Fluidized catalytic conversion process for maximizing propylene production.

Abstract

A fluid catalytic conversion method for maximizing the production of propylene, comprising the steps of: 1) contacting a heavy feedstock with a catalytic conversion catalyst having a temperature of 650°C or higher for reaction; 2) contacting a hydrocarbon oil feedstock having an olefin content of 50% by weight or more with a catalytic conversion catalyst after the reaction of step 1); 3) separating a first catalytic cracking distillate and a second catalytic cracking distillate from the resulting reacted petroleum gas; 4) separating an olefin-rich stream from the first catalytic cracking distillate; and 5) recycling the olefin-rich stream. The method of the present application can effectively improve the yield of propylene and realize the effective utilization of petroleum resources.

Description

Detailed Description of the Invention

[0001] CROSS-REFERENCE TO RELATED APPLICATIONS This application claims priority to Chinese Patent Application No. 202110032113.X, entitled "Catalytic Conversion Method for Maximizing Propylene Production," filed on January 11, 2023, and Chinese Patent Application No. 202110296904.3, entitled "Catalytic Conversion Method for Maximizing Propylene Production," filed on March 19, 2023, the contents of which are incorporated herein by reference in their entireties.

[0002] [Technical Field] This application relates to the field of fluid catalytic conversion, and more particularly to a fluid catalytic conversion process for maximizing propylene production.

[0003] [Background technology] The petrochemical industry is an important pillar industry, providing large quantities of chemical raw materials for industry, agriculture, transportation, national defense, and other sectors. Propylene and ethylene are the two most important basic raw materials for the modern petrochemical industry. However, as oil field production increases, the available yield of conventional crude oil gradually decreases, crude oil quality deteriorates, and crude oil tends to become inferior and heavy. Propylene is an important organic chemical raw material, primarily used in the preparation of acrylonitrile, propylene oxide, acetone, and other compounds. Ethylene and propylene are increasingly used as important chemical intermediates for the preparation of various important organic chemical raw materials, synthetic resins, synthetic rubber, and various microchemicals.

[0004] The traditional method of preparing ethylene and propylene by steam cracking has a high demand for light hydrocarbons such as naphtha. Existing crude oil is usually heavy, and light chemical oils have difficulty meeting the demand for ethylene and propylene feedstocks. Research institutes predict that the global average annual growth rate of the gasoline pool will be less than 1% from 2018 to 2026, while propylene will increase by approximately 4%. The appropriate use of high-carbon olefins in the refinery process of preparing ethylene and propylene by cracking can help petrochemical companies achieve their goals of improving quality and efficiency, as well as shortening the time required for energy transition.

[0005] CN101092323A discloses a method for preparing ethylene and propylene from a mixture of C4-C8 olefins. The method involves reacting the mixture at a reaction temperature of 400-600°C and an absolute pressure of 0.02-0.3 MPa, and recycling 30-90 wt% of the C4 fraction to the reactor after separation in a separator for further cracking. The method improves olefin conversion by recycling primarily the C4 fraction, and the resulting ethylene and propylene account for more than 62% of the total weight of the olefin feedstock. However, the method suffers from problems including a relatively low ethylene / propylene ratio, a lack of flexibility in adjusting according to market demand, low reaction selectivity, a high butylene content in the product, and energy consumption for C4 separation.

[0006] CN101239878A is a method using a mixture rich in C4+ olefins as raw material, and the reaction temperature is 400-680°C, the reaction pressure is -0.09MPa-1.0MPa, and the reaction time is 0.1-50h. -1 wherein the resulting product has an ethylene / propylene ratio of less than 0.41, and as the temperature increases, the ethylene / propylene ratio increases and the production of hydrogen, methane, and ethane increases.

[0007] Therefore, there is a need in the art for new fluid catalytic conversion processes to increase ethylene and propylene production and improve ethylene and propylene selectivity.

[0008] Summary of the Invention It is an object of the present application to provide a method for maximizing the production of propylene from a hydrocarbon-containing feedstock, which is effective in increasing the yield of propylene while providing high ethylene yields, high selectivity, and low dry gas yields.

[0009] In order to achieve the above-mentioned object, the present application provides a fluid catalytic conversion method for maximizing propylene production, comprising the following steps 1) to 6): 1) introducing a heavy feedstock into a first reaction zone of a fluid catalytic conversion reactor, contacting the heavy feedstock with a catalytic conversion catalyst having a temperature of 650°C or higher, and reacting under first catalytic conversion reaction conditions; 2) introducing a hydrocarbon oil feedstock having an olefin content of 50 wt.% or more into a second reaction zone downstream of the fluid catalytic conversion reactor of the first reaction zone after the reaction of step 1), contacting the catalytic conversion catalyst from the first reaction zone and reacting under second catalytic conversion conditions; 3) separating the effluent of the fluidized catalytic conversion reactor to obtain a reaction product and a used catalyst, and performing a first separation on the reaction product to obtain ethylene, propylene, butylene, a first catalytic cracking distillate, and a second catalytic cracking distillate; the first catalytic cracking distillate has an initial boiling point greater than 20°C, the second catalytic cracking distillate has an end boiling point less than 550°C, and the cut point between the first catalytic cracking distillate and the second catalytic cracking distillate is within the range of 140°C to 250°C; 4) subjecting the first catalytic cracking distillate to a second separation to obtain an olefin-rich stream having a C5+ olefin content of at least 50 wt%; 5) recycling at least a portion of the olefin-rich stream to step 2) for further reaction; and 6) recycling at least a portion of the butylenes separated in step 3) upstream of the fluidized catalytic conversion reactor where the heavy feedstock is introduced and contacted for reaction with the catalytic conversion catalyst under third catalytic conversion conditions, wherein the first catalytic conversion conditions include a reaction temperature of 500 to 800°C, a reaction pressure of 0.05 to 1 MPa, a reaction time of 0.01 to 100 seconds, and a weight ratio of the catalytic conversion catalyst to the heavy feedstock of (1 to 200):1; and the second catalytic conversion conditions include a reaction temperature of 400-680°C, a reaction pressure of 0.05-1 MPa, a reaction time of 0.01-100 seconds, and a weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock of (1-100):1.

[0010] The third catalytic conversion conditions include a reaction temperature of 650-800°C, a reaction pressure of 0.05-1 MPa, a reaction time of 0.01-10 seconds, and a weight ratio of the catalytic conversion catalyst to the butylene of (20-200):1.

[0011] Preferably, the method further comprises the following step 2a): 2a) introducing an oxygen-containing organic compound into said second reaction zone of said fluidized catalytic conversion reactor for reaction contact with a catalytic conversion catalyst therein under fourth catalytic conversion conditions comprising: Reaction temperature of 300-550°C, Reaction pressure of 0.05 to 1 MPa, Response time of 0.01 to 100 seconds, a weight ratio of the catalytic conversion catalyst to the oxygen-containing organic compound of (1-100):1;

[0012] Preferably, the method further comprises the following step 7): 7) hydrotreating the second catalytic cracking distillate to obtain a hydrogenated catalytic cracking distillate, and recycling the hydrogenated catalytic cracking distillate to the first reaction zone of the fluid catalytic conversion reactor for further reaction.

[0013] In the fluid catalytic conversion method of the present application, heavy feedstock oil and olefin-rich hydrocarbon oil feedstock are subjected to high-temperature cracking, respectively, and butylene and olefin-rich streams in the separated products are recycled to the reactor for further reaction. Meanwhile, catalytic gas oil with a high boiling point in the separated products can be subjected to hydrotreating and then subjected to further reaction. By refining the low-value olefins produced by chemical processing through a specific route, the yield of propylene can be effectively increased, the efficient use of petroleum resources can be realized, and the traditional method of preparing propylene by steam cracking, which involves high energy consumption, can be replaced. Meanwhile, the method of the present application also has the advantages of high ethylene yield, high selectivity, and low dry gas yield.

[0014] Other features and advantages of the present application are explained in detail in the following detailed description.

[0015] BRIEF DESCRIPTION OF THE DRAWINGS The drawings that form part of this specification are provided to aid in understanding the application and should not be considered limiting. The application may be read with reference to the drawings in combination with the following detailed description. In the drawings: FIG. 1 shows a schematic flow diagram of a preferred embodiment of the fluid catalytic conversion process of the present application; FIG. 2 shows a schematic flow diagram of another preferred embodiment of the fluid catalytic conversion process of the present application; and FIG. 3 shows a schematic flow diagram of yet another preferred embodiment of the fluid catalytic conversion process of the present application.

[0016] [Table 1]

[0017] Detailed Description of the Invention The present application is described in further detail below with reference to the drawings and specific embodiments thereof. It should be noted that the specific embodiments of the present application are provided for illustrative purposes only and are not intended to be limiting in any way.

[0018] Any specific numerical value, including the endpoints of a numerical range described in the context of this application, is not limited to that exact value. However, it should be interpreted as including all values close to the exact value, for example, all values within ±5% of the exact value. Furthermore, with respect to any numerical range described herein, any combination between the endpoints of the range, between each endpoint and any specific value within the range, or between any two specific values within the range, can be made to provide one or more new numerical ranges. Here, the new numerical ranges should also be considered to be specifically described in this application.

[0019] Unless otherwise specified, terms used herein have the same meaning as commonly understood by one of ordinary skill in the art; if a term is defined herein and that definition differs from the common understanding in the art, the definition provided herein shall control.

[0020] In this application, the term "C4+" means having at least 4 carbon atoms, for example, the term "C4+ olefin" refers to an olefin having at least 4 carbon atoms, while the term "C4+ fraction" refers to a fraction whose compounds have at least 4 carbon atoms. Correspondingly, the term "C5+" means having at least 5 C atoms.

