Method for increasing yields of ethylene and propylene

Abstract

The invention relates to a method for increasing the yield of ethylene and propylene. The method comprises the following steps: contacting an olefin-containing raw material with a hot regenerated first catalyst in a dense-phase bed layer of a first reactor to carry out a first catalytic conversion reaction; enabling the reacted material flow to be in contact with a heat collector arranged at the downstream of the dense-phase bed layer so as to carry out instantaneous cooling, cooling the reacted material flow to a target temperature of 450-530 DEG C, and then carrying out separation treatment on the cooled material flow so as to obtain first reaction oil gas rich in ethylene and propylene; part of the first catalyst in the dense-phase bed layer is taken out from a slide valve at the bottom of the first reactor, so that the density of the first catalyst in the dense-phase bed layer is 100-700 kg / m < 3 >. By increasing the density of the catalyst and reducing the bed height of the first catalyst, the contact reaction time of the hydrocarbon oil raw material and the first catalyst is shortened, the catalytic cracking reaction is enhanced, the yield of low-carbon olefin is increased, and the yield of dry gas in the product is reduced.

Description

Technical Field

[0001] The invention belongs to the field of petrochemical industry, and in particular, relates to a method for increasing the production of ethylene and propylene. Background Art

[0002] Low-carbon olefins represented by ethylene and propylene are important basic chemical raw materials and are widely used in various fields. At present, the main raw materials for producing low-carbon olefins are petroleum hydrocarbons such as naphtha, ethane, light diesel and hydrogenated tail oil. The typical process is to use naphtha steam cracking. However, the output of ethylene and propylene in this process is low, which makes it difficult to meet the market's growing demand for low-carbon olefins.

[0003] CN114763483A provides a method for producing light olefins using olefin-rich raw materials, wherein the hydrocarbon oil raw material is contacted with a catalyst at a temperature above 650°C for a catalytic conversion reaction, and the butenes and olefin-rich streams separated from the reaction oil gas are returned to the reactor for further reaction to increase the yield of light olefins.

[0004] CN114763483A discloses a method for producing low-carbon olefins by reacting olefin-rich raw materials in a riser reactor, wherein the olefin-rich raw materials are first subjected to high-temperature cracking, and then the olefins in the separated products are returned to the reactor for further reaction to produce ethylene and propylene.

[0005] However, the existing methods for producing low-carbon olefins have problems such as low ethylene and propylene content in the products obtained and excessive energy consumption. Summary of the invention

[0006] The object of the present invention is to provide a method which can simultaneously increase the yield of ethylene and propylene and reduce the reaction energy consumption.

[0007] In order to achieve the above object, the present invention provides a method for increasing the production of ethylene and propylene, the method comprising the following steps:

[0008] The olefin-containing raw material is brought into contact with the hot regenerated first catalyst in the dense phase bed of the first reactor to carry out a first catalytic conversion reaction; the post-reaction logistics is brought into contact with a heat collector arranged downstream of the dense phase bed to carry out instantaneous cooling, so that the post-reaction logistics is cooled to a target temperature of 450 to 530° C., and then the cooled logistics is separated and treated to obtain a first reaction oil gas rich in ethylene and propylene;

[0009] A portion of the first catalyst in the dense phase bed is taken out from the slide valve at the bottom of the first reactor so that the density of the first catalyst in the dense phase bed is 100 to 700 kg / m 3 .

[0010] Optionally, the first reactor includes, from bottom to top, a first lifting section, an expansion section, a dense phase fluidization section, a reduction section and a second lifting section; the angle between the side wall of the expansion section and the axis is 5 to 60 degrees; the ratio of the diameter to the height of the dense phase fluidization section is (1 to 20): 1; the angle between the side wall of the reduction section and the axis is 5 to 60 degrees; a material distributor is also provided at the head of the dense phase fluidization section; the cooling zone is provided at the tail of the dense phase fluidization section; the heat exchanger is provided along the circumference of the tail of the first reactor; the cooling medium in the heat exchanger is selected from one or more of liquefied gas, crude gasoline, stabilized gasoline, diesel, heavy diesel and water.

[0011] Optionally, the time for the reaction flow to cool down to the target temperature is within 2 seconds, preferably within 1 second, and more preferably within 0.5 seconds.

[0012] Optionally, the catalyst density of the first catalyst in the dense phase bed of the first reactor is 300-650 kg / m 3 , more preferably 500 to 650 kg / m 3 .

