa. The conditions for the second catalytic conversion reaction include: a reaction temperature of 600-660 ℃, preferably 620-640 ℃; a catalyst-to-oil ratio of 5-50, preferably 10-40; an oil-gas residence time of 0.5-1.5 s, preferably 0.5-1.0 s; and a reaction pressure of 0.15-0.30 MPa, preferably 0.15-0.20 MPa. The conditions for the third catalytic conversion reaction include: a reaction temperature of 560-640 °C, preferably 580-620 °C; and a weight hourly space velocity of 1-30 h⁻¹. -1 Preferably 3-20 h -1 The reaction pressure is 0.15-0.30 MPa, preferably 0.15-0.20 MPa.

5. The method according to claim 1, wherein, The first catalytic conversion catalyst and the second catalytic conversion catalyst each independently comprise a molecular sieve mixture, a heat-resistant inorganic oxide, and clay; Based on the weight of the first catalytic conversion catalyst or the second catalytic conversion catalyst, each of the first catalytic conversion catalyst or the second catalytic conversion catalyst independently contains 1-60% by weight of molecular sieve mixture, 5-99% by weight of heat-resistant inorganic oxide and 0-70% by weight of clay, and the sum of the weight contents of each component is 100%. Based on the total weight of the molecular sieve mixture, the molecular sieve mixture contains 80-100% by weight of the MPZ multi-level porous molecular sieve.

6. The method according to claim 5, wherein, In the surface XPS elemental analysis of the MPZ multi-level porous molecular sieve, n1 / n2 is less than or equal to 0.08, where n1 represents the number of moles of phosphorus in the MPZ multi-level porous molecular sieve, and n2 represents the total number of moles of silicon and aluminum in the MPZ multi-level porous molecular sieve. Preferably, after hydrothermal aging at 800 °C and 100% steam for 17 h, the area of ​​strong acid center peaks with desorption temperatures above 200 °C in the NH3-TPD spectrum accounts for more than or equal to 45% of the total area of ​​acid center peaks.

7. The method according to claim 1, wherein, The method for hydrotreating the deasphalted oil includes the steps of: introducing the deasphalted oil into a fixed-bed reactor and contacting it with a hydrotreating catalyst; The hydrotreating conditions for the deasphalted oil include: a reaction temperature of 300-460 °C, a reaction pressure of 6-25 MPa, and a liquid hourly space velocity of 0.10-1.0 h⁻¹. -1 The hydrogen-to-oil volume ratio is 100-1500; Optionally, based on the total weight of the hydrogenation catalyst, the hydrogenation catalyst comprises 0-30% by mass of a metal active component and a support, wherein the metal active component is selected from one or more of nickel, tungsten, cobalt and molybdenum, preferably from one or more of nickel-cobalt combination, nickel-tungsten-cobalt combination, nickel-molybdenum combination and cobalt-molybdenum combination.

8. The method according to claim 1, wherein, The method further includes: regenerating the spent catalyst by coking to obtain a regenerated catalyst; The regenerated catalyst is divided into at least two streams and returned to the first reactor and the second reactor for recycling, respectively. Optionally, the conditions for coke regeneration include: a temperature of 650-750 ℃ ​​and a pressure of 0.15-0.30 MPa.

9. The method according to claim 1, wherein, The method further includes: returning a portion of the oil slurry to the first reactor for catalytic conversion reaction; The distance between the inlet of the slurry and the outlet of the first reactor is 1 / 3 to 1 / 2 of the height of the first reactor.

10. A system for use in the method according to any one of claims 1-9, characterized in that, The system includes a supercritical extraction unit, a hydrogenation unit, a catalytic cracking unit, and a product separation unit; The supercritical extraction and separation unit includes a first extraction tower, a second extraction tower, a solvent recovery tower, a first stripping tower, a second stripping tower, and a third stripping tower; the hydrogenation treatment unit includes one reactor or multiple reactors in series, preferably a fixed-bed reactor. The catalytic cracking unit includes at least a first reactor and a second reactor, wherein the first reactor and the second reactor are selected from one or more combinations of bubbling bed reactor, dense phase bed reactor, turbulent bed reactor, fast bed reactor, conveying bed reactor, and riser reactor; Preferably, the first reactor is a riser reactor and the second reactor is a riser reactor; the first reactor and the second reactor are arranged in parallel and connected in series with the dense phase bed reactor respectively.

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