[0021] In the context of this application, in addition to the explicitly mentioned subject matter, any unmentioned subject matter or matter shall be considered to be the same as that known in the art without any modifications. Furthermore, any of the embodiments described herein can be freely combined with one or more of the embodiments described herein. Any technical solution or idea obtained in this way shall be considered to be part of the original disclosure or original description of this application, and shall not be considered to be a new matter not disclosed or anticipated herein, unless it is obvious to a person skilled in the art that such a combination is obviously unreasonable.

[0022] All patent and non-patent literature cited herein, including but not limited to textbooks and journal articles, is hereby incorporated by reference in its entirety.

[0023] As described above, the present invention provides a fluid catalytic conversion process for maximizing propylene production, comprising the following steps 1) to 6): 1) introducing a heavy feedstock into a first reaction zone of a fluid catalytic conversion reactor and contacting the heavy feedstock with a catalytic conversion catalyst having a temperature of 650°C or greater for reaction; 2) introducing a hydrocarbon oil feedstock having an olefin content of 50 wt.% or more into a second reaction zone downstream of said fluid catalytic conversion reactor of the first reaction zone and contacting it with said catalytic conversion catalyst from said first reaction zone after the reaction of step 1) for reaction; 3) separating the effluent of the fluidized catalytic conversion reactor to obtain a reaction product and a used catalyst, and performing a first separation on the reaction product to obtain ethylene, propylene, butylene, a first catalytic cracking distillate, and a second catalytic cracking distillate; the first catalytic cracking distillate has an initial boiling point greater than 20°C, the second catalytic cracking distillate has an end boiling point less than 550°C, and the cut point between the first catalytic cracking distillate and the second catalytic cracking distillate is within the range of 140°C to 250°C; 4) subjecting the first catalytic cracking distillate to a second separation to obtain an olefin-rich stream having a C5+ olefin content of at least 50 wt%; 5) recycling at least a portion of the olefin-rich stream to step 2) for further reaction; and 6) recycling at least a portion of the butylenes separated in step 3) upstream of the fluidized catalytic conversion reactor where the heavy feedstock is introduced and contacted with the catalytic conversion catalyst for reaction.

[0024] In a preferred embodiment, the reaction in step 1) is carried out under first catalytic conversion conditions which may include a reaction temperature of 500-800°C, a reaction pressure of 0.05-1 MPa, a reaction time of 0.01-100 seconds, and a weight ratio of the catalytic conversion catalyst to the heavy feedstock oil of (1-200):1; and the reaction in step 2) is carried out under second catalytic conversion conditions which may include a reaction temperature of 400-680°C, a reaction pressure of 0.05-1 MPa, a reaction time of 0.01-100 seconds, and a weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock of (1-100):1.

[0025] In a more preferred embodiment, the first catalytic conversion conditions may include a reaction temperature of 510-780°C, a reaction pressure of 0.1-0.8 MPa, a reaction time of 0.1-80 seconds, and a weight ratio of the catalytic conversion catalyst to the heavy feedstock oil of (3-180):1; and the second catalytic conversion conditions may include a reaction temperature of 450-650°C, a reaction pressure of 0.1-0.8 MPa, a reaction time of 0.1-80 seconds, and a weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock of (3-70):1.

[0026] According to the present application, the heavy feedstock oil used in step 1) may be that commonly used in the art and is not strictly limited herein.For example, the heavy feedstock oil may be a petroleum hydrocarbon and / or a mineral oil; the petroleum hydrocarbon may be selected from the group consisting of vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum residue, atmospheric residue, heavy aromatic raffinate, or a combination thereof; the mineral oil may be selected from coal liquid oil, oil sand oil, shale oil, or a combination thereof.

[0027] In a preferred embodiment, the hydrocarbon oil feedstock used in step 2) may have an olefin content of 80 wt.% or more, preferably 90 wt.% or more; more preferably, the hydrocarbon oil feedstock may be a pure olefin feedstock.

[0028] According to the present application, the hydrocarbon oil feedstock can be obtained from various sources and is not particularly limited herein. In some embodiments, the olefins in the hydrocarbon oil feedstock can be obtained from a C4+ fraction produced by dehydrogenation of an alkane feedstock, a C4+ fraction produced by a catalytic cracking unit in an oil refinery, a C4+ fraction produced by a steam cracking unit in an ethylene plant, a C4+ olefin-rich by-product fraction from an MTO process, and a C4+ olefin-rich by-product fraction from an MTP process. In a preferred embodiment, the alkane feedstock for dehydrogenation can be selected from naphtha, aromatic raffinate, light hydrocarbons, or a combination thereof.

[0029] According to the present application, the dehydrogenation of an alkane feedstock can be carried out by contacting the alkane feedstock with a dehydrogenation catalyst. The dehydrogenation conditions used herein are a reactor inlet temperature of 400-700°C, a dehydrogenation time of 200-5000 h -1 and a reaction pressure of 0 to 1.0 MPa.

[0030] Preferably, the dehydrogenation catalyst comprises a support, an active component supported on the support, and a promoter; the support may be present in an amount of 60 to 90 wt %, the active component may be present in an amount of 8 to 35 wt %, and the promoter may be present in an amount of 0.1 to 5 wt %, based on the total weight of the dehydrogenation catalyst.

[0031] More preferably, the support may be alumina containing a modifier; the modifier may be present in an amount of 0.1 to 2 wt % based on the total weight of the dehydrogenation catalyst, and the modifier may be La and / or Ce; the active component may be platinum and / or chromium; the promoter may be a composition of bismuth and an alkali metal component, or a composition of bismuth and an alkaline earth metal component, wherein the molar ratio of bismuth to the active component is 1:(5 to 50); the molar ratio of bismuth to the alkali metal component is 1:(0.1 to 5); and the molar ratio of bismuth to the alkaline earth metal component is 1:(0.1 to 5). Particularly preferably, the alkali metal component may be one or more selected from Li, Na, and K; and the alkaline earth metal component may be one or more selected from Mg, Ca, and Ba.

[0032] According to the present application, butylene introduced into a fluidized catalytic conversion reactor for further reaction is contacted with a high-temperature catalytic conversion catalyst before the heavy feedstock. Because the difficulty of cracking hydrocarbons increases with the decrease in the number of carbon atoms and the energy required to crack butylene is relatively high, butylene is first contacted with a high-temperature catalytic conversion catalyst. As a result, the butylene conversion rate and the selectivity for ethylene and propylene products can be improved, the generation of by-products caused by the co-feeding of butylene and heavy feedstock can be avoided, and highly efficient resource utilization can be achieved.

[0033] In a preferred embodiment, the reaction in step 6) is carried out under third catalytic conversion conditions, which may include a reaction temperature of 650-800°C, a reaction pressure of 0.05-1 MPa, a reaction time of 0.01-10 seconds, and a weight ratio of catalytic conversion catalyst to butylene of (20-200):1. More preferably, the third catalytic conversion conditions include a reaction temperature of 680-780°C, a reaction pressure of 0.1-0.8 MPa, a reaction time of 0.05-8 seconds, and a weight ratio of catalytic conversion catalyst to butylene of (30-180):1.

[0034] In some preferred embodiments, the fluid catalytic conversion process of the present application further comprises the step of 2a) introducing the oxygen-containing organic compound into a second reaction zone of the fluid catalytic conversion reactor and contacting it with a catalytic conversion catalyst for reaction therein.

[0035] Preferably, the reaction of step 2a) is carried out under fourth catalytic conversion conditions, which may include a reaction temperature of 300-550°C, a reaction pressure of 0.05-1 MPa, a reaction time of 0.01-100 seconds, and a weight ratio of catalytic conversion catalyst to oxygen-containing organic compound feedstock of (1-100):1. More preferably, the fourth catalytic conversion conditions include a reaction temperature of 400-530°C, a reaction pressure of 0.1-0.8 MPa, a reaction time of 0.1-80 seconds, and a weight ratio of catalytic conversion catalyst to oxygen-containing organic compound feedstock of (3-80):1.

[0036] In such a preferred embodiment of the present application, the oxygen-containing organic compound may be fed alone or mixed with other feedstocks. For example, the oxygen-containing organic compound may be mixed with the hydrocarbon oil feedstock and then fed to the second reaction zone of the fluid catalytic conversion reactor, or the oxygen-containing organic compound may be fed to the second reaction zone downstream of the fluid catalytic conversion reactor where the hydrocarbon oil feedstock is introduced.

[0037] Particularly preferably, the oxygen-containing organic compound comprises at least one of methanol, ethanol, dimethyl ether, methyl ethyl ether, and ethyl ether. For example, the oxygen-containing organic compound, such as methanol or dimethyl ether, can be obtained from coal-based or natural gas-based synthesis gas.

[0038] In some preferred embodiments, the fluid catalytic conversion method of the present application further includes the step of: 7) hydrotreating the second catalytic cracking distillate obtained in step 3) to obtain a hydrogenated catalytic cracking distillate, and recycling the hydrogenated catalytic cracking distillate to the first reaction zone of the fluid catalytic conversion reactor for further reaction. In this embodiment, the second catalytic cracking distillate is subjected to hydrotreating and then to further reaction. As a result, side reactions that produce low-molecular-weight alkanes and coke can be further reduced, the yields of ethylene and propylene can be improved, and carbon atoms can be effectively utilized.

