A process for the preparation of a high octane gasoline component

Abstract

The present application relates to the technical field of petroleum chemical industry, in particular to a method for preparing high-octane gasoline component. The method comprises the following steps: (1) introducing organic oxygen-containing compound and naphtha into a moving bed reaction zone filled with aromatization catalyst to perform contact reaction, obtaining reaction product I, and separating the reaction product I to obtain a gas phase stream, an oil phase stream and an aqueous phase stream; (2) introducing spent catalyst flowing out from the bottom of the moving bed reaction zone into a regenerator to perform regeneration treatment, obtaining regenerated catalyst, and recycling the regenerated catalyst into the moving bed reaction zone to participate in the contact reaction. The method provided by the present application can improve the yield of high-octane gasoline, and meanwhile, high-quality liquefied gas is obtained.

Description

TECHNICAL FIELD

[0001] The present application relates to the technical field of petroleum chemical industry, in particular to a method for preparing high-octane gasoline component. BACKGROUND

[0002] With the rapid development of China's economy, the demand for oil resources is increasing. At present, more than 70% of the oil processed in China is imported. In recent years, the coal-based / bio-based oxygen-containing compound industry has developed rapidly. For example, as the most important carbon-based chemical basic product, the methanol industry has developed rapidly. It is particularly important and urgent to develop new technologies for utilizing coal-based / bio-based oxygen-containing compounds represented by methanol to expand their utilization and promote the sustainable development of coal chemical industry represented by methanol and bio-chemical industry in China.

[0003] Therefore, it is of great significance to develop a technology for producing high-octane gasoline from coal-based, bio-based and other oxygen-containing compounds to alleviate the pressure of China's oil energy and improve the energy structure.

[0004] CN104910957A discloses a process for preparing high-octane high-purity gasoline from naphtha and methanol. Naphtha enters a first reactor for reaction, and the reaction product is mixed with methanol in a mixer to reach 300-400 DEG C and enters a second reactor. The reaction product of the second reactor is cooled by a first condenser and separated into gasoline A and gas by a first gas-liquid separator. The gas is heated to 350-450 DEG C by a heater after heat exchange with the reaction product of a third reactor, and enters the third reactor. The reaction product of the third reactor is cooled by a second condenser and enters a second gas-liquid separator to obtain gasoline B and gas. More than 90% of the carbon and hydrogen in methanol is converted into high-octane high-purity gasoline, and 2.46-2.53 tons of methanol are converted into one ton of high-octane high-purity gasoline. This method utilizes the alkylation and isomerization reactions of naphtha and methanol under the action of a catalyst to increase the octane number of condensate or naphtha from 93 to 95.

[0005] CN104419441A discloses a method for producing high-octane gasoline blending component by alkylation of light aromatic hydrocarbons and alcohol / ether oxygen-containing compounds. The raw material light aromatic hydrocarbons enter the reactor from the top, and the alcohol / ether oxygen-containing compounds are divided into multiple paths and enter the reaction bed layers of the multi-stage cold shock fixed bed reactor. After mixing with the light aromatic hydrocarbons entering from the top, the reaction occurs on the special catalyst to generate high-octane alkyl aromatic hydrocarbon gasoline blending component. This method mainly utilizes the alkylation reaction of alcohol / ether oxygen-containing compounds and aromatic hydrocarbons to increase the gasoline yield.

[0006] CN101314731A discloses a light hydrocarbon non-hydrogen aromatization method, comprising the following steps: (1) mixing naphtha and C3-C5 light hydrocarbons and then entering a moving bed reaction zone to contact an aromatization catalyst for aromatization reaction; the spent catalyst flowing out from the bottom of the moving bed reaction zone enters a catalyst regenerator from the top through a spent catalyst elevator, a spent catalyst lifting pipeline, a spent catalyst separation hopper and a pressure change and flow control zone, and then passes through a buffer zone, a coking zone, a drying zone and a cooling zone in the catalyst regenerator from top to bottom, and the spent catalyst is regenerated by feeding a regeneration gas containing oxygen into the coking zone, and then the regenerated catalyst is lifted and enters the moving bed reactor from the top.

[0007] CN109569703A discloses a catalyst for producing a gasoline component from naphtha and methanol aromatization, which comprises a carrier and active components in the following contents calculated based on the carrier: Ag 0.1-5.0 mass%, VA group element oxide 1.0-15.0 mass%, rare earth element oxide 0.1-3.0 mass%. The carrier comprises 40-80 mass% ZSM-5 zeolite, 3-30 mass% aluminum oxide and 3-30 mass% amorphous aluminum silicate. The catalyst improves the gasoline yield in the reaction product and the aromatic hydrocarbon yield in the gasoline through the synergistic effect of Ag element, VA group element and rare earth element, and also improves the single-pass service life of the catalyst and has good performance after regeneration.

[0008] CN107573966A discloses a method for producing a high-octane gasoline component from Fischer-Tropsch synthesis naphtha, which comprises the following steps: the Fischer-Tropsch synthesis naphtha is contacted with a catalyst under the conditions of 0.1-2.0 MPa, 300-500°C and a feed mass space velocity of 0.1-10 h-1, and then the C2 or less gas and liquefied gas are separated, and the obtained liquid phase product is a high-octane gasoline component. The method can convert the Fischer-Tropsch synthesis naphtha into high-octane gasoline and high-quality liquefied gas.

[0009] CN107964416A discloses a method and system for realizing gasoline production by coupling methanol and light naphtha, which comprises the following steps: the raw materials methanol and light naphtha are processed through the steps of pressure boosting, preheating, vaporization, superheating, catalysis, cooling and separation, and finally a gasoline product is obtained. Meanwhile, the byproduct cycle gas produced in the production process is recycled and reused, and a heat exchange loop is designed in the system to improve the utilization rate of materials and energy in the system.

[0010] CN102746877A discloses a method for preparing gasoline from methanol and light petroleum fraction, mainly solves the problems of the prior art that the light petroleum fraction cannot or is difficult to be made into gasoline, and the incorporation amount of methanol is low in the process of preparing gasoline, and the equipment is easy to be corroded and carbon deposition under the condition of not modifying the engine. The method comprises the following steps: contacting the raw material with a solid acid catalyst under the conditions of a reaction temperature of 300-500 DEG C, a reaction pressure of normal pressure-10.0 MPa, and a weight space velocity of the raw material of 0.3-10 h-1, so that the methanol and the light petroleum fraction are subjected to alkylation reaction on the catalyst to obtain a gasoline product containing less than 100 ppm of oxygen-containing compounds by weight, thereby solving the problems and being applicable to the industrial production of gasoline.