[0013] Optionally, the first catalyst is a mesopore molecular sieve catalyst and / or a small pore molecular sieve catalyst; the mesopore molecular sieve catalyst is selected from one or more of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38 and ZSM-48; the small pore molecular sieve catalyst is selected from one or more of SAPO-34, SAPO-11 and SAPO-47.

[0014] Optionally, the olefin content in the olefin-containing raw material is 50 to 100% by weight, preferably 80 to 100% by weight; the olefin-containing raw material is a gas phase feed; the olefin in the olefin-containing raw material is C4 to C 12 Olefins, preferably C4-C6 olefins;

[0015] Optionally, the conditions of the first catalytic conversion reaction include: a reaction temperature of 550 to 800°C, preferably 580 to 750°C, and further preferably 600 to 700°C; a reaction pressure of 0.01 to 1MPa, preferably 0.1 to 0.8MPa, and further preferably 0.2 to 0.5MPa; a reaction time of 0.05 to 10 seconds, preferably 0.1 to 5 seconds, and further preferably 0.2 to 2 seconds; a weight ratio of the first catalyst to the olefin-containing raw material is (5 to 50):1, preferably (10 to 35):1, and further preferably (12 to 30):1.

[0016] Optionally, the method further comprises: stripping and charring the first catalyst after the reaction, and then transporting it to a regenerator for charring regeneration.

[0017] Optionally, the method further comprises the following steps: contacting the heavy hydrocarbon oil with a second catalyst in a second reactor to perform a second catalytic conversion reaction to obtain reaction oil gas containing light oil and a catalyst to be produced;

[0018] Then, the reaction oil gas containing light oil is mixed with the first reaction oil gas, and the mixed oil gas is subjected to a first separation to obtain dry gas, light olefins and C5 and above distillate oil, and the C5 and above distillate oil is subjected to a second separation and the separated olefin-rich components are returned to the first reactor; the heavy hydrocarbon oil is selected from one or more of vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum residue oil, atmospheric residue oil, heavy aromatic raffinate oil, coal liquefaction oil, oil sand oil and shale oil; the second catalyst is the regenerated first catalyst or a mixed catalyst of the first catalyst after reaction and the regenerated first catalyst.

[0019] Optionally, when the temperature of the reaction oil gas containing light oil and the catalyst to be regenerated in the upper area of ​​the reaction zone of the second reactor exceeds 530°C, a cooling zone is set downstream of the reaction zone to cool the reaction oil gas containing light oil and the catalyst to be regenerated to below 530°C.

[0020] Optionally, the method further comprises: stripping the first catalyst and the catalyst to be regenerated after the reaction respectively, and then transporting them to a regenerator for char regeneration.

[0021] Optionally, the conditions of the second catalytic conversion reaction include: a reaction temperature of 500-600°C, preferably 530-580°C; a reaction pressure of 0.01-1MPa, preferably 0.05-1MPa; a reaction time of 0.01-100s, preferably 0.1-80s; preferably, a weight ratio of the second catalyst to the heavy hydrocarbon oil is (1-100):1, preferably (3-70):1.

[0022] Through the above technical scheme, the present invention reduces the bed height of the first catalyst by increasing the catalyst density, shortens the contact reaction time between the hydrocarbon oil feedstock and the first catalyst, enhances the catalytic cracking reaction, and significantly inhibits the thermal cracking reaction, thereby achieving the goal of increasing the production of light olefins and reducing the dry gas yield; by rapidly cooling and separating the post-reaction logistics, the deep reaction can be inhibited and the by-product yield can be reduced. At the same time, taking out the first catalyst from the bottom of the reaction zone can not only flexibly adjust the catalyst density, but also save energy consumption for cooling the post-reaction logistics.

[0023] Other features and advantages of the present invention will be described in detail in the following detailed description. BRIEF DESCRIPTION OF THE DRAWINGS

[0024] The accompanying drawings are used to provide a further understanding of the present invention and constitute a part of the specification. Together with the following specific embodiments, they are used to explain the present invention but do not constitute a limitation of the present invention. In the accompanying drawings:

[0025] Figure 1 A schematic diagram of an embodiment of the first reactor provided by the present invention.

[0026] Figure 2 A schematic diagram of another embodiment of the first reactor provided by the present invention.

[0027] Figure 3 A schematic diagram of an embodiment of the method provided by the present invention.