[0039] Preferably, the hydrotreating conditions are a hydrogen partial pressure of 3.0 to 20.0 MPa, a reaction temperature of 300 to 450°C, a volume ratio of hydrogen to oil of 300 to 2000, and a hydrotreating time of 0.1 to 3.0 h. -1 The volumetric space velocity may include

[0040] According to the present application, the hydrogenation catalyst used in step 7) may be one commonly used in the art and is not particularly limited herein. For example, the hydrogenation catalyst may include a support, a metal component, and an additive, optionally supported on the support. Preferably, the hydrogenation catalyst includes 20 to 90 wt. % of the support, 10 to 80 wt. % of the supported metal, and 0 to 10 wt. % of the additive, based on the total weight of the hydrogenation catalyst. More preferably, the support is alumina and / or amorphous silica-alumina, the metal component is a Group VIB metal and / or a Group VIII metal, and the additive is at least one selected from fluorine, phosphorus, titanium, and platinum; and even more preferably, the Group VIB metal is Mo and / or W, and the Group VIII metal is Co and / or Ni. Particularly preferably, the additive is present in an amount of from 0 to 10% by weight, the Group VIB metal is present in an amount of from 12 to 39% by weight, and the Group VIII metal is present in an amount of from 1 to 9% by weight, based on the total weight of the hydrogenation catalyst.

[0041] In a preferred embodiment, the fluid catalytic conversion process of the present application further comprises the step of 8) regenerating the spent catalyst obtained by separation in step 3) by burning coke to obtain a regenerated catalyst having a temperature of 650°C or higher, and thereafter recycling the regenerated catalyst upstream of the first reaction zone of the fluid catalytic conversion reactor for use as a catalytic conversion catalyst.

[0042] In a preferred embodiment, the catalytic conversion catalyst used herein may comprise, based on the weight of the catalyst, from 1 wt. % to 50 wt. % molecular sieve, from 5 wt. % to 99 wt. % inorganic oxide, and from 0 wt. % to 70 wt. % clay.

[0043] In a further preferred embodiment, a molecular sieve functions as an active component in the catalytic conversion catalyst, and the molecular sieve may include at least one of a macroporous molecular sieve, a mesoporous molecular sieve, and a microporous molecular sieve.

[0044] In some further preferred embodiments, the mesoporous molecular sieve may be a ZSM molecular sieve, for example, the ZSM molecular sieve may be at least one selected from the group consisting of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, and ZSM-48; the microporous molecular sieve may be a SAPO molecular sieve and / or a SSZ molecular sieve, for example, the SAPO molecular sieve may be SAPO-34, SAPO- The SSZ molecular sieve may be at least one selected from the group consisting of SSZ-13, SSZ-39, and SSZ-62; the macroporous molecular sieve may be selected from REY molecular sieve, REHY molecular sieve, ultrastable Y molecular sieve, high silica Y molecular sieve, beta molecular sieve, and other molecular sieves of similar structure, or mixtures thereof.

[0045] In a particularly preferred embodiment, the molecular sieve comprises from 40 to 100% by weight, preferably from 50 to 100% by weight, of mesoporous molecular sieves, from 0 to 30% by weight, preferably from 0 to 25% by weight, of microporous molecular sieves, and from 0 to 30% by weight, preferably from 0 to 25% by weight, of macroporous molecular sieves, based on the total weight of the molecular sieves.

[0046] In a further preferred embodiment, an inorganic oxide functions as a binder in the catalytic conversion catalyst, and preferably, the inorganic oxide may be selected from silica (SiO2) and / or alumina (Al2O3).

[0047] In a further preferred embodiment, clay functions as a matrix in the catalytic conversion catalyst, and preferably, the clay may be selected from kaolin and / or halloysite.

[0048] In a more preferred embodiment, the catalytic conversion catalyst may also contain a modifying element to further improve the performance of the catalytic conversion catalyst. The modifying element may be at least one selected from a nonmetallic element, a transition metal element, and a rare earth metal element; even more preferably, the nonmetallic element may be phosphorus, and the transition metal element may be selected from iron, cobalt, and nickel, and the content of the modifying element is 0.1 to 3 wt% of the catalytic conversion catalyst.

[0049] In a preferred embodiment, the olefin-rich stream separated in step 4) has an olefin content of 80 wt% or more, more preferably a C5+ olefin content of 80 wt% or more. The higher the olefin content in the olefin-rich stream, the better the purification effect and the better the resource utilization.

[0050] In some preferred embodiments, the second separation in step 4) further comprises: separating an olefin-depleted stream, a first olefin-rich stream having a lower boiling point, and a second olefin-rich stream having a higher boiling point from the first catalytic cracking distillate; the cut point between the first stream and the second stream is in the range of 140 to 200°C; and step 5) further comprises introducing the first olefin-rich stream into a second reaction zone for further reaction, and introducing the second olefin-rich stream into a third reaction zone downstream of the fluid catalytic conversion reactor of the second reaction zone for further reaction.

[0051] After numerous experiments, the inventors of the present application found that olefins with long carbon chains have a poorer ability to inhibit methane formation while producing propylene through cracking than olefins with short carbon chains. Taking propylene as the target product for illustration, the longer the carbon chain of the olefin molecule subjected to catalytic cracking, the easier it is to crack. To avoid an increase in by-products (such as methane in the product caused by the first cracking of the long carbon chains into small molecules), the olefins with long carbon chains may be cracked in a fourth reaction zone under relatively mild conditions to obtain C5-C9 olefins with short carbon chains, which are then recycled to the reactor for further cracking. It is beneficial to improve the yield of propylene and reduce the yield of methane. In the preferred embodiment described above, a first stream and a second stream containing C5+ olefins with different distillation ranges are introduced into different reaction zones. Specifically, the first stream with a lower boiling point is introduced into the third reaction zone, and the second stream with a higher boiling point is introduced into the fourth reaction zone. As a result, olefins with longer carbon chains can be prevented from being cracked into smaller molecules all at once, improving the yield of propylene.

[0052] In a further preferred embodiment, the conditions for the further reaction of the first stream introduced into the third reaction zone may include a reaction temperature of 600 to 750°C, a reaction pressure of 0.05 to 1 MPa, a reaction time of 0.01 to 100 seconds, and a weight ratio of the catalytic conversion catalyst to the first stream of (1 to 140):1; and the conditions for the further reaction of the second stream introduced into the fourth reaction zone may include a reaction temperature of 400 to 650°C, a reaction pressure of 0.05 to 1 MPa, a reaction time of 0.01 to 100 seconds, and a weight ratio of the catalytic conversion catalyst to the second stream of (1 to 100):1.

[0053] In a preferred embodiment, the fluidized catalytic conversion reactor may be selected from a riser reactor, which may be an isodiameter riser reactor or a diameter conversion riser reactor, a fluidized bed reactor, which may be an isolinear velocity fluidized bed reactor or an isodiameter fluidized bed reactor, an up-transfer line, a down-transfer line, or a combination of two or more thereof, and the diameter conversion riser reactor may be, for example, a riser reactor as described in Chinese Patent CN1078094C.

[0054] In a preferred embodiment, as shown in FIG. 1, the fluid catalytic conversion process of the present application is carried out as follows: The prelift medium is introduced through pipeline 101 from the bottom of the fluidized catalytic conversion reactor (riser reactor) 102, and the regenerated catalytic conversion catalyst from pipeline 117 moves upward along the fluidized catalytic conversion reactor 102 under the lifting action of the prelift medium, and the heavy feedstock oil is injected through pipeline 103 together with atomized steam from pipeline 104 into the bottom of the first reaction zone I of the reactor 102, where it comes into contact with and reacts with a high-temperature catalyst having a temperature of 650°C or higher, and then moves further upward.

[0055] A hydrocarbon oil feedstock having an olefin content of 50% by weight or more is injected into the lower middle section of the fluid catalytic conversion reactor 102 through a pipeline 105 together with atomized steam from a pipeline 106, and mixed with the stream from the first reaction zone I in a second reaction zone II, where the hydrocarbon oil feedstock contacts and reacts with the high-temperature catalyst and moves upward.

[0056] The resulting reaction products and deactivated spent catalyst are sent through a discharge section 107 to a cyclone separator 108 in the separator, where the spent catalyst and reaction products are separated. The reaction products are sent to a plenum chamber 109, and the fine catalyst powder is returned to the separator through a dipleg. The spent catalyst in the separator is sent to a stripping section 110, where it is contacted with stripping steam from a pipeline 111. The stripped product vapor from the spent catalyst passes through the cyclone separator and is then sent to the plenum chamber 109. The stripped spent catalyst is sent to a regenerator 113 through a standpipe 112, and main air is introduced into the regenerator through a pipeline 116 to burn the coke on the spent catalyst and regenerate the deactivated spent catalyst. The exhaust gas is sent to an exhaust gas turbine through a pipeline 115. The regenerated catalyst is sent to the reactor 102 through a pipeline 117.

[0057] The reaction product (reaction product vapor) is sent to the subsequent fractionator 120 through the reactor vapor line 119, the separated hydrogen, methane and ethane are withdrawn through the pipeline 121, ethylene is withdrawn through the pipeline 122, propylene is withdrawn through the pipeline 123, butylene is recycled to the bottom of the reactor 102 through the pipeline 124 for further reaction, propane and butane are withdrawn through the pipeline 125, and the first catalytic cracking distillate is sent to the olefin separator 128 through the pipeline 126 to separate the olefins. The olefin-depleted stream is withdrawn through pipeline 129, the olefin-rich stream is sent to the bottom of the second reaction zone II of the reactor 102 through pipeline 130 for further reaction, the second catalytic cracking distillate is sent to the hydrotreater 131 through pipeline 127, light components and hydrogenated catalytic cracking distillate are obtained after hydrotreating, the light components are withdrawn through pipeline 118, and the hydrogenated catalytic cracking distillate is withdrawn through pipeline 114 and optionally introduced into the first reaction zone I of the reactor 102 for further reaction.

[0058] FIG. 2 shows another preferred embodiment of the present application, in which the fluidized catalytic conversion reactor further comprises a third reaction zone downstream of the second reaction zone II, and the third reaction zone is appropriately enlarged when the amount of olefin-rich feedstock from an external source is relatively large.

