CN115895710A

CN115895710A - Catalytic conversion method and device for producing low-carbon olefin

Info

Publication number: CN115895710A
Application number: CN202111165835.9A
Authority: CN
Inventors: 马文明; 朱根权; 杨超
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2021-09-30
Filing date: 2021-09-30
Publication date: 2023-04-04

Abstract

The invention discloses a catalytic conversion method for producing low-carbon olefin, which comprises the following steps: introducing a first hydrocarbon raw material into a first riser reactor to contact with a catalyst and perform catalytic cracking to obtain a first oil mixture, and separating the first oil mixture to obtain a first oil gas and a first reactant; introducing a second hydrocarbon raw material into a second riser reactor to contact with the catalyst and carry out catalytic cracking to obtain a second oil agent mixture, and separating the second oil agent mixture to obtain a second oil gas and a second spent agent; introducing the first reaction oil gas and the second spent agent into a reverse flow reactor for contact and catalytic cracking to obtain a third oil agent mixture, and separating the third oil agent mixture to obtain a third oil gas and a third spent agent; and carrying out a regeneration reaction on the third spent catalyst to obtain a regenerant, and feeding the regenerant into the first riser reactor and the second riser reactor.

Description

Catalytic conversion method and device for producing low-carbon olefin

Technical Field

The disclosure relates to the field of petrochemical industry, in particular to a catalytic conversion method and a catalytic conversion device for producing low-carbon olefins.

Background

Low-carbon olefins such as ethylene and propylene are basic chemical raw materials, and are mainly derived from steam cracking, catalytic cracking, methanol-to-olefin and alkane dehydrogenation devices at present. With the adoption of new light raw materials in the steam cracking process, the distribution of products will change, for example, ethane is used as the steam cracking raw material, the proportion of ethylene in the products is obviously improved compared with naphtha, and the yield of propylene is reduced. The catalytic cracking process can produce more low-carbon olefins, and is an effective supplementary measure for preparing ethylene by steam thermal cracking. However, the yield of light olefins in the conventional catalytic cracking process is not high, and is not more than 15% of the feedstock, which makes it difficult to meet the market demand, and therefore, it is necessary to develop a catalytic cracking technology capable of processing heavy feedstock and producing more light olefins.

US patent document US5997728 discloses a process for using a large amount of a shape selective cracking aid in the catalytic cracking of heavy feedstocks. Said adjuvant is formed from amorphous matrix and ZSM-5 zeolite 12.40%, and its system storage quantity is at least 10%, so that the ZSM-5 content in the catalyst is greater than 3%. The method can greatly improve the propylene and the butylene without additionally increasing the yield of aromatic hydrocarbon and losing the yield of gasoline.

Chinese patent document CN1031834A discloses a catalytic conversion method for producing low-carbon olefins. The method takes petroleum fractions, residual oil or crude oil with different boiling ranges as raw materials, takes a mixture containing Y zeolite and quinary ring high-silicon zeolite as a catalyst, adopts a fluidized bed or a moving bed as a reactor, and has the following reaction conditions: the temperature is 500-650 ℃, the pressure is 0.15-0.30MPa, and the weight hourly space velocity is 0.2-20 hours ^-1 And the catalyst is in a catalyst-oil ratio of 2-12, and the reacted catalyst is returned to the reactor for recycling after being burnt and regenerated. Compared with the conventional catalytic cracking and steam cracking, the method can obtain more propylene and butylene.

Chinese patent document CN102690683A discloses a catalytic cracking method for producing propylene. The method adopts a double-riser configuration, the first riser reactor is used for treating heavy hydrocarbon oil, a catalyst containing Y-type zeolite and beta-type zeolite is used, the second riser reactor is used for treating light hydrocarbon, and a selective zeolite with the pore diameter less than 0.7nm is used. The method adopts two different catalysts, and divides the stripping zone and the regeneration zone into two independent parts by the partition plates respectively, thereby increasing the complexity of the device and being not beneficial to operation.

Chinese patent document CN102206509a discloses a hydrocarbon catalytic conversion method for producing propylene and light aromatic hydrocarbons. The method adopts a combined reactor form of a double lifting pipe and a fluidized bed reactor, wherein heavy hydrocarbons and a cracking catalyst containing modified beta zeolite are in contact reaction in a first reactor, C4 hydrocarbon fractions and/or light gasoline fractions and the cracking catalyst containing the modified beta zeolite are introduced into a third reactor for continuous reaction after being in contact reaction in a second reactor, and the third reactor is the fluidized bed reactor and creates conditions for secondary cracking reaction of the gasoline fractions, thereby improving the yield of propylene and light aromatic hydrocarbons.