[0028] Figure 4 A schematic diagram of another embodiment of the method provided by the present invention.

[0029] Figure 5 A schematic diagram of an embodiment of the catalytic conversion method provided by the present invention.

[0030] Figure 6 A schematic diagram of another embodiment of the catalytic conversion method provided by the present invention.

[0031] Description of Reference Numerals

[0032] A The first reactor B The second reactor

[0033] 1 pipeline 2 pipeline 3 material distributor

[0034] 4 Heat exchanger 5 Pipeline 6 Pipeline

[0035] 7 slide valve 8 pipeline 9 stripping section

[0036] 10 second lifting section 11 horizontal pipe 12 settler

[0037] 13 cyclone separator 14 air collecting chamber 15 inclined tube

[0038] 16 regenerator 17 pipeline 18 pipeline

[0039] 19 pipeline 20 large oil and gas pipeline 21 distillation tower

[0040] 22 pipeline 23 pipeline 24 regeneration inclined pipe

[0041] 25 pipeline 26 pipeline 27 pipeline

[0042] 28 pipeline 29 pipeline 30 olefin separation unit

[0043] 31 pipeline 32 pipeline 33 first lifting section

[0044] 34 cooling zone 35 distributor 36 pipeline

[0045] 37 pipeline 38 pipeline DETAILED DESCRIPTION

[0046] The specific embodiments of the present invention are described in detail below. It should be understood that the specific embodiments described herein are only used to illustrate and explain the present invention, and are not used to limit the present invention.

[0047] The present invention provides a method for increasing the production of ethylene and propylene, the method comprising the following steps:

[0048] The olefin-containing raw material is brought into contact with the hot regenerated first catalyst in the dense phase bed of the first reactor to carry out a first catalytic conversion reaction; the post-reaction logistics is brought into contact with a heat collector arranged downstream of the dense phase bed to carry out instantaneous cooling, so that the post-reaction logistics is cooled to a target temperature of 450-530° C., and then the cooled logistics is separated and treated to obtain a first reaction oil gas rich in ethylene and propylene; a portion of the first catalyst in the dense phase bed is taken out from the slide valve at the bottom of the first reactor so that the density of the first catalyst in the dense phase bed is 100-700 kg / m 3 .

[0049] The inventors have found through a large number of experiments that the product distribution of olefin catalytic cracking in different types of reactors and on catalysts of different densities is significantly different. It is also found that when the density is greater than 100 kg / m 3 When the catalytic reaction is carried out on a catalyst, the yield of methane in the olefin catalytic cracking products can be reduced, and the yield and selectivity of light olefins can be increased at the same time.

[0050] The present invention reduces the bed height of the first catalyst by increasing the catalyst density, shortens the contact reaction time between the hydrocarbon oil feedstock and the first catalyst, enhances the catalytic cracking reaction, and significantly inhibits the thermal cracking reaction, thereby achieving the goal of increasing the production of light olefins and reducing the dry gas yield; by rapidly cooling and separating the post-reaction logistics, the deep reaction can be inhibited and the by-product yield can be reduced. At the same time, taking out the first catalyst from the bottom of the reaction zone can not only flexibly adjust the catalyst density, but also save the energy consumption for cooling the post-reaction logistics.

[0051] The first catalyst is preheated and then fed into the first reactor, wherein the energy for preheating the catalyst can be provided by electric heating, fuel combustion, etc. For example, by-product gas and / or inferior heavy oil produced by the method of the present invention is used as fuel, or fuel oil or fuel gas produced by other devices in the refinery, so that the temperature of the catalyst entering the first reactor is at an appropriate temperature.

[0052] Among them, the first reactor includes, from bottom to top, a first lifting section, an expansion section, a dense phase fluidization section, a reduction section and a second lifting section; the angle between the side wall of the expansion section and the axis is 5 to 60 degrees; the ratio of the diameter to the height of the dense phase fluidization section is (1 to 20): 1; the angle between the side wall of the reduction section and the axis is 5 to 60 degrees; a material distributor is also provided at the head of the dense phase fluidization section; the cooling zone is provided at the tail of the dense phase fluidization section.

[0053] The heat collector is arranged along the circumference of the tail of the first reactor; the cooling medium in the heat collector is selected from one or more of liquefied gas, crude gasoline, stabilized gasoline, diesel, heavy diesel and water. The heat collector is arranged along the circumference of the cooling zone, or is evenly spaced along the circumference of the cooling zone, which is conducive to the cooling medium in the heat collector to instantly reduce the temperature of the reaction product, and the reaction product is cooled to 450-530°C within 2 seconds, so as to avoid deep reaction and secondary reaction caused by excessive temperature of the reaction product.