[0059] In this embodiment, as shown in FIG. 2, the fluid catalytic conversion process of the present application is carried out as follows: The prelift medium is introduced from the first reaction zone I of the fluid catalytic conversion reactor 202 through pipeline 201, the regenerated catalytic conversion catalyst from pipeline 217 moves upward along the reactor 202 under the lifting action of the prelift medium, and the heavy feedstock oil together with the atomized steam from pipeline 204 is injected through pipeline 203 into the bottom of the first reaction zone I of the reactor 202, where it comes into contact with and reacts with the high-temperature catalyst having a temperature of 650°C or higher and moves upward.

[0060] A hydrocarbon oil feedstock having an olefin content of 50% by weight or more is injected into the lower middle section of the fluid catalytic conversion reactor 202 through a pipeline 205 together with atomized steam from a pipeline 206, and mixed with the stream from the first reaction zone I in the second reaction zone II, and the hydrocarbon oil feedstock contacts and reacts with the high-temperature catalyst and moves upward.

[0061] The reaction products and deactivated spent catalyst produced in the reactor 202 are sent through a discharge section 207 to a cyclone separator 208 in the separator, where they are separated from the spent catalyst. The reaction products are sent to a plenum chamber 209, and the fine catalyst powder is returned to the separator through a dipleg. The spent catalyst in the separator is sent to a stripping section 210, where it is contacted with stripping steam from a pipeline 211. The stripped product steam from the spent catalyst passes through the cyclone separator and is then sent to the plenum chamber 209. The stripped spent catalyst is sent to a regenerator 213 through a standpipe 212, and main air is introduced into the regenerator through a pipeline 216 to burn the coke on the spent catalyst and regenerate the deactivated spent catalyst. The exhaust gas is sent to an exhaust gas turbine through a pipeline 215. The regenerated catalyst is recycled to the bottom of the reactor 202 through a pipeline 217. The reaction product is sent to the subsequent fractionator 220 through the reactor vapor line 219, the separated hydrogen, methane and ethane are withdrawn through the pipeline 221, ethylene is withdrawn through the pipeline 222, propylene is withdrawn through the pipeline 223, butylene is recycled to the bottom of the reactor 202 through the pipeline 224 for further reaction, propane and butane are withdrawn through the pipeline 225, the first catalytic cracking distillate is sent to the olefin separator 228 through the pipeline 226, an olefin-depleted stream, a first olefin-rich stream and a second olefin-rich stream are obtained by separation, wherein the cut point between the first stream and the second stream is in the range of 140-200°C. The olefin-depleted stream is withdrawn via pipeline 229, a first olefin-rich stream having a lower boiling point is introduced via pipeline 230 into a second reaction zone II of reactor 202 for further reaction, and a second olefin-rich stream having a higher boiling point is introduced via pipeline 231 into a third reaction zone III of reactor 202 for further reaction.The second catalytic cracking distillate is introduced into the hydrotreater 232 through the pipeline 227, and light components and hydrogenated catalytic cracking distillate are obtained after hydrotreatment, in which the light components are withdrawn through the pipeline 218 and the hydrogenated catalytic cracking distillate is withdrawn through the pipeline 214 and optionally sent to the bottom of the first reaction zone I of the reactor 202 for further reaction.

[0062] In yet another preferred embodiment, as shown in FIG. 3, the fluid catalytic conversion process of the present application is carried out as follows: The prelift medium is introduced from the bottom of the fluid catalytic conversion reactor (riser reactor) 302 through pipeline 301, the regenerated catalytic conversion catalyst from pipeline 317 moves upward along the fluid catalytic conversion reactor 302 under the lifting action of the prelift medium, and the heavy feedstock oil together with atomized steam from pipeline 304 is injected through pipeline 303 into the bottom of the first reaction zone I of the fluid catalytic conversion reactor 302, where it comes into contact with the high-temperature catalyst having a temperature of 650°C or higher for reaction and reacts, and moves further upward.

[0063] A hydrocarbon oil feedstock having an olefin content of more than 50 wt% is injected into the lower middle part of the fluid catalytic conversion reactor 302 through a pipeline 305 together with atomized steam from a pipeline 306, and mixed with the stream from the first reaction zone I in a second reaction zone II, and the hydrocarbon oil feedstock contacts and reacts with the high-temperature catalyst and moves upward.

[0064] An oxygen-containing organic compound (e.g., methanol) is injected into the lower middle section of the second reaction zone II through pipeline 307 downstream from where the hydrocarbon oil feedstock is injected and mixed with the stream therein. The oxygen-containing organic compound reacts with the catalyst and travels upward.

[0065] The resulting reaction product and deactivated spent catalyst are sent through a discharge section 308 to a cyclone separator 309 in the separator, where the spent catalyst and reaction product are separated. The reaction product is sent to a plenum chamber 310, and the fine catalyst powder is returned to the separator through a dipleg. The spent catalyst in the separator is sent to a stripping section 311, where it is contacted with stripping steam from a pipeline 312. The product steam stripped from the spent catalyst passes through the cyclone separator and is then sent to the plenum chamber 310. The stripped spent catalyst is sent to a regenerator 314 through a standpipe 313, and main air is introduced into the regenerator through a pipeline 316 to burn the coke on the spent catalyst and regenerate the deactivated spent catalyst. The exhaust gas is sent to an exhaust gas turbine through a pipeline 315. The regenerated catalyst is recycled to the bottom of the reactor 302 through a pipeline 317.

[0066] The reaction product (reaction product vapor) is sent through reactor vapor line 319 to the subsequent fractionator 320 for separation, the separated hydrogen, methane, and ethane are withdrawn through pipeline 321, ethylene is withdrawn through pipeline 322, propylene is withdrawn through pipeline 323, butylene is recycled to the bottom of reactor 302 through pipeline 324 for further reaction, propane and butane are withdrawn through pipeline 325, and unconverted oxygen-containing organic compounds are withdrawn through pipeline 326 and, optionally, recycled to the lower middle section of the second reaction zone II of reactor 302 for further reaction; the first catalytic cracking distillate is sent through pipeline 327 to olefin separator 329 for separation. As a result, an olefin-depleted stream and an olefin-rich stream are obtained, the olefin-depleted stream is withdrawn through pipeline 331, and the olefin-rich stream is sent to the bottom of the second reaction zone II of the reactor 302 through pipeline 330 for further reaction; the second catalytic cracking distillate is sent to the hydrotreater 332 through pipeline 328, and light components and hydrogenated catalytic cracking distillate are obtained after separation, the light components are withdrawn through pipe 318, and the hydrogenated catalytic cracking distillate is withdrawn through pipeline 333, optionally mixed with heavy feedstock, and then sent to the bottom of the first reaction zone I of the reactor 302 for further reaction.

[0067] In a particularly preferred embodiment, the present application provides the following technical solutions: A1, a catalytic conversion method for maximizing propylene production, comprising the following steps S1 to S3: S1, contacting a heavy feedstock with a catalytic conversion catalyst having a temperature of 650°C or higher to carry out a first catalytic conversion reaction in a first reaction zone of a catalytic conversion reactor to obtain a first mixed stream; S2, contacting a hydrocarbon oil feedstock having an olefin content of 50 wt% or more with the first mixed flow in a second reaction zone of the catalytic conversion reactor, and carrying out a second catalytic conversion reaction to obtain a reaction product vapor and a spent catalyst; the second reaction zone is located downstream of the first reaction zone; S3, performing a first separation on the reaction product vapor to obtain ethylene, propylene, butylene, a first catalytic cracking distillate, and a second catalytic cracking distillate; the first catalytic cracking distillate has a first temperature higher than 20°C, the second catalytic cracking distillate has a last temperature lower than 550°C, and the cut point between the first catalytic cracking distillate and the second catalytic cracking distillate is between 140°C and 250°C, and performing a second separation on the first catalytic cracking distillate to obtain an olefin-rich stream; and separately introducing the butylene and the olefin-rich stream into the catalytic conversion reactor for further reaction.

[0068] A2. The method according to item A1, wherein in step S3, the butylene introduced into the catalytic conversion reactor for further reaction is contacted with the catalytic conversion catalyst before the heavy feedstock.

[0069] A3. The method according to item A1, wherein the olefins in the olefin-rich stream are C4+ olefins; The olefin content of the olefin-rich stream is from 50 wt% to 100 wt%.

[0070] A4. The method according to item A1, wherein the butylene is introduced into the first reaction zone of the catalytic conversion reactor for reaction, and the olefin-rich stream is introduced into the second reaction zone of the catalytic conversion reactor for reaction.

[0071] A5, the catalytic conversion reactor further comprises a reaction zone A and a reaction zone B; the reaction zone A and the reaction zone B are sequentially arranged downstream of the second reaction zone; the second separation includes separating a first olefin-rich stream and a second olefin-rich stream from the first catalytic cracking distillate; a cut point between the first stream and the second stream is between 140°C and 200°C; the butylene is introduced into the first reaction zone for reaction; the first stream is introduced into reaction zone A for further reaction; The method according to item A1, wherein the second stream is introduced into reaction zone B for further reaction.

[0072] A6, the method includes the step of regenerating the spent catalyst by burning coke to obtain a regenerated catalyst; and The method according to claim A1, further comprising preheating the regenerated catalyst and then recycling it to the catalytic conversion reactor.

[0073] A7. The method according to item A1, further comprising: hydrotreating the second catalytic cracked distillate to obtain a hydrogenated product, and separating the hydrogenated catalytic cracked distillate from the hydrogenated product; introducing the hydrotreated catalytic cracking distillate into the first reaction zone for further reaction.