Chinese patent document CN103131464a discloses a hydrocarbon catalytic conversion method for producing propylene and light aromatic hydrocarbons. The method comprises the steps of enabling petroleum hydrocarbon and a catalytic cracking catalyst to react in a contact manner in a lifting pipe, enabling reaction effluent to enter a fluidized bed reactor without separation, enabling the reaction effluent to contact with the introduced catalyst subjected to pore channel modification treatment to carry out oligomerization, cracking and aromatization reactions, separating to obtain a product containing low-carbon olefin and light aromatic hydrocarbon, separating the carbon deposited catalyst into two parts after steam stripping and regeneration, enabling one part of the carbon deposited catalyst to return to the lifting pipe for recycling, sending the other part of the carbon deposited catalyst to a catalyst pore channel modification area, contacting and reacting with a contact agent, and sending the other part of the carbon deposited catalyst to a fluidized bed for use. The method has high heavy oil conversion capacity and high propylene selectivity for heavy hydrocarbon raw materials.

The above technology promotes the conversion of heavy hydrocarbon feedstock and improves the selectivity of low carbon olefins by adjusting the catalyst formulation and adopting a combined reactor form combining a riser and a fluidized bed, but the yield of low carbon olefins still needs to be further improved, and the generation of methane and coke cannot be inhibited.

Disclosure of Invention

The present disclosure aims to provide a catalytic conversion method and apparatus for producing low-carbon olefins, thereby increasing the yield of low-carbon olefins.

In order to achieve the above object, a first aspect of the present disclosure provides a catalytic conversion method for producing lower olefins, the method comprising:

s1, introducing a first hydrocarbon raw material into a first riser reactor to contact with a catalyst and perform a first catalytic cracking reaction to obtain a first oil mixture, and performing first separation on the first oil mixture to obtain a first reaction oil gas and a first catalyst to be generated;

s2, introducing a second hydrocarbon raw material into a second riser reactor to contact with the catalyst and perform a second catalytic cracking reaction to obtain a second oil mixture, and performing second separation on the second oil mixture to obtain a second reaction oil gas and a second spent catalyst;

s3, introducing the first reaction oil gas and the second spent catalyst into a reverse flow reactor to contact with each other and perform a third catalytic cracking reaction to obtain a third oil mixture, and performing third separation on the third oil mixture to obtain a third reaction oil gas and a third spent catalyst;

and S4, introducing the third spent catalyst into a regenerator after steam stripping to perform a regeneration reaction to obtain a regenerated catalyst, and feeding the regenerated catalyst into the first riser reactor and the second riser reactor.

Alternatively, the weight ratio of the second hydrocarbonaceous feedstock to the first hydrocarbonaceous feedstock is in the range of from 0.05 to 0.20, preferably from 0.08 to 0.15; the first hydrocarbon raw material is selected from one or a mixture of more than one of vacuum wax oil, atmospheric residue oil, vacuum residue oil, coking wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil, distillate oil obtained by F-T synthesis and animal and vegetable oil; the second hydrocarbon feedstock is a mixture of C4-C8 hydrocarbons; the mixture of C4-C8 hydrocarbons has an olefin content of more than 50% by weight, preferably more than 60% by weight.

Optionally, the reaction temperature of the first riser reactor is 520-620 ℃, preferably 540-600 ℃; the agent-oil ratio is 2-25, preferably 3-20; the reaction time is from 1 to 15 seconds, preferably from 2 to 10 seconds.

Alternatively, the second riser reactor has a reaction temperature of 560 to 660 ℃, preferably 580 to 640 ℃, a catalyst-to-oil ratio of 3 to 40, preferably 5 to 30, and a reaction time of 0.5 to 10 seconds, preferably 1 to 5 seconds.

Optionally, the first reaction oil gas is introduced into the reverse-flow reactor from the bottom of the reverse-flow reactor, and the second spent catalyst is introduced into the reverse-flow reactor from the top of the reverse-flow reactor; the reaction temperature of the reverse-flow reactor is 540-640 ℃, preferably 560-620 ℃; the density of the catalyst is 50-400kg/m ³ Preferably 150 to 250kg/m ³ (ii) a The oil gas residence time is 0.5-10 seconds, preferably 2-5 seconds.