[0011] CN1651141A discloses an aromatization catalyst and its preparation and application. HZSM-5 molecular sieve and REY molecular sieve modified by Zn, phosphorus and rare earth elements are used as the active components of the catalyst. The catalyst is used for the moving bed aromatization reaction of inferior gasoline with low octane number under the non-hydrogen condition. The reaction temperature is 350-550 DEG C, the pressure is 0.1-0.5 MPa, the space velocity is 0.5-5 h. The regeneration temperature of the catalyst is 400-550 DEG C, the residence time in the regenerator is 1-600 minutes, and air is used for regeneration. SUMMARY

[0012] The purpose of the present application is to provide a method for preparing high-octane gasoline components, which can improve the yield of high-octane gasoline components while obtaining high-quality liquefied gas.

[0013] In order to achieve the above-mentioned purpose, the present application provides a method for preparing high-octane gasoline components, which comprises the following steps:

[0014] (1) introducing an organic oxygen-containing compound and naphtha into a moving bed reaction zone filled with an aromatization catalyst to perform a contact reaction, to obtain a reaction product I, and separating the reaction product I to obtain a gas phase stream, an oil phase stream and an aqueous phase stream;

[0015] (2) introducing spent catalyst flowing out from the bottom of the moving bed reaction zone into a regenerator to perform a regeneration treatment, to obtain regenerated catalyst, and recycling the regenerated catalyst into the moving bed reaction zone to participate in the contact reaction;

[0016] The organic oxygen-containing compound is selected from at least one of alcohols, ethers and esters; and the mass ratio of the use amount of the organic oxygen-containing compound to the naphtha is 1:0.1-10.

[0017] The aromatization catalyst comprises a carrier and an active component, and the active component contains 0.5-15% of VA group element oxide and 0.1-2% of rare earth element oxide by mass based on the total mass of the carrier.

[0018] The method provided by the application can convert the organic oxygen-containing compound and the naphtha into high-octane gasoline under a non-hydrogen atmosphere, using a specific type of aromatization catalyst, thereby effectively solving the utilization problem of the light naphtha and the organic oxygen-containing compound.

[0019] In particular, the method provided by the application can improve the coking capacity of the regenerator, realize real-time regeneration of the catalyst, maintain the activity of the catalyst, improve the yield of high-octane gasoline, and obtain high-quality liquefied gas. BRIEF DESCRIPTION OF DRAWINGS

[0020] Figure 1 is a process flow chart of a particularly preferred method for preparing high-octane gasoline components provided by the application;

[0021] Figure 2 is a separation and fractionation process flow chart of a particularly preferred reaction product I provided by the application;

[0022] Figure 3 is a process flow chart of another particularly preferred method for preparing high-octane gasoline components provided by the application.

[0023] BRIEF DESCRIPTION OF DRAWINGS

[0024] 101, heat exchanger I 102, heating furnace 103, first reactor

[0025] 104, second reactor 105, elevator I 201, separation hopper I

[0026] 202, downpipe I 203, metering hopper 204, regenerator

[0027] 205, preheating zone 206, downpipe II 207, first coking zone

[0028] 208, downpipe III 209, second coking zone 210, downpipe IV

[0029] 211, drying zone 212, downpipe V 213, cooling zone

[0030] 214, lower hopper 215, elevator II 216, separation hopper II

[0031] 217, desulfurization tank 218, heat exchanger II 219, air cooler

[0032] 220. Dryer; 221. Regeneration fan; 222. Electric heater I

[0033] 223. First Space 224. Second Space 225. Annular Gap

[0034] 226. Electric heater II 227. Electric heater III 228. Cooler I

[0035] 301. Cooler II; 302. Three-phase separator; 303. Compressor

[0036] 304, Second Separator; 305, Pump; 306, Stabilizer Tower Detailed Implementation

[0037] The endpoints and any values ​​of the ranges disclosed herein are not limited to the precise ranges or values, and these ranges or values ​​should be understood to include values ​​close to these ranges or values. For numerical ranges, the endpoint values ​​of the various ranges, the endpoint values ​​of the various ranges and individual point values, and individual point values ​​can be combined with each other to obtain one or more new numerical ranges, which should be considered as specifically disclosed herein.

[0038] In this invention, unless otherwise stated, the moving bed reaction zone is connected to the regenerator via pipelines, and all reactors are connected to each other via pipelines.

[0039] As mentioned above, the present invention provides a method for preparing high-octane gasoline components, the method comprising:

[0040] (1) An organic oxygen-containing compound and naphtha are introduced into a moving bed reaction zone packed with an aromatization catalyst to carry out a contact reaction to obtain reaction product I, and the reaction product I is separated to obtain a gas phase stream, an oil phase stream and an aqueous phase stream.

[0041] (2) The catalyst flowing out from the bottom of the moving bed reaction zone is introduced into the regenerator for regeneration treatment to obtain a regenerated catalyst, and the regenerated catalyst is recycled back to the moving bed reaction zone to participate in the contact reaction;

[0042] The organic oxygen-containing compound is selected from at least one of alcohols, ethers, and esters; the mass ratio of the organic oxygen-containing compound to the naphtha is 1:0.1-10.

[0043] The aromatization catalyst comprises a support and an active component, and based on the total mass of the support, the active component contains 0.5-15% by mass of Group VA element oxides and 0.1-2% by mass of rare earth element oxides.

[0044] In the present application, the VA group element oxide refers to an oxide formed by a VA group element and an oxygen element, and the rare earth element oxide refers to an oxide formed by a rare earth element and an oxygen element.

[0045] The present application does not have special requirements for the separation method of the reaction product I, and known methods in the art can be used, preferably, the reaction product I is separated by a three-phase separator to obtain an aqueous phase, an oil phase and a gas phase. In order to reduce the influence of water on the catalyst and reduce the energy consumption of the reactor, the aqueous phase is discharged from the system.

[0046] According to a particularly preferred embodiment of the present application, in step (1), the organic oxygen-containing compound is selected from at least one of methanol, dimethyl ether, ethanol, glycerol, butanol, and isobutanol.

[0047] Preferably, in step (1), the mass ratio of the organic oxygen-containing compound to the naphtha is 1:0.2-5. The inventors found during research that the specific embodiment under this preferred condition can obtain a high-octane gasoline component with higher yield.

[0048] Preferably, in step (1), the moving bed reaction zone comprises at least two moving bed radial reactors connected in series.