[0074] A8, the conditions of the hydrotreating are a hydrogen partial pressure of 3.0 to 20.0 MPa, a reaction temperature of 300 to 450 ° C, a volume ratio of hydrogen to oil of 300 to 2000, and a hydrotreating time of 0.1 to 3.0 h -1 The method according to item A7, comprising a volumetric space velocity of

[0075] A9. The method according to item A1, wherein the catalytic conversion reactor is selected from a riser reactor, a constant linear velocity fluidized bed, a constant diameter fluidized bed, an upward transfer line, and a downward transfer line, or a combination of two thereof connected in series; the riser reactor is preferably a diameter conversion riser reactor.

[0076] A10, the conditions of the first catalytic conversion reaction include a reaction temperature of 500 to 800°C, a reaction pressure of 0.05 to 1 MPa, a reaction time of 0.01 to 100 seconds, and a weight ratio of the catalytic conversion catalyst to the heavy feedstock oil of (1 to 200):1; The conditions of the second catalytic conversion reaction include a reaction temperature of 400 to 680°C, a reaction pressure of 0.05 to 1 MPa, a reaction time of 0.01 to 100 seconds, and a weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock of (1 to 100):1; Preferably, the conditions of the first catalytic conversion reaction include a reaction temperature of 510 to 780°C, a reaction pressure of 0.1 to 0.8 MPa, a reaction time of 0.1 to 80 seconds, and a weight ratio of the catalytic conversion catalyst to the heavy feedstock oil of (3 to 180); The method according to Item A1, wherein the conditions of the second catalytic conversion reaction include a reaction temperature of 450 to 650°C, a reaction pressure of 0.1 to 0.8 MPa, a reaction time of 0.1 to 80 seconds, and a weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock of (3-70):1.

[0077] A11, the conditions for the further reaction of the butylene introduced into the catalytic reactor include a reaction temperature of 650-800°C, a reaction pressure of 0.05-1 MPa, a reaction time of 0.01-10 seconds, and a weight ratio of the catalytic conversion catalyst to the butylene of (20-200):1; Preferably, the conditions include a reaction temperature of 680 to 780°C, a reaction pressure of 0.1 to 0.8 MPa, a reaction time of 0.05 to 8 seconds, and a weight ratio of the catalytic conversion catalyst to the butylene of (30 to 180):1.

[0078] A12, the hydrocarbon oil feedstock has an olefin content of 80 wt% or more, preferably 90 wt% or more, and more preferably, the hydrocarbon oil feedstock is a pure olefin feedstock; The method according to item A1, wherein the heavy feedstock is a petroleum hydrocarbon and / or a mineral oil; the petroleum hydrocarbon is at least one selected from the group consisting of vacuum gas oil, atmospheric gas oil, coked gas oil, deasphalted oil, vacuum resid, atmospheric resid, and heavy aromatic raffinate; and the mineral oil is at least one selected from the group consisting of coal liquid oil, oil sands oil, and shale oil.

[0079] A13, the olefins in the hydrocarbon oil feedstock are obtained from a C4+ fraction produced by dehydrogenation of an alkane feedstock, a C4+ fraction from a catalytic cracking unit in an oil refinery, a C4+ fraction from a steam cracking unit in an ethylene plant, a C4+ olefin-rich by-product fraction from an MTO process, and a C4+ olefin-rich by-product fraction from an MTP process; The method according to item A1 or A12, wherein the alkane feedstock is at least one selected from the group consisting of naphtha, aromatic raffinate, and light hydrocarbons.

[0080] A14, the catalytic conversion catalyst comprises, based on the weight of the catalytic conversion catalyst, 1 to 50 wt% of a molecular sieve, 5 to 99 wt% of an inorganic oxide, and 0 to 70 wt% of a clay; The molecular sieve comprises one or more of a macroporous molecular sieve, a mesoporous molecular sieve, and a microporous molecular sieve; The method according to Item A1, wherein the catalytic conversion catalyst further contains 0.1 to 3 wt. % of an active metal based on the weight of the catalytic conversion catalyst; and the active metal is one or more selected from the group consisting of Group VIII metals, Group IVA metals, and rare earth metals.

[0081] B1. A catalytic conversion method for maximizing propylene production, comprising the following steps S1 to S4: S1, contacting a heavy feedstock with a catalytic conversion catalyst having a temperature of 650°C or higher, and carrying out a first catalytic conversion reaction in a first reaction zone of a catalytic conversion reactor under first catalytic conversion conditions to obtain a first mixed stream; S2, contacting a hydrocarbon oil feedstock having an olefin content of 50 wt% or more and an oxygen-containing organic compound feedstock with the first mixed stream from the first reaction zone in a second reaction zone of the catalytic conversion reactor, and conducting a second catalytic conversion reaction under second catalytic conversion conditions to obtain a reaction product vapor and a spent catalyst; S3, performing a first separation on the reaction product vapor to obtain ethylene, propylene, butylene, oxygen-containing organic compounds, a first catalytic cracking distillate, and a second catalytic cracking distillate; the initial boiling point of the first catalytic cracking distillate is greater than 20°C and less than 140°C, the final boiling point of the second catalytic cracking distillate is greater than 250°C and less than 550°C, and the cut point between the first catalytic cracking distillate and the second catalytic cracking distillate is 140°C to 250°C; performing a second separation on the first catalytic cracking distillate to obtain an olefin-rich stream; S4, recycling the olefin-rich stream to the catalytic conversion reactor for further reaction.

[0082] B2. The method according to item B1, wherein the method comprises the steps of: passing the reaction product vapor through a product fractionator for a first separation to separate ethylene, propylene, butylene, the oxygen-containing organic compounds, the first catalytic cracking distillate, and the second catalytic cracking distillate; passing the first catalytic cracking distillate through an olefin separator for second separation to separate the olefin-rich stream; recycling the olefin-rich stream to the first reaction zone of the catalytic conversion reactor for further reaction.

[0083] B3. The method according to item B1 or B2, wherein the method further comprises the following steps: recycling the separated butylenes to the first reaction zone of the catalytic conversion reactor for further reaction; Preferably, the butylenes that are recycled to the catalytic conversion reactor for further reaction are contacted with the catalytic conversion catalyst before the heavy feedstock.

[0084] B4, the conditions for the further reaction of the butylene recycled to the catalytic reactor include a reaction temperature of 650-800°C, a reaction pressure of 0.05-1 MPa, a reaction time of 0.01-10 seconds, and a weight ratio of the catalytic conversion catalyst to the recycled butylene of (20-200):1; Preferably, the conditions include a reaction temperature of 680 to 780°C, a reaction pressure of 0.1 to 0.8 MPa, a reaction time of 0.05 to 8 seconds, and a weight ratio of the catalytic conversion catalyst to the recycled butylene of (30 to 180):1.

[0085] B5, the first catalytic conversion conditions include a reaction temperature of 500-800°C, a reaction pressure of 0.05-1 MPa, a reaction time of 0.01-100 seconds, and a weight ratio of the catalytic conversion catalyst to the heavy feedstock of (1-200):1; Preferably, the first catalytic conversion conditions include a reaction temperature of 510 to 780°C, a reaction pressure of 0.1 to 0.8 MPa, a reaction time of 0.1 to 80 seconds, and a weight ratio of the catalytic conversion catalyst to the heavy feedstock of (3 to 180):1.

[0086] B6, the second catalytic conversion conditions include a reaction temperature of 300 to 680°C, a reaction pressure of 0.01 to 1 MPa, a reaction time of 0.01 to 100 seconds, and a weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock of (1 to 100):1; and a weight ratio of the catalytic conversion catalyst to the oxygen-containing organic compound of (1 to 100):1; Preferably, the second catalytic conversion conditions include a reaction temperature of 400 to 650°C, a reaction pressure of 0.05 to 1 MPa, a reaction time of 0.1 to 80 seconds, and a weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock of (3 to 70):1; and a weight ratio of the catalytic conversion catalyst to the oxygen-containing organic compound of (3 to 50):1; Preferably, the reaction temperature of the first catalytic conversion reaction is 0 to 380°C higher than the reaction temperature of the second catalytic conversion reaction, and more preferably, the reaction temperature of the first catalytic conversion reaction is 10 to 370°C higher than the reaction temperature of the second catalytic conversion reaction. The method according to Item B1 or B5.

[0087] B7. The method according to item B1 or B5, wherein the second reaction zone is divided into an upstream portion of the second reaction zone and a downstream portion of the second reaction zone along the flow direction of the reaction stream, and is bounded by the supply position of the oxygen-containing organic compound feedstock, and the downstream portion of the second reaction zone is downstream of the supply position of the oxygen-containing organic compound feedstock; the method further comprises the following steps: contacting the first mixed stream from the first reaction zone with the hydrocarbon oil feedstock having an olefin content of 50 wt.% or more in the upstream portion of the second reaction zone to carry out a catalytic conversion reaction to obtain a second mixed stream; and then contacting the second mixed stream with the oxygen-containing organic compound feedstock in the downstream portion of the second reaction zone to carry out a catalytic conversion reaction to obtain a reaction product vapor and spent catalyst.

[0088] B8, the catalytic conversion conditions for the reaction between the hydrocarbon oil feedstock and the first mixed stream in the upstream portion of the second reaction zone include: a reaction temperature of 400 to 680°C, preferably 450 to 650°C; a reaction pressure of 0.05 to 1 MPa, preferably 0.1 to 0.8 MPa; a reaction time of 0.01 to 100 seconds, preferably 0.1 to 80 seconds; a weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock of (1-100):1, preferably (3-70):1; Catalytic conversion conditions for the reaction between the oxygen-containing organic compound feedstock and the second mixed stream in the downstream portion of the second reaction zone include: a reaction temperature of 300 to 550°C, preferably 400 to 530°C; a reaction pressure of 0.01 to 1 MPa, preferably 0.05 to 1 MPa; a reaction time of 0.01 to 100 seconds, preferably 0.1 to 80 seconds; a weight ratio of the catalytic conversion catalyst to the oxygen-containing organic compound feedstock of (1-100):1, preferably (3-50):1; The reaction temperature in the upstream portion of the second reaction zone is 0 to 250°C higher than the reaction temperature in the downstream portion of the second reaction zone, and preferably the reaction temperature in the upstream portion of the second reaction zone is 10 to 240°C higher than the reaction temperature in the downstream portion of the second reaction zone.