[0054] Wherein, the catalyst density of the first catalyst in the dense phase bed of the first reactor is 300-650 kg / m 3 , more preferably 500 to 650 kg / m 3 In some specific embodiments of the present invention, the catalyst density of the first catalyst in the dense phase bed is 300 kg / m 3 、350kg / m 3 , 400kg / m 3 , 450kg / m 3 , 500kg / m 3 、550kg / m 3 、600kg / m 3 、650kg / m 3 Or any value within the aforementioned range.

[0055] Wherein, the first catalyst is a mesopore molecular sieve catalyst and / or a small pore molecular sieve catalyst; the mesopore molecular sieve catalyst is selected from one or more of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38 and ZSM-48; the small pore molecular sieve catalyst is selected from one or more of SAPO-34, SAPO-11 and SAPO-47.

[0056] The olefin content in the olefin-containing raw material is 50 to 100% by weight, preferably 80 to 100% by weight; the olefin-containing raw material is a gas phase feed; the olefin in the olefin-containing raw material is C4 to C 12 Olefins, preferably C4-C6 olefins.

[0057] Among them, the conditions of the first catalytic conversion reaction include: reaction temperature of 550-800°C, preferably 580-750°C, and more preferably 600-700°C; reaction pressure of 0.01-1MPa, preferably 0.1-0.8MPa, and more preferably 0.2-0.5MPa; reaction time of 0.05-10 seconds, preferably 0.1-5 seconds, and more preferably 0.2-2 seconds; the weight ratio of the first catalyst to the olefin-containing raw material is (5-50):1, preferably (10-35):1, and more preferably (12-30):1.

[0058] The method further comprises: stripping the first catalyst after the reaction, and then transporting it to a regenerator for char regeneration.

[0059] The method further comprises the following steps: contacting the heavy hydrocarbon oil with the second catalyst in the second reactor for a second catalytic conversion reaction to obtain reaction oil gas containing light oil and a catalyst to be regenerated; then mixing the reaction oil gas containing light oil with the first reaction oil gas, and subjecting the mixed oil gas to a first separation to obtain dry gas, light olefins and C5 and above distillate oil, subjecting the C5 and above distillate oil to a second separation and returning the separated olefin-rich components to the first reactor; the heavy hydrocarbon oil is selected from one or more of vacuum gas oil, atmospheric gas oil, coking gas oil, deasphalted oil, vacuum residue oil, atmospheric residue oil, heavy aromatic raffinate oil, coal liquefaction oil, oil sand oil and shale oil; the second catalyst is the regenerated first catalyst or a mixed catalyst of the first catalyst after reaction and the regenerated first catalyst.

[0060] Wherein, when the temperature of the reaction oil gas containing light oil and the catalyst to be regenerated in the upper region of the reaction zone of the second reactor exceeds 530° C., a cooling zone is provided downstream of the reaction zone to cool the reaction oil gas containing light oil and the catalyst to be regenerated to below 530° C. By providing a cooling zone in the second reactor to timely remove part of the heat of the product of the heavy oil reaction, deep reaction and secondary reaction caused by excessively high temperature of the reaction product can be avoided.

[0061] The method further comprises: stripping the first catalyst and the catalyst to be regenerated after the reaction, and then charring and regenerating them, and then returning the regenerated catalysts to the first reactor and the second reactor respectively; preferably, the conditions of the second catalytic conversion reaction include: reaction temperature of 500-600°C, preferably 530-580°C; reaction pressure of 0.01-1MPa, preferably 0.05-1MPa; reaction time of 0.01-100s, preferably 0.1-80s; preferably, the weight ratio of the second catalyst to the heavy hydrocarbon oil is (1-100):1, preferably (3-70):1.

[0062] Figure 1 is a schematic structural diagram of the first reactor of the present invention. It should be noted that: Figure 1 The device shown is merely an example of a device that can be applied to a method of an embodiment of the present invention, and does not mean that the method of the embodiment of the present invention cannot use other devices or equipment.