[0049] Preferably, in step (1), the number of moving bed radial reactors is 2-5.

[0050] The arrangement of the moving bed radial reactors connected in series in the present application is not particularly required, as long as it can meet the requirements of the present application, for example, it can be stacked vertically or placed side by side. When the number of moving bed radial reactors is 4, two moving bed radial reactors can be stacked vertically as a group, and two groups of reactors can be placed side by side, and a catalyst lift is provided between each group.

[0051] According to a particularly preferred embodiment of the present application, in step (1), the naphtha is mixed with all the organic oxygen-containing compounds and introduced into the first moving bed radial reactor of the moving bed reaction zone, and sequentially flows through each of the subsequent moving bed radial reactors.

[0052] According to a particularly preferred embodiment of the present application, in step (1), the organic oxygen-containing compound is first divided into at least two streams, and the first stream of organic oxygen-containing compound is mixed with the naphtha and introduced into the first moving bed radial reactor.

[0053] Preferably, in step (1), the mass ratio of the first stream of organic oxygen-containing compound to the naphtha is 1:0.5-5.

[0054] It should be noted that the present application does not have special requirements for the way of introducing the remaining each stock of the organic oxygen-containing compound into the moving bed reaction zone, and any known way in the art can be used for introduction, for example, all the remaining each stock of the organic oxygen-containing compound can be introduced into any one of the downstream reactors, or the remaining each stock of the organic oxygen-containing compound can be introduced into a plurality of downstream reactors respectively, and the amount of each stock of the organic oxygen-containing compound can be evenly distributed, or gradually decreased along the flow direction of the liquid phase material, or gradually increased along the flow direction of the liquid phase material.

[0055] Preferably, in step (1), the content of the VA group element oxide in the active component is 1.5-12% by mass, and the content of the rare earth element oxide is 0.2-1.5% by mass.

[0056] Preferably, in step (1), the rare earth element oxide is lanthanum oxide and / or cerium oxide, and optionally further contains at least one of praseodymium oxide and neodymium oxide.

[0057] According to a particularly preferred embodiment of the present application, in step (1), the rare earth element oxide contains 20-60% by mass of lanthanum oxide, 40-80% by mass of cerium oxide, 0-10% by mass of praseodymium oxide, and 0-10% by mass of neodymium oxide, based on the total mass of the rare earth element oxide.

[0058] According to another particularly preferred embodiment of the present application, in step (1), the rare earth element oxide contains 20-40% by mass of lanthanum oxide, 40-60% by mass of cerium oxide, 10-18% by mass of praseodymium oxide, and 2-10% by mass of neodymium oxide, based on the total mass of the rare earth element oxide.

[0059] Preferably, in step (1), the carrier contains 40-75% by mass of ZSM-5 zeolite, 3-35% by mass of alumina, and 8-35% by mass of amorphous aluminum silicate, based on the total mass of the carrier.

[0060] More preferably, in step (1), the carrier contains 45-75% by mass of ZSM-5 zeolite, 10-30% by mass of alumina, and 10-30% by mass of amorphous aluminum silicate, based on the total mass of the carrier.

[0061] In the present application, ZSM-5 zeolite, alumina and silica sol are shaped and then calcined to form a carrier containing alumina and amorphous aluminum silicate, which not only increases the strength of the carrier, but also adjusts the distribution of B acid and L acid, and can inhibit the excessive aromatization reaction of olefins in the Fischer-Tropsch synthesis naphtha raw material. In addition, after the catalyst is modified by adding VA group elements and rare earth elements, the stability and carbon deposition resistance of the catalyst are significantly improved, and the catalyst can adapt to the high-temperature hydrothermal atmosphere caused by a large amount of water generated in the aromatization process of organic oxygen-containing compounds, so that the catalyst maintains high activity in this environment and prolongs the single-pass reaction life. In addition, it also has good regeneration performance.

[0062] Preferably, in step (1), the initial boiling point of the naphtha is 30-80℃, and the final boiling point is 130-200℃.

[0063] According to a particularly preferred embodiment of the present application, in step (1), the naphtha is selected from at least one of straight-run naphtha, raffinate, condensate, Fischer-Tropsch synthesis oil, and naphtha fraction produced by hydrogenation or coking process.

[0064] Preferably, in step (1), the contact reaction conditions at least include: reaction temperature is 250-550℃, mass space velocity is 0.2-2.0h -1 , reaction pressure is 0.3-1.0MPa.

[0065] More preferably, in step (1), the contact reaction conditions at least include: reaction temperature is 300-450℃, mass space velocity is 0.3-1.5h -1 , reaction pressure is 0.4-0.8MPa.

[0066] Preferably, in step (2), the method further comprises: introducing the oil phase stream and the gas phase stream into a fractionation zone for fractionation to obtain dry gas, liquefied gas and gasoline.

[0067] Preferably, in step (2), the regenerator includes, in sequence along the flow direction of the spent catalyst, a preheating zone, a decoking zone, a drying zone and a cooling zone, and a regeneration gas containing oxygen is introduced into the decoking zone to coke the spent catalyst.

[0068] Preferably, in step (2), the average residence time of the spent catalyst in the decoking zone is 10-600min, preferably 60-480min.

[0069] Preferably, in step (2), the decoking zone includes at least two segments of decoking zones arranged vertically from top to bottom.

[0070] Preferably, in step (2), the coke-burning zone comprises a first coke-burning zone and a second coke-burning zone arranged vertically from top to bottom, and the oxygen content of the regeneration gas in the second coke-burning zone is higher than that in the first coke-burning zone, and the inlet temperature of the regeneration gas in the second coke-burning zone is higher than that in the first coke-burning zone.

[0071] It should be noted that the coke-burning flue gas generated by the coke-burning zone of the regenerator is discharged from the top of the regenerator, part of which returns to the preheating zone of the regenerator as preheating gas, and the other part is mixed with nitrogen after desulfurization and drying, and is divided into two streams, and air is mixed into each stream as regeneration gas, and the oxygen content of the regeneration gas entering the second coke-burning zone is higher than that of the regeneration gas entering the first coke-burning zone, and the two streams of regeneration gas are heated and then enter the two coke-burning zones, respectively, and the flue gas generated by the two coke-burning zones is discharged from the bottom of the regenerator, between the two coke-burning zones, or from the top.