Optionally, the method further comprises: and carrying out first separation on the first oil mixture through a separation baffle at the top of the first riser reactor to obtain first reaction oil gas and a first catalyst to be regenerated, and introducing the first catalyst to be regenerated into a regenerator after steam stripping to carry out regeneration reaction to obtain a first regenerated catalyst.

Optionally, the method further comprises: and introducing the second oil agent mixture into a quick separation device connected with the tail end of the second riser reactor for second separation to obtain second reaction oil gas and a second spent catalyst, and leading the second reaction oil gas out of a gas collection chamber.

Optionally, the catalyst comprises the MFI structure molecular sieve, clay, and binder; the content of the MFI structure molecular sieve is 20 to 60 wt%, preferably 30 to 50 wt%, based on the total weight of the catalyst; the clay content is 10-70 wt%, preferably 15-45 wt%; the content of the binder is 10-40 wt%, preferably 20-35 wt%; the MFI structure molecular sieve is selected from at least one of ZRP molecular sieves, ZRP molecular sieves containing phosphorus, ZRP molecular sieves containing rare earth, ZRP molecular sieves containing phosphorus and alkaline earth metal and ZRP molecular sieves containing phosphorus and transition metal, and the ZRP zeolite containing phosphorus and rare earth is preferred; the clay is at least one of kaolin, montmorillonite and bentonite; the binder is at least one selected from silica sol, aluminum sol and pseudo-boehmite, and preferably, the binder is double-aluminum binder of the aluminum sol and the pseudo-boehmite.

A second aspect of the present disclosure provides an apparatus of a catalytic conversion method for producing lower olefins, the catalytic conversion apparatus including a combined reactor, a stripper, and a regenerator, the combined reactor including a first riser reactor, a second riser reactor, and a reverse-flow reactor;

the top of the first riser reactor is provided with a first separation device, an oil-gas distributor and a catalyst dipleg; the tail end of the second riser reactor is provided with a second separation device and a catalyst distributor; the stripper is positioned below the reverse-flow reactor and is communicated with the reverse-flow reactor;

the regenerator is connected with the stripper through a spent agent conveying pipe; the regenerator is connected with the first riser reactor and the second riser reactor through a regenerant delivery pipe respectively.

Optionally, the reverse-flow reactor is an equal diameter reactor and/or a variable diameter reactor, and the ratio of the average diameter to the height of the reverse-flow reactor is 1:0.5 to 5, preferably 1:1-3.

Through the technical scheme, the method provided by the disclosure has the advantages that the countercurrent reactor is arranged, so that the first reaction oil gas obtained by the first riser reactor and the second spent catalyst obtained by the second riser reactor are subjected to countercurrent contact reaction, the heavy oil catalytic cracking intermediate product is promoted to be further converted into the low-carbon olefin, and the yield of the low-carbon olefin is improved. The device provided by the present disclosure is used for the separation of the first mixed oil mixture by arranging a separation baffle at the top of the first riser reactor, wherein an oil-gas distributor is arranged for the uniform distribution of the first reaction oil gas, and a catalyst dipleg is arranged for the introduction of the first catalyst to be generated into the stripper. And arranging a quick separation device at the tail end of the second riser reactor for separating the second oil agent mixture, wherein a catalyst distributor is arranged for uniformly distributing the second spent catalyst. The measures can promote the heavy oil catalytic cracking intermediate product to be further converted into the low-carbon olefin, so that the yield of the low-carbon olefin is improved.

Additional features and advantages of the disclosure will be set forth in the detailed description which follows.

Drawings

The accompanying drawings, which are included to provide a further understanding of the disclosure and are incorporated in and constitute a part of this specification, illustrate embodiments of the disclosure and together with the description serve to explain the disclosure without limiting the disclosure. In the drawings:

fig. 1 is a schematic structural view of a catalytic conversion device according to an embodiment of the present disclosure.