[0063] like Figure 1 As shown, the first reactor A includes a first lifting section 33, an expansion section, a dense phase fluidization section, a reduction section and a second lifting section 10 in order from bottom to top along its axial direction. Among them, the diameter of the first lifting section is 0.2 to 5m, and the height of the first lifting section accounts for 5 to 20% of the total height of the first reactor. The angle between the side wall of the expansion section and the axis is 5 to 60 degrees; the height of the expansion section accounts for 1 to 10% of the height of the first reactor, the diameter and height ratio of the dense phase fluidization section is (1 to 20): 1, and its height accounts for 20 to 70% of the total height of the first reactor. One or more gas-solid mixed fluid distributors are provided at the bottom of the reaction zone. The reduction section is located at the junction of the dense phase fluidization section and the second lifting section, and the angle between the side wall of the reduction section and the axis is 5 to 60 degrees, and the height of the reduction section accounts for 5 to 20% of the total height of the reactor. The structure of the second lifting section is similar to the top outlet part of the lifting tube reactor, and the ratio of the diameter of the second lifting section to the diameter of the reaction zone is (0.2-0.8):1, and the height of the second lifting section accounts for 0-20% of the total height of the reactor.

[0064] exist Figure 1 In the reaction device shown, the pre-lifting medium enters the first lifting section 33 from the bottom of the first reactor A through pipeline 1. The first catalyst preheated to above 650°C from pipeline 24 accelerates upward along the first lifting section under the action of the pre-lifting medium. The olefin-containing raw material is injected into the bottom of the reaction zone of the first reactor through pipeline 2, and after being distributed by the material distributor 3, it contacts with the hot first catalyst to react. The reacted first catalyst is controlled by the slide valve 7 and removed from the first reactor A through pipeline 8. Several heat collectors are arranged along the circumference of the cooling zone. The cooling medium is introduced into the heat collector through the pipeline. The reaction product moves upward and contacts with the heat collector to cool down. The reaction product is cooled to 450-550°C within 2 seconds to avoid deep reaction and secondary reaction caused by excessive temperature of the reaction product.

[0065] The olefin-containing feedstock can also be introduced into the reaction zone from other locations in the reaction zone. For example, in another embodiment of the present invention, Figure 2 As shown, the olefin-containing feedstock is injected into the material distributor of the first reactor A through pipeline 36, and after being distributed through the material distributor, it contacts with the hot first catalyst and reacts.

[0066] After the reaction, the first catalyst is removed from the first reactor and can be sent to a regenerator for regeneration after stripping, or it can be mixed with the regenerated catalyst and returned to the reactor for reaction. Figure 3 As shown, in a specific embodiment of the present invention, the first catalyst after the reaction is controlled by the slide valve 7 and introduced into the stripping section 9 from the bottom of the reaction zone through the pipeline 8 after gas-solid separation. The first catalyst after the reaction entering the stripping section 9 contacts with the stripping medium from the pipeline 25 to remove the reaction oil gas on the catalyst. The catalyst after the stripping enters the regenerator 16 through the inclined pipe 15, contacts with the oxygen-containing gas from the pipeline 17 to burn and regenerate, so as to burn off the coke loaded on the catalyst and obtain the regenerated first catalyst. The supplementary fuel is introduced into the regenerator through the pipeline 18, and the regenerated first catalyst enters the first reactor A through the regeneration inclined pipe 24 after the heat is supplemented by the combustion of the fuel.

[0067] The reaction oil and gas obtained from the steam stripping of the first catalyst after the reaction in the settler 12 is separated by the cyclone separator 13 and then enters the gas collecting chamber 14, and then enters the subsequent distillation tower 21 through the large oil and gas pipeline 20 for the first separation. The separated hydrogen, methane and ethane are led out through the pipeline 22, ethylene is led out through the pipeline 23, propylene is led out through the pipeline 26, butene is led out through the pipeline 27 and returned to or not returned to the first reactor, propane and butane are led out through the pipeline 28, and the first catalytic cracking distillate oil with an initial boiling point of any temperature between 20 and 140°C is introduced into the olefin separation device 30 through the pipeline 29 for the second separation treatment to obtain a logistics containing no olefins and a logistics rich in olefins. The logistics containing no olefins is led out through the pipeline 31, and the logistics rich in olefins is returned to the first reactor A through the pipeline 32 as a raw material for further reaction.

[0068] In another embodiment of the present invention, Figure 4 As shown, the reacted first catalyst is controlled by slide valve 7 and enters the bottom of first lifting section 33 through pipeline 8 to mix with the regenerated first catalyst from pipeline 24, and then accelerates upward into the reaction zone under the lifting effect of the pre-lifting medium.