[0072] The present application adopts one-stage preheating and two-stage coke-burning, which improves the regeneration capacity, and can separately control the temperature of each coke-burning zone and the oxygen content of the inlet gas, so that the operation flexibility of the regenerator is large, the coke-burning is more sufficient, and in the case that the oxygen content of the inlet gas of the coke-burning zone is low, the local overheating point generated in the catalyst coke-burning process is reduced, the coke-burning is more uniform, the specific surface area loss of the catalyst is reduced, and the service life is prolonged.

[0073] Preferably, in step (2), the conditions of the first coke-burning zone at least include: the inlet temperature of the regeneration gas is 350-600℃, more preferably 400-500℃; the oxygen content of the regeneration gas is 0.1-3.0 vol.%, more preferably 0.3-2.0 vol.%; and the pressure is 0.1-2.0 MPa, more preferably 0.3-1.0 MPa.

[0074] Preferably, in step (2), the conditions of the second coke-burning zone at least include: the inlet temperature of the regeneration gas is 400-600℃, more preferably 440-550℃; the oxygen content of the regeneration gas is 0.2-5.0 vol.%, more preferably 0.3-3.0 vol.%; and the pressure is 0.1-2.0 MPa, more preferably 0.3-1.0 MPa.

[0075] It should be noted that the cooling gas is introduced into the cooling zone of the regenerator, the gas from the cooling zone is mixed with the drying gas, and then enters the drying zone after being heated; and the gas discharged from the drying zone can be discharged from the system or returned to the cooling zone after being cooled and dried for recycling, that is, the gas discharged from the drying zone is returned to the cooling zone after being cooled and dried. The present application does not have special requirements for the types of the cooling gas and the drying gas, as long as they meet the requirements of the present application. In order to effectively avoid the problem of over-temperature in the drying zone, the cooling gas and the drying gas used are both nitrogen.

[0076] Preferably, in step (2), the temperature of the drying zone inlet gas is 120-600°C, more preferably 250-500°C.

[0077] Preferably, in step (2), the temperature of the cooling zone outlet gas is not more than 200°C.

[0078] The following Figure 1 and Figure 2 illustrate a particularly preferred reaction process, catalyst regeneration process, gas circulation process in the regenerator and reaction product separation and fractionation process.

[0079] As Figure 1 shown, the reaction process is as follows: naphtha and all organic oxygen compounds are introduced into heat exchanger I 101, heated in furnace 102 after heat exchange with reaction products, then enter first reactor 103, radially contact with the aromatization catalyst therein and perform contact reaction, after reaction, the material directly enters second reactor 104, radially contacts with the aromatization catalyst therein and performs contact reaction, after reaction, the material enters heat exchanger I 101, after heat exchange with the feedstock, is sent to the fractionation system through a pipeline.

[0080] The catalyst regeneration process is as follows: spent catalyst flowing out from the bottom of second reactor 104 enters elevator I 105, is lifted to spent catalyst separation hopper I 201 with circulating nitrogen, in separation hopper I 201, the spent catalyst and the nitrogen gas stream are separated, the nitrogen gas stream carrying part of the dust enters the dust collection system, the dust therein is separated from the nitrogen, the separated nitrogen is recycled as catalyst lifting gas (not shown in the figure), the spent catalyst enters metering hopper 203 through downcomer I 202, enters regenerator 204 through a pipeline for regeneration, in regenerator 204, the spent catalyst sequentially passes through preheating zone 205 (axial bed), downcomer II 206, a section of coking zone 207 (radial bed) surrounded by a screen, downcomer III 208, a second section of coking zone 209 (radial bed) surrounded by a screen, downcomer IV 210, drying zone 211 (axial bed), downcomer V 212, cooling zone 213 (axial bed) from top to bottom. The regenerated catalyst from regenerator 204 passes through lower hopper 214, enters regenerated catalyst elevator II 215 through a pipeline, the regenerated catalyst is lifted to regenerated catalyst separation hopper II 216 at the top of first reactor 103 with circulating nitrogen, in separation hopper II 216, the regenerated catalyst is separated from the nitrogen gas stream, the nitrogen gas stream carrying part of the dust enters the dust collection system through a pipeline, the regenerated catalyst enters first reactor 103, in first reactor 103, the catalyst enters second reactor 104 by gravity, then enters spent catalyst elevator I 105 by gravity, to restart a new cycle.

[0081] The nitrogen gas used for lifting the catalyst in the above-mentioned catalyst regeneration process can be supplemented at the spent catalyst lifter I 105, lifter II 218 and lower hopper 214.

[0082] The gas circulation process in the regenerator is as follows: the flue gas from the top of the regenerator, part of which enters the preheating zone 205 of the regenerator as preheating gas, and the remaining part enters the desulfurization tank 217 containing flue gas desulfurizer, which is a calcium-based desulfurizer or other high-temperature flue gas desulfurizer known in the art. The flue gas after SO2 removal from the desulfurization tank 217 is discharged from the pipeline, part of which is directly discharged from the system, and the remaining part enters the heat exchanger II 218, exchanges heat with the cooled and dehydrated regenerated gas, and then enters the dryer 220 through the air cooler 219. After the water content in the gas is reduced to below 2000 ppm (preferably below 1000 ppm, and more preferably below 200 ppm), the cold air or the mixture of air and nitrogen is mixed to become the regenerated gas. The regenerated gas is sent to the regenerator fan 221, and then enters the heat exchanger II 218 to exchange heat with the hot flue gas generated in the coking zone. Then part of the regenerated gas enters the electric heater I 222 and is heated to 350-600°C (preferably 400-500°C), and then enters the top of the first coking zone 207 through the annular gap 225, with the oxygen content in the regenerated gas controlled at 0.1-3.0 vol% (preferably 0.3-2.0 vol%). The regenerated gas passes through the bed of spent catalyst in a centripetal radial direction, contacts the spent catalyst, and the flue gas generated after coking enters the first space 223, and then is discharged from the regenerator for recycling, with a small amount of the discharged gas entering the regenerated gas preheating zone to preheat the catalyst. The other part of the regenerated gas is mixed with air or a mixture of nitrogen and air, and then enters the electric heater II 226 through the pipeline and is heated to 400-600°C, preferably 440-550°C, with the oxygen content in the regenerated gas controlled at 0.2-5.0 vol%, preferably 0.3-3.0 vol%. The regenerated gas enters the second space 224 through the pipeline, passes through the bed of spent catalyst in the second coking zone in a centrifugal radial direction, contacts the catalyst, and the generated gas is recycled through the pipeline.

[0083] The nitrogen gas from the pipeline enters the cooling zone, cools the catalyst, and the gas from the cooling zone 213 enters the electric heater III 227 through the pipeline, is heated to 120-600°C, preferably 250-500°C, and then enters the catalyst drying zone 211. The gas discharged from the drying zone 211 is discharged from the system or enters the cooler I 228 after cooling and is recycled.