Description of the reference numerals

1 first riser reactor 2 second riser reactor

3 counter-current reactor 4 settler

5 stripper 6 regenerator

11 first hydrocarbon feedstock 12 first pre-lift gas line

13 recycle oil 14 separation baffle

15 oil gas distributor 16 catalyst dipleg

21 second hydrocarbon feedstock 22 second pre-lift gas line

23 quick split device 24 catalyst distributor

41 first cyclone separator 42 second cyclone separator

43 first plenum 44 separate system lines

51 spent agent transfer line 52 stripping gas

53 stripping baffle 61 main air

62 second regenerated catalyst line 63 first regenerated catalyst line

64 third cyclone separator 65 fourth cyclone separator

66 second gas collecting chamber 67 regeneration flue gas outlet

Detailed Description

The following describes in detail specific embodiments of the present disclosure. It should be understood that the detailed description and specific examples, while indicating the present disclosure, are given by way of illustration and explanation only, not limitation.

A first aspect of the present disclosure provides a catalytic conversion process for producing lower olefins, the process comprising:

According to the method, the countercurrent reactor is arranged, so that the first reaction oil gas obtained by the first riser reactor and the second spent catalyst obtained by the second riser reactor are subjected to countercurrent contact reaction, the heavy oil catalytic cracking intermediate product is promoted to be further converted into the low-carbon olefin, and the yield of the low-carbon olefin is improved.

According to the present disclosure, the weight ratio of the second hydrocarbonaceous feedstock to the first hydrocarbonaceous feedstock may be from 0.05 to 0.20, preferably from 0.08 to 0.15; the first hydrocarbon raw material can be one or a mixture of more than one of vacuum wax oil, atmospheric residue oil, vacuum residue oil, coking wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil, distillate oil obtained by F-T synthesis and animal and vegetable oil. The first hydrocarbon feedstock undergoes a cracking reaction primarily in the first riser reactor, converting from a macromolecular reactant to a small molecule product.

According to the present disclosure, the second hydrocarbonaceous feedstock can be a mixture of C4-C8 hydrocarbons; the mixture of C4-C8 hydrocarbon refers to low molecular hydrocarbon which takes C4-C8 fraction as main component and exists in gas form at normal temperature and normal pressure, and comprises alkane, alkene and alkyne. The mixture of C4-C8 hydrocarbons comprises a mixture of C4-C8 hydrocarbons produced by the apparatus of the present invention and may also comprise a mixture of C4-C8 hydrocarbons produced by other apparatus processes, with a mixture of C4-C8 hydrocarbons produced by the apparatus of the present invention being preferred. The mixture of C4-C8 hydrocarbons is preferably a mixture of C4-C8 hydrocarbons rich in olefins, which mixture of C4-C8 hydrocarbons may have an olefin content of more than 50 wt.%, preferably more than 60 wt.%.

According to the present disclosure, the reaction temperature of the first riser reactor may be 520 to 620 ℃, preferably 540 to 600 ℃; the agent-oil ratio can be 2-25, preferably 3-20; the reaction time may be 1 to 15 seconds, preferably 2 to 10 seconds.

According to the present disclosure, the reaction temperature of the second riser reactor may be 560 to 660 ℃, preferably 580 to 640 ℃, the agent-to-oil ratio may be 3 to 40, preferably 5 to 30, and the reaction time may be 0.5 to 10 seconds, preferably 1 to 5 seconds.

According to the present disclosure, the first reaction oil gas may be introduced into the reverse-flow reactor from the bottom of the reverse-flow reactor, and the second spent catalyst may be introduced into the reverse-flow reactor from the top of the reverse-flow reactor; the reaction temperature of the reverse-flow reactor may be 540 to 640 ℃, preferably 560 to 620 ℃; the density of the catalyst can be 50-400kg/m ³ Preferably 150 to 250kg/m ³ (ii) a The hydrocarbon residence time may be in the range of 0.5 to 10 seconds, preferably 2 to 5 seconds.

According to the present disclosure, the method may further include: and carrying out first separation on the first oil mixture through a separation baffle at the top of the first riser reactor to obtain first reaction oil gas and a first catalyst to be regenerated, and introducing the first catalyst to be regenerated into a regenerator after steam stripping to obtain a first regenerated catalyst. The separation baffle is arranged at the top of the first riser reactor, so that the first mixed oil agent mixture is separated.

According to the present disclosure, the method may further include: and introducing the second oil agent mixture into a quick separation device connected with the tail end of the second riser reactor for second separation to obtain second reaction oil gas and a second spent catalyst, and leading the second reaction oil gas out of a gas collection chamber.