[0089] B9. The method of claim B1, wherein the method further comprises recycling the separated oxygen-containing organic compound to the second reaction zone of the catalytic conversion reactor for further reaction.

[0090] B10. The method of any one of paragraphs B1 to B9, wherein the catalytic conversion reactor is a riser reactor, preferably a diameter conversion riser reactor.

[0091] B11. The method according to item B1, further comprising the steps of: regenerating the spent catalyst by burning coke to obtain a regenerated catalyst; and recycling the regenerated catalyst to the first reaction zone of the catalytic conversion reactor as the catalytic conversion catalyst.

[0092] B12, the heavy feedstock is a petroleum hydrocarbon and / or a mineral oil; the petroleum hydrocarbon is at least one selected from vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum resid, atmospheric resid, and heavy aromatic raffinate; the mineral oil is at least one selected from the group consisting of coal liquid oil, oil sands oil, and shale oil; the hydrocarbon oil feedstock has an olefin content of 80 wt% or more, preferably 90 wt% or more, more preferably the hydrocarbon oil feedstock is a pure olefin feedstock; Optionally, the oxygen-containing organic compound feedstock comprises at least one of methanol, ethanol, dimethyl ether, methyl ethyl ether, and diethyl ether.

[0093] B13, the olefins in the hydrocarbon oil feedstock are obtained from at least one of a C5+ fraction produced by an alkane dehydrogenation unit, a C5+ fraction produced by a catalytic cracking unit in an oil refinery, a C5+ fraction produced by a steam cracking unit in an ethylene plant, a C5+ olefin-rich by-product fraction of an MTO process, and a C5+ olefin-rich by-product fraction of an MTP process; Optionally, the alkane feedstock of the alkane dehydrogenation unit is obtained from at least one of naphtha, aromatic raffinate, and other light hydrocarbons.

[0094] B14. The catalytic conversion catalyst comprises, based on the weight of the catalytic conversion catalyst, 1 to 50 wt. % of a molecular sieve, 5 to 99 wt. % of an inorganic oxide, and 0 to 70 wt. % of a clay; The molecular sieve comprises one or more of a macroporous molecular sieve, a mesoporous molecular sieve, and a microporous molecular sieve; The method according to Item B1, wherein the catalytic conversion catalyst further contains 0.1 to 3 wt. % of an active metal based on the weight of the catalytic conversion catalyst; and the active metal is one or more selected from the group consisting of Group VIII metals, Group IVA metals, and rare earth metals.

[0095] B15. The method according to item B1, further comprising the steps of: hydrotreating the second catalytic cracked distillate under hydrogenating conditions to obtain a hydrogenated catalytic cracked distillate; introducing the hydrotreated catalytic cracking distillate into a first reaction zone of the catalytic conversion reactor for further reaction; Here, the hydrotreating conditions are a hydrogen partial pressure of 3.0 to 20.0 MPa, a reaction temperature of 300 to 450°C, a volume ratio of hydrogen to oil of 300 to 2000, and a reaction time of 0.1 to 3.0 h. -1 Including volumetric space velocity of; Optionally, the hydrogenation catalyst comprises, based on the total weight of the hydrogenation catalyst, 20 to 90 wt % of a support, 10 to 80 wt % of a supported metal, and 0 to 10 wt % of an additive; the support is alumina and / or amorphous silica-alumina, the additive is at least one selected from the group consisting of fluorine, phosphorus, titanium, and platinum, and the supported metal is a Group VIB metal and / or a Group VIII metal; Preferably, the Group VIB metal is Mo or / and W, and the Group VIII metal is Co or / and Ni.

[0096] B16, wherein the olefins in the olefin-rich stream are C5+ olefins; The process according to item B1, wherein the olefin-rich stream has a C5+ olefin content of 50 wt% or more, preferably 80 wt% or more.

[0097] [Example] The present invention will be described in more detail below with reference to examples. All raw materials used in the examples are commercially available.

[0098] Raw materials and catalysts The properties of heavy feedstock a and heavy feedstock b used in the following examples are shown in Tables 1-1 and 1-2, respectively.

[0099] [Table 2]

[0100] [Table 3]

[0101] The preparation or source of the various catalysts used in the following examples and comparative examples is as follows: Catalytic conversion catalyst A 969 g of halloysite (China Kaolin Clay Co., Ltd., solids content 73%) was slurried in 4300 g of deionized water, and 781 g of pseudoboehmite (CHALCO Shandong Co., Ltd., solids content 64%) and 144 ml of hydrochloric acid (30% concentration, specific gravity 1.56) were added. The mixture was stirred evenly and aged at 60 °C for 1 hour while maintaining the pH at 2-4. The mixture was cooled to room temperature and 5000 g of the prepared slurry was added. The slurry contained 1600 g of mesoporous ZSM-5 zeolite and macroporous Y molecular sieve (Sinopec Catalyst Co., Ltd., Qilu Branch). The weight ratio of mesoporous ZSM-5 zeolite to macroporous Y molecular sieve was 9:1. The mixture was stirred evenly, spray-dried, and washed to remove free Na+ to obtain the catalyst. The obtained catalyst was aged at 800°C in 100% steam, and the aged catalyst was designated as catalyst A. The properties of catalyst A are shown in Table 2.

[0102] Catalytic conversion catalyst B: an industrial product available from Sinopec Catalyst Co., Ltd., Qilu Branch, under the trade name CEP-1, the properties of which are shown in Table 2.

[0103] Catalytic conversion catalyst C: an industrial product available from Sinopec Catalyst Co., Ltd., Qilu Branch, under the trade name CHP-1, the properties of which are shown in Table 2.

[0104] Hydrotreating catalyst D Ammonium metatungstate ((NH4)2W4O 13Nickel nitrate (Ni(NO)·18H2O, chemically pure) and nickel nitrate (Ni(NO)·18H2O, chemically pure) were weighed and mixed with water to obtain 200 ml of solution. The solution was added to 50 g of alumina support and impregnated at room temperature for 3 hours. During the impregnation, the impregnated solution was treated with ultrasound for 30 minutes, cooled, filtered, and dried in a microwave oven for approximately 15 minutes. The catalyst contained the following components: 30.0 wt. % WO3, 3.1 wt. % NiO, and the remainder was alumina, designated Catalyst D.

[0105] Hydrodesulfurization catalyst E 1000 g of pseudoboehmite manufactured by the ChangLing Branch of Sinopec Catalyst Co., Ltd. was weighed, and 1000 ml of an aqueous solution containing 10 ml of chemically pure nitric acid was added. The mixture was extruded using a twin-screw extruder, dried at 120°C for 4 hours, and calcined at 800°C for 4 hours to obtain a catalyst support. The resulting material was impregnated with 900 ml of an aqueous solution containing 120 g of ammonium fluoride for 2 hours, dried at 120°C for 3 hours, and calcined at 600°C for 3 hours. After cooling to room temperature, the resulting material was impregnated with 950 ml of an aqueous solution containing 133 g of ammonium metamolybdate for 3 hours, dried at 120°C for 3 hours, and calcined at 600°C for 3 hours. After cooling to room temperature, the obtained catalyst was impregnated for 4 hours with 900 ml of an aqueous solution containing 180 g of nickel nitrate and 320 g of ammonium metatungstate, and the fluorided alumina support was impregnated for 4 hours with a mixed aqueous solution containing 0.1 wt % of ammonium metamolybdate (chemically pure) and 0.1 wt % of nickel nitrate (chemically pure) relative to the catalyst support, dried at 120°C for 3 hours, and calcined at 600°C for 4 hours to obtain catalyst E.

[0106] [Table 4]

[0107] Example 1 For the riser reactor pilot plant, experiments were carried out according to the scheme shown in Figure 1 as follows: The heavy feedstock was contacted with the catalytic conversion catalyst A at the bottom of a first reaction zone for reaction under conditions including a reaction temperature of 600°C, a reaction pressure of 0.1 MPa, a reaction time of 3 seconds, and a weight ratio of catalyst to feedstock of 5:1. The 1-pentene feedstock was contacted with the catalytic conversion catalyst A from the first reaction zone at the bottom of a second reaction zone for reaction under conditions including a reaction temperature of 530°C, a reaction pressure of 0.1 MPa, a reaction time of 10 seconds, and a weight ratio of catalyst to 1-pentene feedstock of 45:1.

[0108] The resulting reaction products were separated from the spent catalyst, the spent catalyst was regenerated by burning coke in a regenerator, and the reaction products were separated to obtain ethylene, propylene, butylene, an olefin-rich stream, a second catalytic cracking distillate having a boiling point above 250°C, and the like.

[0109] The second catalytic cracked distillate was heated at a temperature of 350°C, a hydrogen partial pressure of 18 MPa, and a temperature of 1.5 h. -1 and a volumetric hydrogen to oil ratio of 1500 to obtain a hydrogenated catalytic cracking distillate.

[0110] The separated butylene was recycled to the bottom of the riser reactor for further cracking under conditions including a reaction temperature of 650°C, a catalyst to butylene weight ratio of 100:1, and a reaction time of 0.2 seconds; the olefin-rich stream was recycled to the bottom of the second reaction zone for further cracking; and the hydrotreated catalytic cracking distillate was mixed with the heavy feedstock and then recycled to the first reaction zone for reaction. The reaction conditions and product distribution are listed in Table 3.