[0084] As Figure 2As shown, the separation and fractionation process of the reaction product I is as follows: the reaction product I from the reaction system enters the product cooler II 301 through a pipeline, is then cooled, and enters the three-phase separator 302, where it is separated into a gas phase, an oil phase, and an aqueous phase. The aqueous phase is discharged from the three-phase separator 302 through a pipeline at the bottom of the three-phase separator 302, and the gas phase is compressed by the compressor 303 through a pipeline at the top of the three-phase separator 302, and then enters the second separator 304. Dry gas is discharged from the system through a pipeline at the top of the second separator 304. The material at the bottom of the second separator 304 is mixed with the oil phase separated from the three-phase separator 302 through a pipeline, and then enters the stabilizer 306 through a pipeline by the pump 305. The liquefied gas is discharged from the system through a pipeline at the top of the stabilizer, and the gasoline component is discharged from the system through a pipeline at the bottom. The dry gas refers to hydrogen, methane, ethane, and ethylene.

[0085] The following will be described in detail by examples. In the following examples, the various raw materials used are commercially available products, unless otherwise specified. Figure 3 A particularly preferred reaction process of the present application is described below, which is similar to the process flow in Figure 1 The difference is that naphtha and part of the organic oxygen-containing compounds are introduced into the heat exchanger I 101, heated by the heater 102 after heat exchange with the reaction product, and then enter the first reactor 103 to contact with the aromaticization catalyst therein and perform a contact reaction. After the reaction, the material is mixed with the remaining part of the organic oxygen-containing compounds, enters the second reactor 104, contacts with the aromaticization catalyst therein, and performs a contact reaction. After the reaction, the material enters the heat exchanger I 101 to exchange heat with the raw material, and then is discharged from the system through a pipeline.

[0086] Under the process flow conditions shown in Figure 3 The gas circulation process in the regenerator is the same as that in Figure 1 The separation and fractionation process of the reaction product I is the same as that in Figure 2

[0087] The following will be described in detail by examples. In the following examples, the various raw materials used are commercially available products, unless otherwise specified.

[0088] HZSM-5 zeolite powder: purchased from Shanghai Huaheng Chemical Factory;

[0089] Pseudo-boehmite powder: purchased from Sasol Company;

[0090] Aluminum hydroxide powder: purchased from Qilu Catalyst Factory;

[0091] Mixed rare earth chloride: purchased from Aladdin;

[0092] Reformed PS-Ⅵ noble metal catalyst: purchased from Hunan Jianchang Petrochemical Co., Ltd.;

[0093] In the following examples, the content of the active component and the carrier in the aromaticization catalyst is obtained by X-ray fluorescence method.​

[0094] In the absence of the opposite case, the properties of the naphtha in the following examples are as shown in Table 1.

[0095] Table 1

[0096] Properties Data Density (20°C), g / cm 3 ]] 0.7446 Sulfur content, ppm ＜0.5 Nitrogen content, ppm ＜0.5 Bromine number, g Br / 100 g ＜0.1 Initial boiling point 84.5 10％ 100.8 30％ 109.2 50％ 118.8 90％ 15.8 Final boiling point 165.6 Hydrocarbon group composition, mass % Alkanes 46.44 Naphthenes 42.52 Aromatics 10.84

[0097] Preparation Example 1

[0098] This preparation example illustrates the preparation process of the aromatization catalyst S1.

[0099] (1) Take 67.6 grams of pseudoboehmite (produced by Sasol). SB powder (75% by mass) was mixed with 380g of a 1.1% by mass nitric acid aqueous solution under stirring at 100rpm. After 2 hours of gelation, an alumina sol with an alumina content of 11% by mass was obtained. 50g of hydrogen-form ZSM-5 zeolite (SiO2 / Al2O3 molar ratio of 56) was added to the sol, and the mixture was stirred at 100rpm for 3 hours to obtain a zeolite-containing slurry with a solid content of 20% by mass. The slurry was then drop-formed into spheres in an oil-ammonia column at a dropping temperature of 15℃. The oil phase of the oil-ammonia column was kerosene with a thickness of 10cm, and the ammonia phase had a thickness of 200cm and an ammonia concentration of 6% by mass. The wet spheres were removed from the bottom of the ammonia layer, dried at 60℃ for 10 hours, and activated at 550℃ for 3 hours to obtain a small spherical composite carrier with an average particle size of 1.8mm.

[0100] (2) The composite carrier prepared above was loaded into a tubular reactor, heated to 550°C in an air flow at 0.1 MPa, and treated with water vapor for 4 hours.

[0101] (3) Take 100g of the carrier after steam treatment, impregnate it in 50ml of phosphoric acid solution with a concentration of 100mg / ml for 1 hour, and dry the impregnated solid material at 120℃ for 2 hours. Then add 100ml of mixed rare earth chloride (wherein, the content of lanthanum oxide is 40% by mass and the content of cerium oxide is 60% by mass) aqueous solution with a concentration of 10mg / ml, impregnate it at 80℃ for 2 hours, and dry the impregnated solid material at 120℃ for 8 hours. Then calcine it at 550℃ for 4 hours to obtain aromatization catalyst S1.

[0102] In the aromatization catalyst S1, based on the total mass of the support, the active component contains 5.29% by mass of P2O5 and 0.53% by mass of rare earth element oxides (of which, the content of lanthanum oxide is 40% by mass and the content of cerium oxide is 60% by mass).

[0103] The carrier contains 69.8% by mass of ZSM-5 zeolite, 12.2% by mass of alumina, and 18% by mass of amorphous aluminum silicate, based on the total mass of the carrier.

[0104] Comparative Preparation Example 1

[0105] This comparative preparation example refers to the method of Preparation Example 1, except that in step (2), 5 g of rare earth chloride is applied to obtain the aromatization catalyst DS1.

[0106] In the aromatization catalyst DS1, the active component contains 5.29% by mass of P2O5 and 2.5% by mass of rare earth element oxide (of which the content of lanthanum oxide is 40% by mass and the content of cerium oxide is 60% by mass), based on the total mass of the carrier.

[0107] The carrier contains 69.8% by mass of ZSM-5 zeolite, 12.2% by mass of alumina, and 18% by mass of amorphous aluminum silicate, based on the total mass of the carrier.