[0069] The first reactor can also be coupled to the second reactor. The first catalyst after reaction taken out from the first reactor is stripped and sent to the regenerator for regeneration, and then the regenerated catalyst is sent to the first reactor A and the second reactor, or the reacted catalyst is mixed with the regenerated catalyst from the regenerator and sent to the second reactor. In one embodiment of the present invention, Figure 5As shown, the pre-lifting medium enters the first lifting section 33 from the bottom of the first reactor A through pipeline 1, and the preheated first catalyst from pipeline 24 accelerates upward along the first lifting section under the lifting effect of the pre-lifting medium. The olefin-containing raw material is injected into the bottom of the reaction zone of the first reactor A through pipeline 2, and contacts with the hot first catalyst after being distributed by the material distributor 3 to react. The reacted oil and gas contact with the heat collector 4 in the cooling zone to cool down. The reacted first catalyst is controlled by the slide valve 7 and sent to the stripping section 9 from the bottom of the reaction zone of the first reactor A through pipeline 8 after gas-solid separation. The reacted first catalyst contacts with the stripping medium from pipeline 25 to remove the reaction oil and gas loaded on the catalyst. The reaction oil and gas steamed from the catalyst is introduced into the settler 12 through the second lifting section, and then separated by the cyclone separator 13 and enters the gas collecting chamber 14. The reacted first catalyst after stripping enters the regenerator 16 for regeneration, and enters the first reactor A through pipeline 24 and the second reactor B through pipeline 19 after supplementing heat.

[0070] The heavy crude oil is injected into the bottom of the second reactor B through pipeline 5 together with the atomized steam from pipeline 6, and reacts with the regenerated first catalyst introduced through pipeline 19. The product of the heavy crude oil reaction contacts the cooling medium from pipeline 38 for instantaneous cooling, and enters the gas collecting chamber after being separated from the reaction oil and gas from the first reactor through a cyclone separator, and is sent to the fractionation tower 21 through a large oil and gas pipeline 20 for the first separation. The separated hydrogen, methane and ethane are led out through pipeline 22, and ethylene is led out through pipeline 23. The first catalytic cracking distillate oil having an initial boiling point of any temperature between 20 and 140° C. is introduced into an olefin separation unit 30 via a pipeline 29 for a second separation treatment to obtain an olefin-free stream and an olefin-rich stream. The olefin-free stream is drawn out via a pipeline 31, and the olefin-rich stream is returned to the first reactor via a pipeline 32 as a raw material for further reaction.

[0071] In another embodiment, Figure 6 As shown, the first catalyst to be regenerated in the first reactor has different destinations. The first catalyst to be regenerated is controlled by the slide valve 7 to enter the second reactor through the pipeline 37, mixes with the regenerated catalyst from the pipeline 19, and contacts and reacts with the heavy feedstock oil from the pipeline 5. The reacted catalyst to be regenerated enters the stripping section 9 for stripping and is sent to the regenerator for regeneration through the pipeline 15, and then returns to the first reactor and the second reactor after preheating.

[0072] The present invention is further described in detail below by way of examples, but the present invention is not limited thereby. The raw materials used in the examples can all be obtained through commercial routes.

[0073] The catalyst used in the examples and comparative examples is a commercial catalyst produced by the Qilu Branch of Sinopec Catalyst Co., Ltd., with a trade name of TCC-1. The properties of the heavy crude oil (heavy oil I) used in the examples and comparative examples are shown in Table 1.

[0074] Table 1

[0075]

[0076] Example 1

[0077] The experiment of this example was Figure 3 The first reactor shown is carried out on a scaled-up unit.

[0078] The hot raw material 1-octene enters the bottom of the reaction zone of the first reactor and mixes with the preheated 1-octene at 650°C and a density of 400 kg / m 3 The oil and gas after the reaction are quickly cooled by the heat exchanger in the cooling zone to below 530°C, and the reaction oil and gas and the catalyst are quickly separated to obtain oil and gas rich in ethylene and propylene, which are sent to the fractionation tower for separation. The catalyst after the reaction is taken out from the bottom of the reaction zone of the reactor by the slide valve. The reaction conditions and product distribution are shown in Table 2.