[0111] Comparative Example 1 Anticipating that 1-pentene was not added at the bottom of the second reaction zone and that an olefin-rich stream was not recycled to the riser reactor, an experiment was carried out in a pilot plant riser reactor as described in Example 1 as follows: Heavy feedstock a was contacted with catalytic conversion catalyst A at the bottom of the first reaction zone for reaction under conditions including a reaction temperature of 600°C, a reaction pressure of 0.1 MPa, a reaction time of 3 seconds, and a weight ratio of catalyst to feedstock of 5:1.

[0112] The resulting reaction product was separated from the spent catalyst, the spent catalyst was regenerated by burning coke in a regenerator, and the regenerated catalyst was recycled to the bottom of the riser reactor; the reaction product was separated to obtain ethylene, propylene, butylene, and a catalytic cracking distillate having a boiling point above 250°C.

[0113] The catalytic cracking distillate was heated at 350°C under a hydrogen partial pressure of 18 MPa for 1.5 hours. -1 and a volumetric space velocity of 1,500 and a volumetric ratio of hydrogen to oil of 1,500 to obtain a hydrogenated catalytic cracking distillate oil.

[0114] The separated butylene was recycled to the bottom of the riser reactor for further cracking under conditions including a reaction temperature of 650°C, a catalyst to butylene weight ratio of 100:1, and a reaction time of 0.2 seconds; the hydrotreated catalytic cracking distillate was mixed with the heavy feedstock and then recycled to the first reaction zone for reaction. The reaction conditions and product distribution are listed in Table 3.

[0115] Example 2 For the riser reactor pilot plant, experiments were carried out according to the scheme shown in Figure 1 as follows: The heavy feedstock was contacted with the catalytic conversion catalyst A at the bottom of the first reaction zone for reaction under conditions including a reaction temperature of 600°C, a reaction pressure of 0.1 MPa, a reaction time of 3 seconds, and a weight ratio of catalyst to feedstock of 5:1.

[0116] The resulting reaction product was separated from the spent catalyst, the spent catalyst was regenerated by burning the coke in a regenerator, and the regenerated catalyst was recycled to the bottom of the riser reactor; the reaction product was separated to obtain ethylene, propylene, butylene, an olefin-rich stream, a second catalytic cracking distillate having a boiling point above 250°C, etc.

[0117] The second catalytic cracked distillate was heated at a temperature of 350°C, a hydrogen partial pressure of 18 MPa, and a temperature of 1.5 h. -1 and a volumetric space velocity of 1,500 and a volume ratio of hydrogen to oil of 1,500, to obtain a hydrogenated catalytic cracking distillate oil.

[0118] The separated butylene was recycled to the bottom of the riser reactor for further cracking under conditions including a reaction temperature of 650°C, a catalyst to feed weight ratio of 100:1, and a reaction time of 0.2 seconds; the olefin-rich stream was recycled to the bottom of the second reaction zone for further cracking under conditions including a reaction temperature of 530°C, a reaction pressure of 0.1 MPa, and a reaction time of 10 seconds; the hydrotreated catalytic cracking distillate was mixed with the heavy feedstock and then recycled to the first reaction zone for reaction. The reaction conditions and product distribution are listed in Table 3.

[0119] Comparative Example 2 An experiment was conducted on a pilot plant of a riser reactor. Heavy feedstock a was contacted with catalytic conversion catalyst B at the bottom of the riser reactor for reaction under conditions including a reaction temperature of 610°C, a reaction pressure of 0.1 MPa, a reaction time of 6 seconds, and a weight ratio of catalyst to feedstock of 16.9:1.

[0120] The resulting reaction product was separated from the spent catalyst, and the spent catalyst was regenerated by burning the coke in a regenerator. The regenerated catalyst was recycled to the bottom of the riser reactor; after separating the reaction product, it was not subjected to hydrotreating or further reaction. The reaction conditions and product distribution are listed in Table 3.

[0121] Example 3 The experiment was carried out as described in Example 2, except that the second catalytic cracking distillate oil having a boiling point above 250°C using a heavier heavy feedstock b was contacted with the hydrodesulfurization catalyst E in a hydrodesulfurization reactor and reacted under conditions including a reaction pressure of 6.0 MPa, a reaction temperature of 350°C, a hydrogen-to-oil volume ratio of 350, and a volumetric space velocity of 2.0 h to obtain a low-sulfur hydrotreated catalytic cracking distillate oil that was recovered as a gas oil component without being recycled to the riser reactor for further reaction. The reaction conditions and product distribution are listed in Table 3.

[0122] Comparative Example 3 An experiment was conducted on a pilot plant of a riser reactor. Heavy feedstock b was reacted with catalytic conversion catalyst C at the bottom of the riser reactor under conditions including a reaction temperature of 530°C, a reaction pressure of 0.1 MPa, a reaction time of 6 seconds, and a weight ratio of catalyst to feedstock of 5:1.

[0123] The obtained reaction product was separated from the spent catalyst, and the spent catalyst was regenerated by burning coke in a regenerator, and the regenerated catalyst was recycled to the bottom of the riser reactor; the reaction product obtained after separation was not recycled to the riser reactor for further reaction, and the hydrotreatment of the second catalytic cracking distillate was the same as in Example 3. The reaction conditions and product distribution are listed in Table 3.

[0124] Example 4 The experiment was carried out as described in Example 1, except that the reaction conditions shown in Table 3 were used. The reaction conditions and product distribution are listed in Table 3.

[0125] Example 5 The experiment was carried out as described in Example 1, except that the reaction conditions shown in Table 3 were used. The reaction conditions and product distribution are listed in Table 3.

[0126] Example 6 The experiment was carried out as described in Example 1, except that the reaction conditions shown in Table 3 were used. The reaction conditions and product distribution are listed in Table 3.

[0127] [Table 5] JPEG0007723100000006.jpg233169

[0128] The results shown in Table 3 show that cracking olefin feedstock at high temperatures results in higher propylene yields, while simultaneously increasing ethylene yields, and the higher the olefin content in the feedstock, the better the improvement effect.

[0129] Example 7 For the riser reactor pilot plant, experiments were carried out according to the scheme shown in Figure 2 as follows: The heavy feedstock was contacted with the catalytic conversion catalyst A at the bottom of the first reaction zone I for reaction under conditions including a reaction temperature of 600°C, a reaction pressure of 0.1 MPa, a reaction time of 3 seconds, and a weight ratio of catalyst to feedstock of 5:1. The 1-octene feedstock was contacted with the catalytic conversion catalyst A from the first reaction zone at the bottom of the second reaction zone for reaction under conditions including a reaction temperature of 530°C, a reaction pressure of 0.1 MPa, a reaction time of 10 seconds, and a weight ratio of catalyst to 1-octene of 45:1.

[0130] The resulting reaction product is separated from the spent catalyst, the spent catalyst is regenerated by burning coke in a regenerator, and the regenerated catalyst is recycled to the bottom of the riser reactor; the reaction product is separated to obtain ethylene, propylene, butylene, a first olefin-rich stream, a second olefin-rich stream, a second catalytic cracking distillate having a boiling point above 250°C, etc.

[0131] The second catalytic cracking distillate was heated at a temperature of 350°C, a hydrogen partial pressure of 18 MPa, and a temperature of 1.5 h. -1 and a volumetric hydrogen to oil ratio of 1500 to obtain a hydrogenated catalytic cracking distillate.

[0132] The separated butylene was recycled to the bottom of the riser reactor for further cracking under conditions including a reaction temperature of 650 ° C, a catalyst to feed weight ratio of 100:1, and a reaction time of 0.2 seconds; the first olefin-rich stream (i.e., a stream containing small olefins) having a boiling point below 140 ° C was recycled to the bottom of the second reaction zone for further cracking; the second olefin-rich stream (i.e., a stream containing large olefins) having a boiling point above 140 ° C was recycled to the bottom of the third reaction zone for further cracking under conditions including a reaction temperature of 500 ° C and a reaction time of 10 seconds; the hydrotreated catalytic cracking distillate was mixed with heavy feedstock and then recycled to the first reaction zone for reaction. The reaction conditions and product distribution are listed in Table 4.

[0133] Example 8 For the riser reactor pilot plant, experiments were carried out according to the scheme shown in Figure 3 as follows: Heavy feedstock a was contacted with catalytic conversion catalyst A at the bottom of the first reaction zone for reaction under conditions including a reaction temperature of 600°C, a reaction pressure of 0.1 MPa, a reaction time of 3 seconds, and a weight ratio of catalyst to feedstock of 5:1; 1-pentene feedstock was contacted with catalytic conversion catalyst A from the first reaction zone at the bottom of the second reaction zone for reaction under conditions including a reaction temperature of 530°C, a reaction pressure of 0.1 MPa, a reaction time of 6 seconds, and a weight ratio of catalyst to 1-pentene of 45:1. Methanol was introduced into a lower middle portion of the second reaction zone for reaction under conditions including a reaction temperature of 500°C, a reaction pressure of 0.1 MPa, a reaction time of 3 seconds, and a weight ratio of catalyst to methanol of 10:1.

[0134] The resulting reaction products were separated from the spent catalyst, the spent catalyst was regenerated by burning coke in a regenerator, and the reaction products were separated to obtain ethylene, propylene, butylene, an olefin-rich stream, a second catalytic cracking distillate having a boiling point above 250°C, and the like.

[0135] The second catalytic cracked distillate was heated at a temperature of 350°C, a hydrogen partial pressure of 18 MPa, and a temperature of 1.5 h. -1 and a volumetric hydrogen to oil ratio of 1500 to obtain a hydrogenated catalytic cracking distillate.