[0108] Example 1

[0109] According to the process shown in Figure 1 , the aromatization catalyst S1 prepared in Preparation Example 1 is used, methanol and naphtha shown in Table 1 are used as raw materials, the mass ratio of methanol to naphtha is 1:3, the reaction conditions are: temperature is 350°C, pressure is 0.3 MPa, mass space velocity of raw materials is 0.5 h -1 , and the reaction time is 48 hours.

[0110] The regenerator pressure is 0.7 MPa, the inlet temperature of the regenerating gas in the first coking zone is 470°C, the oxygen content in the regenerating gas is 0.5% by volume; the inlet temperature of the regenerating gas in the second coking zone is 480°C, the oxygen content in the regenerating gas is 0.6% by volume, the volume ratio of the regenerating gas to the spent catalyst is 3000:1, and the residence time of the spent catalyst in the coking zone of the regenerator is 120 minutes.

[0111] The temperature of the inlet gas in the drying zone is 490°C, and the temperature of the outlet gas in the cooling zone is 165°C. The specific catalytic reaction results are shown in Table 2.

[0112] Example 2

[0113] According to the process shown in Figure 1 , the aromatization catalyst S1 prepared in Preparation Example 1 is used, methanol and naphtha shown in Table 1 are used as raw materials, the mass ratio of methanol to naphtha is 4:6, the reaction conditions are: temperature is 350°C, pressure is 0.3 MPa, mass space velocity of raw materials is 0.5 h -1 , and the reaction time is 48 hours.

[0114] The regenerator pressure is 0.7 MPa. The inlet temperature of the regenerated gas in the first coke burning zone is 470℃, and the oxygen content in the regenerated gas is 0.5% by volume. The inlet temperature of the regenerated gas in the second coke burning zone is 480℃, and the oxygen content in the regenerated gas is 0.6% by volume. The volume ratio of regenerated gas to the catalyst to be generated is 3000:1, and the residence time of the catalyst to be generated in the coke burning zone of the regenerator is 120 minutes.

[0115] The inlet gas temperature in the drying zone is 490℃, and the outlet gas temperature in the cooling zone is 165℃. Specific catalytic reaction results are shown in Table 2.

[0116] Example 3

[0117] according to Figure 1 The process flow shown uses the aromatization catalyst S1 prepared in Preparation Example 1, with dimethyl ether and naphtha as raw materials, in a mass ratio of 1:3. The reaction conditions are: temperature 350℃, pressure 0.3 MPa, and feed mass hourly space velocity 0.5 h⁻¹. -1 The reaction time is 48 hours;

[0118] The regenerator pressure is 0.7 MPa. The inlet temperature of the regenerated gas in the first coke burning zone is 470℃, and the oxygen content in the regenerated gas is 0.5% by volume. The inlet temperature of the regenerated gas in the second coke burning zone is 480℃, and the oxygen content in the regenerated gas is 0.6% by volume. The volume ratio of regenerated gas to the catalyst to be generated is 3000:1, and the residence time of the catalyst to be generated in the coke burning zone of the regenerator is 120 minutes.

[0119] The inlet gas temperature in the drying zone is 490℃, and the outlet gas temperature in the cooling zone is 165℃. Specific catalytic reaction results are shown in Table 2.

[0120] Example 4

[0121] according to Figure 1 The process flow shown uses the aromatization catalyst S1 prepared in Preparation Example 1, with glycerol and naphtha as raw materials.

[0122] The mass ratio of glycerol to naphtha was 1:3, and the reaction conditions were: temperature 350℃, pressure 0.3MPa, and feed mass hourly space velocity (WHSV) 0.5h⁻¹. -1 The reaction time is 48 hours;

[0123] The regenerator pressure is 0.7 MPa. The inlet temperature of the regenerated gas in the first coke burning zone is 470℃, and the oxygen content in the regenerated gas is 0.5% by volume. The inlet temperature of the regenerated gas in the second coke burning zone is 480℃, and the oxygen content in the regenerated gas is 0.6% by volume. The volume ratio of regenerated gas to the catalyst to be generated is 3000:1, and the residence time of the catalyst to be generated in the coke burning zone of the regenerator is 120 minutes.

[0124] The inlet gas temperature in the drying zone was 490℃, and the outlet gas temperature in the cooling zone was 165℃. Specific catalytic reaction results are shown in Table 2.

[0125] Example 5

[0126] according to Figure 3 The process flow shown uses the aromatization catalyst S1 prepared in Preparation Example 1, with methanol and naphtha as raw materials. The methanol is divided into two streams. The first stream of methanol and naphtha, with a mass ratio of 1:4, is mixed and introduced into the first reactor for reaction. Then, the reaction product is mixed with the remaining methanol and introduced into the second reactor for reaction. The conditions for both reactions are: temperature 350°C, pressure 0.3 MPa, and feed mass hourly space velocity 0.5 h⁻¹. -1 The reaction time is 48 hours.

[0127] The regenerator pressure is 0.7 MPa, the inlet temperature of the regenerated gas in the first coke burning zone is 470℃, and the oxygen content in the regenerated gas is 0.5% by volume; the inlet temperature of the regenerated gas in the second coke burning zone is 480℃, and the oxygen content in the regenerated gas is 0.6% by volume. The volume ratio of regenerated gas to the catalyst to be generated is 3000:1, and the residence time of the catalyst to be generated in the coke burning zone of the regenerator is 120 minutes.

[0128] The inlet gas temperature in the drying zone was 490℃, and the outlet gas temperature in the cooling zone was 165℃. Specific catalytic reaction results are shown in Table 2.

[0129] Comparative Example 1

[0130] according to Figure 1 The process flow shown uses the aromatization catalyst S1 prepared in Preparation Example 1, with methanol and water as raw materials in a methanol-to-water mass ratio of 4:6. The reaction conditions are: temperature 350℃, pressure 0.3 MPa, and feed mass hourly space velocity (WHSV) 0.5 h⁻¹. -1 The reaction time is 48 hours;

[0131] The regenerator pressure is 0.7 MPa. The inlet temperature of the regenerated gas in the first coke burning zone is 470℃, and the oxygen content in the regenerated gas is 0.5% by volume. The inlet temperature of the regenerated gas in the second coke burning zone is 480℃, and the oxygen content in the regenerated gas is 0.6% by volume. The volume ratio of regenerated gas to the catalyst to be generated is 3000:1, and the residence time of the catalyst to be generated in the coke burning zone of the regenerator is 120 minutes.

[0132] The inlet gas temperature in the drying zone was 490℃, and the outlet gas temperature in the cooling zone was 165℃. Specific catalytic reaction results are shown in Table 2.