[0079] Example 2

[0080] The experiment of this example was Figure 3 The hot raw material 1-octene enters the bottom of the reaction zone of the first reactor and is preheated to 650°C and has a density of 650 kg / m 3 The oil and gas after the reaction are rapidly cooled in the cooling zone to below 530°C, and the reaction oil and gas and the catalyst are rapidly separated to obtain oil and gas rich in ethylene and propylene, which are then sent to the fractionation tower for separation. The catalyst after the reaction is taken out from the bottom of the reaction zone of the reactor by the control of the slide valve. The reaction conditions and product distribution are shown in Table 2.

[0081] Comparative Example 1

[0082] This comparative example was carried out according to the method of Example 1. The difference from Example 1 is that the catalyst density in the first reactor is 90 kg / m 3 , and no heat collector is set, the reaction conditions and product distribution are shown in Table 2.

[0083] Table 2

[0084]

[0085] It can be seen from Table 2 that by adopting the method of the present invention, the contents of ethylene, propylene and butene in the reaction oil gas obtained by high-temperature cracking of 1-octene in the first reactor are as high as 69.18%, and ethylene and propylene have high yield and selectivity.

[0086] Example 3

[0087] The experiment of this example was Figure 5 The process shown is carried out on a pilot plant with the first reactor.

[0088] The hot raw material 1-pentene enters the bottom of the reaction zone of the first reactor and mixes with the preheated 1-pentene at a density of 400 kg / m 3 The catalyst of the first reactor is contacted with the catalyst to cause a catalytic conversion reaction; the catalyst after the reaction is controlled by a slide valve from the bottom of the reaction zone of the first reactor, and is introduced into the stripping section of the second reactor through gas-solid separation to obtain a catalyst to be regenerated; the heavy raw material and the catalyst are contacted and reacted at the bottom of the second reactor, the reaction temperature is 560°C, the reaction pressure is 0.2MPa, the reaction time is 6 seconds, and the weight ratio of the catalytic conversion catalyst A to the heavy oil I is 5:1. The second reaction oil gas and the catalyst to be regenerated are obtained; the catalyst to be regenerated enters the regenerator for regeneration; the reaction product is cut according to the distillation range in the product fractionation device to obtain products such as ethylene, propylene, butene and crude olefin-containing logistics. The crude olefin-containing logistics continues to enter the olefin separation device to obtain an olefin-rich logistics, the olefin content in the olefin-rich logistics is 80% by weight, and the olefin-rich logistics is introduced into the second reactor from the bottom of the reaction zone of the second reactor. The reaction conditions and product distribution are shown in Table 3.

[0089] Comparative Example 2

[0090] This comparative example was carried out according to the method of Example 3. The difference between this comparative example and Example 3 is that the catalyst density in the first reactor is 90 kg / m 3 , and no heat collector is set. The reaction conditions and product distribution are shown in Table 3.

[0091] Table 3

[0092]

[0093]

[0094] As can be seen from Table 3, by coupling the catalytic conversion of light feedstocks and heavy oils, the proportion of ethylene and propylene in the obtained reaction oil and gas is increased, especially the yield of ethylene in the product, thus realizing the recycling of heavy oil and reducing costs.

[0095] The preferred embodiments of the present invention are described in detail above. However, the present invention is not limited to the specific details in the above embodiments. Within the technical concept of the present invention, a variety of simple modifications can be made to the technical solution of the present invention, and these simple modifications all belong to the protection scope of the present invention.

[0096] It should also be noted that the various specific technical features described in the above specific embodiments can be combined in any suitable manner without contradiction. In order to avoid unnecessary repetition, the present invention will not further describe various possible combinations.

[0097] In addition, various embodiments of the present invention may be arbitrarily combined, and as long as they do not violate the concept of the present invention, they should also be regarded as the contents disclosed by the present invention.

Claims

1. A method for increasing the production of ethylene and propylene, characterized in that: The method comprises the following steps: The olefin-containing raw material is brought into contact with the hot regenerated first catalyst in the dense phase bed of the first reactor to carry out a first catalytic conversion reaction; the post-reaction logistics is brought into contact with a heat collector arranged downstream of the dense phase bed to carry out instantaneous cooling, so that the post-reaction logistics is cooled to a target temperature of 450 to 530° C., and then the cooled logistics is separated and treated to obtain a first reaction oil gas rich in ethylene and propylene; A portion of the first catalyst in the dense phase bed is taken out from the slide valve at the bottom of the first reactor so that the density of the first catalyst in the dense phase bed is 100 to 700 kg / m 3 .