[0136] The separated butylene was recycled to the bottom of the riser reactor for further cracking under conditions including a reaction temperature of 650°C, a catalyst to butylene weight ratio of 100:1, and a reaction time of 0.2 seconds; the olefin-rich stream was recycled to the bottom of the second reaction zone for further cracking; and the hydrotreated catalytic cracking distillate was mixed with the heavy feedstock and then recycled to the first reaction zone for reaction. The reaction conditions and product distribution are listed in Table 4.

[0137] [Table 6] JPEG0007723100000008.jpg53169

[0138] As can be seen from the data shown in Table 4, the processes of Examples 7 and 8 of the present application further improved the overall yield of propylene and ethylene, especially the yield of ethylene, and reduced the overall yield of hydrogen, methane, and ethane, compared to Example 1.

[0139] Although the present application has been described in detail above with reference to preferred embodiments, it is not intended to be limited to these embodiments. Various modifications can be made in accordance with the inventive concept of the present application, and these modifications are within the scope of the present application.

[0140] The various technical features described in the above embodiments can be combined as appropriate without contradiction, and in order to avoid unnecessary repetition, the present application does not describe various possible combinations, but such combinations are also within the scope of the present application.

[0141] Furthermore, the various embodiments of the present application can be combined in any manner without departing from the spirit of the present application, and the combined embodiments are considered to be the disclosure content of the present application. [Brief explanation of the drawings]

[0142] [Figure 1] FIG. 1 shows a schematic flow diagram of a preferred embodiment of the fluid catalytic conversion process of the present application. [Figure 2] FIG. 2 shows a schematic flow diagram of another preferred embodiment of the fluid catalytic conversion process of the present application. [Figure 3] FIG. 3 shows a schematic flow diagram of yet another preferred embodiment of the fluid catalytic conversion process of the present application.

Claims

1. A fluid catalytic conversion process for maximizing propylene production, comprising the following steps 1) to 6): 1) introducing a heavy feedstock into a first reaction zone of a fluid catalytic conversion reactor, contacting the heavy feedstock with a catalytic conversion catalyst having a temperature of 650°C or greater, and reacting under first catalytic conversion reaction conditions; 2) introducing a hydrocarbon oil feedstock having an olefin content of 50 wt.% or more into a second reaction zone downstream of the fluid catalytic conversion reactor of the first reaction zone after the reaction of step 1), contacting the catalytic conversion catalyst from the first reaction zone and reacting under second catalytic conversion conditions; 3) separating the effluent of the fluidized catalytic conversion reactor to obtain a reaction product and a spent catalyst, and performing a first separation on the reaction product to obtain ethylene, propylene, butylene, a first catalytic cracking distillate, and a second catalytic cracking distillate; the first catalytic cracking distillate has an initial boiling point greater than 20°C, the second catalytic cracking distillate has an end point less than 550°C, and the cut point between the first catalytic cracking distillate and the second catalytic cracking distillate is in the range of 140°C to 250°C; 4) subjecting the first catalytic cracking distillate to a second separation to obtain an olefin-rich stream having at least 50 wt% C5+ olefins content; 5) recycling at least a portion of the olefin-rich stream to step 2) for further reaction; and 6) recycling at least a portion of the butylenes separated in step 3) upstream of the fluidized catalytic conversion reactor where the heavy feedstock is introduced and contacted for reaction with the catalytic conversion catalyst under third catalytic conversion conditions; wherein the first catalytic conversion conditions include: a reaction temperature of 500 to 800°C; a reaction pressure of 0.05 to 1 MPa; a reaction time of 0.01 to 100 seconds; a weight ratio of the catalytic conversion catalyst to the heavy feedstock of (1 to 200):1; The second catalytic conversion conditions include: a reaction temperature of 400 to 680°C; a reaction pressure of 0.05 to 1 MPa; a reaction time of 0.01 to 100 seconds; a weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock of (1 to 100):1; and The third catalytic conversion conditions include: a reaction temperature of 650 to 800°C; a reaction pressure of 0.05 to 1 MPa; a reaction time of 0.01 to 10 seconds; a weight ratio of said catalytic conversion catalyst to said butylenes of (20-200):1; 2. The first catalytic conversion conditions include: a reaction temperature of 510 to 780°C; a reaction pressure of 0.1 to 0.8 MPa; a reaction time of 0.1 to 80 seconds; a weight ratio of the catalytic conversion catalyst to the heavy feedstock of (3 to 180):1; The second catalytic conversion conditions include: a reaction temperature of 450 to 650°C; a reaction pressure of 0.1 to 0.8 MPa; a reaction time of 0.1 to 80 seconds; a weight ratio of the catalytic conversion catalyst to the hydrocarbon oil feedstock of (3 to 70):1; and 10. The method of claim 1, wherein the third catalytic conversion conditions comprise: a reaction temperature of 680-780°C; a reaction pressure of 0.1 to 0.8 MPa; reaction time of 0.05 to 8 seconds; a weight ratio of said catalytic conversion catalyst to said butylenes of (30-180):1;

3. The method according to claim 1 or 2, further comprising the following step 2a): 2a) introducing an oxygen-containing organic compound into the second reaction zone of the fluidized catalytic conversion reactor for reaction contact with the catalytic conversion catalyst under fourth catalytic conversion conditions comprising: a reaction temperature of 300 to 550°C; a reaction pressure of 0.05 to 1 MPa; a reaction time of 0.01 to 100 seconds; a weight ratio of said catalytic conversion catalyst to said oxygen-containing organic compound of (1-100):1; 4. The method of claim 3, wherein the fourth catalytic conversion conditions include: a reaction temperature of 400 to 530°C; a reaction pressure of 0.1 to 0.8 MPa; a reaction time of 0.1 to 80 seconds; a weight ratio of the catalytic conversion catalyst to the oxygen-containing organic compound of (3 to 80):1; The oxygen-containing organic compound includes at least one of methanol, ethanol, dimethyl ether, methyl ethyl ether, and ethyl ether.

5. The method according to any one of claims 1 to 4, further comprising the following step 7): 7) hydrotreating the second catalytic cracking distillate to obtain a hydrogenated catalytic cracking distillate, and recycling the hydrogenated catalytic cracking distillate to the first reaction zone of the fluid catalytic conversion reactor for further reaction.

6. The method according to claim 5, wherein the hydrotreating conditions include a hydrogen partial pressure of 3.0-20.0 MPa, a reaction temperature of 300-450°C, a volume ratio of hydrogen to oil of 300-2000, and a volumetric space velocity of 0.1-3.0 h −1 .

7. The method according to any one of claims 1 to 6, further comprising the following step 8): 8) regenerating the spent catalyst obtained by the separation in step 3) by burning coke to obtain a regenerated catalyst having a temperature of 650°C or higher, and then recycling the regenerated catalyst upstream of the first reaction zone of the fluidized catalytic conversion reactor for use as the catalytic conversion catalyst.

8. The heavy feedstock used in step 1) is selected from the group consisting of petroleum hydrocarbons and mineral oils; the petroleum hydrocarbons are selected from vacuum gas oils, atmospheric gas oils, coked gas oils, deasphalted oils, vacuum residua, atmospheric residua, heavy aromatic raffinates, or combinations thereof; the mineral oils are selected from coal liquids, oil sands oils, shale oils, or combinations thereof; and The method according to any one of claims 1 to 7, wherein the hydrocarbon oil feedstock used in step 2) has an olefin content of 80 wt% or more.

9. The method of claim 8, wherein the hydrocarbon oil feedstock is a pure olefin feedstock.

10. The method of claim 8, wherein the olefins in the hydrocarbon oil feedstock are obtained from a C4+ fraction produced by dehydrogenation of an alkane feedstock, a C4+ fraction produced by a catalytic cracking unit in an oil refinery, a C4+ fraction produced by a steam cracking unit in an ethylene plant, a C4+ olefin-rich by-product fraction of an MTO process, and a C4+ olefin-rich by-product fraction of an MTP process.

11. The second separation of step 4) further comprises splitting the olefin-rich stream into a first olefin-rich stream having a lower boiling point and a second olefin-rich stream having a higher boiling point; the cut point between the first stream and the second stream is in the range of 140 to 200°C; and 11. The method of any one of claims 1 to 10, wherein step 5) further comprises introducing the first olefin-rich stream into the second reaction zone of the fluidized catalytic conversion reactor for further reaction, and introducing the second olefin-rich stream into a third reaction zone of the fluidized catalytic conversion reactor downstream of the second reaction zone for further reaction.

12. 12. The method of any one of claims 1 to 11, wherein the fluidized catalytic conversion reactor is selected from a riser reactor, a fluidized bed reactor, an up-transfer line, a down-transfer line, or a combination of two or more thereof.

13. The method of claim 12, wherein the fluidized catalytic conversion reactor is a diameter conversion riser reactor.

14. The catalytic conversion catalyst comprises, based on the weight of the catalytic conversion catalyst, 1 to 50 wt. % of a molecular sieve, 5 to 99 wt. % of an inorganic oxide, and 0 to 70 wt. % of clay; The molecular sieve comprises at least one of a macroporous molecular sieve, a mesoporous molecular sieve, and a microporous molecular sieve; and The method according to any one of claims 1 to 13, wherein the catalytic conversion catalyst further comprises 0.1 to 3 wt% of an active metal, based on the weight of the catalytic conversion catalyst; and the active metal is at least one selected from the group consisting of Group VIII metals, Group IVA metals, and rare earth metals.

15. 4. The method of claim 3, wherein the oxygen-containing organic compound is mixed with the hydrocarbon oil feed before being fed to the second reaction zone of the fluid catalytic conversion reactor, or the oxygen-containing organic compound is fed to the second reaction zone of the fluid catalytic conversion reactor downstream from where the hydrocarbon oil feed is introduced.

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