[0133] Comparative Example 2

[0134] according to Figure 1The process flow shown uses the aromatization catalyst S1 prepared in Preparation Example 1, with methanol as the raw material, and the reaction conditions are: temperature 350℃, pressure 0.3MPa, and feed mass hourly space velocity 0.5h⁻¹. -1 The reaction time is 48 hours;

[0135] The regenerator pressure is 0.7 MPa. The inlet temperature of the regenerated gas in the first coke burning zone is 470℃, and the oxygen content in the regenerated gas is 0.5% by volume. The inlet temperature of the regenerated gas in the second coke burning zone is 480℃, and the oxygen content in the regenerated gas is 0.6% by volume. The volume ratio of regenerated gas to the catalyst to be generated is 3000:1, and the residence time of the catalyst to be generated in the coke burning zone of the regenerator is 120 minutes.

[0136] The inlet gas temperature in the drying zone was 490℃, and the outlet gas temperature in the cooling zone was 165℃. Specific catalytic reaction results are shown in Table 2.

[0137] Comparative Example 3

[0138] according to Figure 1 The process flow shown uses the aromatization catalyst DS1 prepared in Comparative Preparation Example 1, with methanol and naphtha as raw materials (as shown in Table 1) in a methanol-to-naphtha mass ratio of 1:3. The reaction conditions are: temperature 350°C, pressure 0.3 MPa, and feed mass hourly space velocity (WHSV) 0.5 h⁻¹. -1 The reaction time is 48 hours;

[0139] The regenerator pressure is 0.7 MPa, the inlet temperature of the regenerated gas in the first coke burning zone is 470℃, and the oxygen content in the regenerated gas is 0.5% by volume; the inlet temperature of the regenerated gas in the second coke burning zone is 480℃, and the oxygen content in the regenerated gas is 0.6% by volume. The volume ratio of regenerated gas to the catalyst to be generated is 3000:1, and the residence time of the catalyst to be generated in the coke burning zone of the regenerator is 120 minutes.

[0140] The inlet gas temperature in the drying zone was 490℃, and the outlet gas temperature in the cooling zone was 165℃. Specific catalytic reaction results are shown in Table 2.

[0141] Comparative Example 4

[0142] according to Figure 1 The process flow shown uses catalyst A prepared in Example 1 of CN109569703A, with methanol and naphtha as shown in Table 1 as raw materials. The mass ratio of methanol to naphtha is 1:3. The reaction conditions are: temperature 350℃, pressure 0.3MPa, and feed mass hourly space velocity 0.5h⁻¹. -1 The reaction time is 48 hours;

[0143] The regenerator pressure is 0.7 MPa. The inlet temperature of the regenerated gas in the first coke burning zone is 470℃, and the oxygen content in the regenerated gas is 0.5% by volume. The inlet temperature of the regenerated gas in the second coke burning zone is 480℃, and the oxygen content in the regenerated gas is 0.6% by volume. The volume ratio of regenerated gas to the catalyst to be generated is 3000:1, and the residence time of the catalyst to be generated in the coke burning zone of the regenerator is 120 minutes.

[0144] The inlet gas temperature in the drying zone is 490℃, and the outlet gas temperature in the cooling zone is 165℃. Specific catalytic reaction results are shown in Table 2.

[0145] Comparative Example 5

[0146] according to Figure 1 The process flow shown uses a modified PS-VI noble metal catalyst, with methanol and naphtha (as shown in Table 1) as feedstocks at a methanol-to-naphtha mass ratio of 1:3. The reaction conditions are: temperature 350℃, pressure 0.3 MPa, and feedstock mass hourly space velocity (WHSV) 0.5 h⁻¹. -1 The reaction time is 48 hours;

[0147] The regenerator pressure is 0.7 MPa. The inlet temperature of the regenerated gas in the first coke burning zone is 470℃, and the oxygen content in the regenerated gas is 0.5% by volume. The inlet temperature of the regenerated gas in the second coke burning zone is 480℃, and the oxygen content in the regenerated gas is 0.6% by volume. The volume ratio of regenerated gas to the catalyst to be generated is 3000:1, and the residence time of the catalyst to be generated in the coke burning zone of the regenerator is 120 minutes.

[0148] The inlet gas temperature in the drying zone is 490℃, and the outlet gas temperature in the cooling zone is 165℃. Specific catalytic reaction results are shown in Table 2.

[0149] Comparative Example 6

[0150] The process was carried out according to Example 2, except that the catalyst regeneration process adopted existing technology, namely, the pressure transformation and flow control zone was located in the lower part of the regenerator; the regenerator consisted of two parts: a coking zone and a drying zone. The flue gas generated from coking was dried, mixed with supplementary air, and returned to the coking zone of the regenerator. The regenerator pressure was 0.35 MPa, the inlet temperature of the regeneration gas in the coking zone was 480°C, the oxygen content in the regeneration gas was 0.75% by volume, the volume ratio of regeneration gas to the catalyst to be regenerated was 3000:1, and the residence time of the catalyst in the coking zone of the regenerator was 120 minutes. Specific catalytic reaction results are shown in Table 2.

[0151] Table 2

[0152]

[0153] Table 2 (continued)

[0154]

[0155] As can be seen from the results of Table 2, the method provided by the present application can be used to prepare high-octane gasoline, the gasoline yield in hydrocarbons can reach 75-85 mass%, the C5 + The research method octane number reaches 93-99, the liquefied gas yield reaches 15-25 mass%, the low-value byproduct dry gas (C1+C2 hydrocarbons and a small amount of hydrogen) yield is 0.2-1.0 mass%, the C5 + aromatic hydrocarbons is 60 mass% or more, and each index meets the environmental protection requirements; meanwhile, the method provided by the present application can improve the coking capacity of the regenerator, maintain the catalyst activity, and the specific surface area of the catalyst is 290 m 2 / g or more after 20 times of regeneration.

[0156] The preferred embodiments of the present application are described in detail above, but the present application is not limited thereto. Within the technical concept of the present application, various simple modifications can be made to the technical solutions of the present application, including that each technical feature is combined in any other suitable manner, and these simple modifications and combinations should also be considered as the disclosed content of the present application, and all fall within the protection scope of the present application.