2. The method according to claim 1, wherein: The first reactor comprises, from bottom to top, a first lifting section, an expansion section, a dense phase fluidization section, a reduction section, and a second lifting section; the angle between the side wall of the expansion section and the axis is 5 to 60 degrees; the ratio of the diameter to the height of the dense phase fluidization section is (1 to 20): 1; the angle between the side wall of the reduction section and the axis is 5 to 60 degrees; The head of the dense phase fluidization section is also provided with a material distributor; The cooling zone is arranged at the tail of the dense phase fluidization section; The heat collector is arranged along the circumference of the tail of the first reactor; the cooling medium in the heat collector is selected from one or more of liquefied gas, crude gasoline, stabilized gasoline, diesel, heavy diesel and water.

3. The method according to claim 1, wherein: The time for the reaction flow to cool down to the target temperature is within 2 seconds, preferably within 1 second, and more preferably within 0.5 seconds.

4. The method according to claim 1, wherein: The catalyst density of the first catalyst in the dense phase bed of the first reactor is 300-650 kg / m 3 , more preferably 500 to 650 kg / m 3 .

5. The method according to claim 1, wherein: The first catalyst is a medium pore molecular sieve catalyst and / or a small pore molecular sieve catalyst; The mesoporous molecular sieve catalyst is selected from one or more of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38 and ZSM-48; The small pore molecular sieve catalyst is selected from one or more of SAPO-34, SAPO-11 and SAPO-47.

6. The method according to claim 1, wherein: The olefin content in the olefin-containing raw material is 50 to 100% by weight, preferably 80 to 100% by weight; The olefin-containing raw material is a gas phase feed; The olefins in the olefin-containing raw material are C4 to C 12 Olefins, preferably C4-C6 olefins; Preferably, the conditions of the first catalytic conversion reaction include: a reaction temperature of 550 to 800°C, preferably 580 to 750°C, and further preferably 600 to 700°C; a reaction pressure of 0.01 to 1MPa, preferably 0.1 to 0.8MPa, and further preferably 0.2 to 0.5MPa; a reaction time of 0.05 to 10 seconds, preferably 0.1 to 5 seconds, and further preferably 0.2 to 2 seconds; a weight ratio of the first catalyst to the olefin-containing raw material of (5 to 50):1, preferably (10 to 35):1, and further preferably (12 to 30):

1.

7. The method according to claim 1, wherein: The method further comprises: stripping the first catalyst after the reaction, and then transporting the first catalyst to a regenerator for char regeneration.

8. The method according to claim 1, wherein: The method further comprises the following steps: The heavy hydrocarbon oil is brought into contact with the second catalyst in the second reactor to perform a second catalytic conversion reaction, thereby obtaining a reaction oil gas containing light oil and a catalyst to be produced; Then, the reaction oil gas containing light oil is mixed with the first reaction oil gas, and the mixed oil gas is subjected to a first separation to obtain dry gas, light olefins and C5 and above distillate oil, and the C5 and above distillate oil is subjected to a second separation and the separated olefin-rich component is returned to the first reactor; The heavy hydrocarbon oil is selected from one or more of vacuum gas oil, atmospheric gas oil, coker gas oil, deasphalted oil, vacuum residue oil, atmospheric residue oil, heavy aromatic raffinate oil, coal liquefaction oil, oil sand oil and shale oil; The second catalyst is the regenerated first catalyst or a mixed catalyst of the reacted first catalyst and the regenerated first catalyst.

9. The method according to claim 8, wherein: When the temperature of the reaction oil gas containing light oil and the catalyst to be regenerated in the upper area of ​​the reaction zone of the second reactor exceeds 530°C, a cooling zone is set downstream of the reaction zone to cool the reaction oil gas containing light oil and the catalyst to be regenerated to below 530°C.

10. The method according to claim 8, wherein: The method further comprises: stripping the first catalyst and the catalyst to be regenerated respectively, and then burning and regenerating them, and then returning the regenerated catalysts to the first reactor and the second reactor respectively; Preferably, the conditions of the second catalytic conversion reaction include: reaction temperature of 500-600°C, preferably 530-580°C; reaction pressure of 0.01-1MPa, preferably 0.05-1MPa; reaction time of 0.01-100s, preferably 0.1-80s; Preferably, the weight ratio of the second catalyst to the heavy hydrocarbon oil is (1-100):1, preferably (3-70):1.

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