Claims

1. A method for preparing high-octane gasoline components, characterized in that, The method includes: (1) Organic oxygen-containing compounds and naphtha are introduced into a moving bed reaction zone packed with aromatization catalyst to carry out a contact reaction to obtain reaction product I, and the reaction product I is separated to obtain a gas phase stream, an oil phase stream and an aqueous phase stream. (2) The unregenerated catalyst flowing out from the bottom of the moving bed reaction zone is introduced into the regenerator for regeneration treatment to obtain a regenerated catalyst, and the regenerated catalyst is recycled back to the moving bed reaction zone to participate in the contact reaction; The organic oxygen-containing compound is selected from at least one of dimethyl ether and glycerol; the mass ratio of the organic oxygen-containing compound to the naphtha is 1:0.1-10. The aromatization catalyst comprises a support and an active component, and based on the total mass of the support, the active component contains 5.29-12% by mass of Group VA element oxides and 0.2-1.5% by mass of rare earth element oxides.

2. The method according to claim 1, wherein, In step (1), the mass ratio of the organic oxygenated compound to the naphtha is 1:0.2-5.

3. The method according to claim 1 or 2, wherein, In step (1), the moving bed reaction zone comprises at least two moving bed radial reactors connected in series.

4. The method according to claim 3, wherein, In step (1), the number of moving bed radial reactors is 2-5.

5. The method according to claim 3, wherein, In step (1), naphtha is mixed with all the organic oxygen-containing compounds and introduced into the first moving bed radial reactor of the moving bed reaction zone, and then flows sequentially through each of the subsequent moving bed radial reactors.

6. The method according to claim 3, wherein, In step (1), the organic oxygen-containing compound is first divided into at least two streams, and the first stream of organic oxygen-containing compound is mixed with the naphtha and introduced into the first moving bed radial reactor.

7. The method according to claim 6, wherein, In step (1), the mass ratio of the first organic oxygenated compound to the naphtha is 1:0.5-5.

8. The method according to claim 1 or 2, wherein, In step (1), the rare earth element oxide is selected from at least one of lanthanum oxide, cerium oxide, praseodymium oxide, and neodymium oxide.

9. The method according to claim 8, wherein, In step (1), based on the total mass of the rare earth element oxides, the rare earth element oxides contain 20-60% by mass of lanthanum oxide, 40-80% by mass of cerium oxide, 0-10% by mass of praseodymium oxide and 0-10% by mass of neodymium oxide.

10. The method according to claim 9, wherein, In step (1), based on the total mass of the rare earth element oxides, the rare earth element oxides contain 20-40% by mass of lanthanum oxide, 40-60% by mass of cerium oxide, 10-18% by mass of praseodymium oxide and 2-10% by mass of neodymium oxide.

11. The method according to claim 1 or 2, wherein, In step (1), based on the total mass of the carrier, the carrier contains 40-75% by mass of ZSM-5 zeolite, 3-35% by mass of alumina and 8-35% by mass of amorphous aluminum silicate.

12. The method according to claim 11, wherein, In step (1), based on the total mass of the carrier, the carrier contains 45-75% by mass of ZSM-5 zeolite, 10-30% by mass of alumina and 10-30% by mass of amorphous aluminum silicate.

13. The method according to claim 1 or 2, wherein, In step (1), the initial boiling point of the naphtha is 30-80℃ and the final boiling point is 130-200℃.

14. The method according to claim 1 or 2, wherein, In step (1), the naphtha is selected from at least one of straight-run naphtha, raffinate, condensate, Fischer-Tropsch synthetic oil, and naphtha fractions produced by hydrotreating or coking processes.

15. The method according to claim 1 or 2, wherein, In step (1), the conditions for the contact reaction include at least: a reaction temperature of 250℃-550℃ and a mass hourly space velocity of 0.2-2.0 h⁻¹. -1 The reaction pressure is 0.3-1.0 MPa.

16. The method according to claim 1 or 2, wherein, In step (1), the conditions for the contact reaction include at least: a reaction temperature of 300℃-450℃ and a mass hourly space velocity of 0.3-1.5 h⁻¹. -1 The reaction pressure is 0.4-0.8 MPa.

17. The method according to claim 1 or 2, wherein, In step (2), the method further includes: introducing the oil phase stream and the gas phase stream into the fractionation zone for fractionation to obtain dry gas, liquefied gas and gasoline.

18. The method according to claim 1 or 2, wherein, In step (2), the regenerator includes a preheating zone, a coking zone, a drying zone and a cooling zone in sequence along the flow direction of the catalyst to be regenerated, and oxygen-containing regeneration gas is introduced into the coking zone so that the catalyst to be regenerated is coked and regenerated.

19. The method according to claim 18, wherein, In step (2), the average residence time of the catalyst to be generated in the coking zone is 10-600 min.

20. The method according to claim 18, wherein, In step (2), the average residence time of the catalyst to be generated in the coking zone is 60-480 min.

21. The method according to claim 18, wherein, In step (2), the charred zone includes at least two charred zones arranged vertically from top to bottom.

22. The method according to claim 21, wherein, In step (2), the coking zone includes a first coking zone and a second coking zone arranged vertically from top to bottom. The oxygen content in the regeneration gas of the second coking zone is higher than that in the regeneration gas of the first coking zone, and the inlet temperature of the regeneration gas of the second coking zone is higher than that of the regeneration gas of the first coking zone.

23. The method according to claim 22, wherein, In step (2), the conditions of the coking zone include at least the following: the inlet temperature of the regenerated gas is 350-600℃; the oxygen content in the regenerated gas is 0.1-3.0% by volume; and the pressure is 0.1-2.0 MPa.

24. The method according to claim 23, wherein, In step (2), the conditions of the coking zone include at least the following: the inlet temperature of the regenerated gas is 400-500℃; the oxygen content in the regenerated gas is 0.3-2.0% by volume; and the pressure is 0.3-1.0 MPa.

25. The method according to claim 22, wherein, In step (2), the conditions of the second-stage coking zone include at least the following: the inlet temperature of the regenerated gas is 400-600℃; the oxygen content in the regenerated gas is 0.2-5.0% by volume; and the pressure is 0.1-2.0 MPa.

26. The method according to claim 22, wherein, In step (2), the conditions of the second-stage coking zone include at least the following: the inlet temperature of the regenerated gas is 440-550℃; the oxygen content in the regenerated gas is 0.3-3.0% by volume; and the pressure is 0.3-1.0 MPa.

27. The method according to claim 18, wherein, In step (2), the temperature of the gas entering the drying zone is 120-600℃.

28. The method according to claim 18, wherein, In step (2), the temperature of the gas entering the drying zone is 250-500℃.

29. The method according to claim 18, wherein, In step (2), the outlet gas temperature of the cooling zone is not greater than 200°C.

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