In the present disclosure, the third separating may include: and separating the third oil mixture in a settler to obtain a third reaction oil gas and a third spent catalyst.

According to the present disclosure, the first hydrocarbon feedstock is preferably a preheated first hydrocarbon feedstock, and further preferably, the temperature of the preheated first hydrocarbon feedstock is 180 to 340 ℃; the second hydrocarbon feedstock is preferably a preheated second hydrocarbon feedstock, and more preferably the preheated second hydrocarbon feedstock has a temperature of from 100 ℃ to 150 ℃.

According to the present disclosure, the catalyst may include the MFI structure molecular sieve, clay, and binder; the MFI structure molecular sieve may be present in an amount of from 20 to 60 wt.%, preferably from 30 to 50 wt.%, based on the total weight of the catalyst; the clay may be present in an amount of 10 to 70 wt%, preferably 15 to 45 wt%; the binder may be present in an amount of 10 to 40 wt%, preferably 20 to 35 wt%; the MFI structure molecular sieve may be at least one selected from the group consisting of ZRP molecular sieves, ZRP molecular sieves containing phosphorus, ZRP molecular sieves containing rare earth, ZRP molecular sieves containing phosphorus and alkaline earth metal, and ZRP molecular sieves containing phosphorus and transition metal, preferably ZRP zeolite containing phosphorus and rare earth; the clay can be at least one selected from kaolin, montmorillonite and bentonite; the binder may be at least one selected from the group consisting of silica sol, aluminum sol, and pseudo-boehmite, and preferably, the binder is a double aluminum binder of aluminum sol and pseudo-boehmite.

In a specific embodiment of the present disclosure, as shown in fig. 1, after the first hydrocarbon feedstock 11 is preheated to 180-340 ℃, it is sprayed into the first riser reactor 1 through a nozzle to perform a first catalytic cracking reaction with the regenerated catalyst entering the bottom of the first riser reactor 1 through a first regenerated catalyst pipeline 63, so as to obtain a first oil mixture. The first oil mixture is separated by a separation baffle 14 at the top of the first riser reactor 1, so that the first reaction oil gas is separated from most or all of the first catalyst to be generated, the separated first catalyst to be generated is introduced into the stripper 5 by a catalyst dipleg 16, and the first reaction oil gas is introduced into the reverse flow reactor 3 by an oil gas distributor 15. After the second hydrocarbon raw material 21 is preheated to 100-150 ℃, the second hydrocarbon raw material is sprayed into the second riser reactor 2 through a nozzle and carries out a second catalytic cracking reaction with the regenerated catalyst entering the bottom of the second riser reactor 2 through a second regenerated catalyst pipeline 62, so as to obtain a second oil mixture. The second oil mixture is separated by the fast separation device 23 connected with the end of the second riser reactor 2, so that the second reaction oil gas is separated from most or all of the second spent catalyst, the separated second reaction oil gas is led out of the device by the gas collection chamber 43, and the obtained second spent catalyst is led into the reverse flow reactor 3 by the catalyst distributor 24. The first reaction oil gas and the second spent catalyst are contacted in the reverse flow reactor 3 and undergo a third catalytic cracking reaction to obtain a third oil mixture, the third oil mixture is separated in the settler 4, the obtained third reaction oil gas is led out of the device through a separation system pipeline 44, the third spent catalyst is led into the stripper 5, and the stripped spent catalyst is led into the regenerator 6 through a spent catalyst conveying pipeline 51 for regeneration and then is recycled. The third reaction oil gas enters the subsequent product separation system through separation system line 44. The catalytic cracking products are separated into products such as dry gas, cracked gas, gasoline, light oil, slurry oil and the like in the product separation system. The product separation system may be any of a variety of separation systems known in the art, and the disclosure is not particularly limited. The cracked gas can be separated and refined in subsequent products to obtain a mixture of polymer-grade propylene products and C4-C8 hydrocarbons. The mixture of C4-C8 hydrocarbons may be partially or completely returned to the second riser reactor 2 for reaction. The spent catalyst separated by the first cyclone separator 41 and the second cyclone separator 42 enters the stripper 5 for stripping. The stripping steam in the stripper 5 can directly enter the settler 4, and the reaction oil gas is separated by the first cyclone 41 and the second cyclone 42 together with other oil gas and then is led out of the reactor through a separation system pipeline 44. The catalyst stripped in the stripper 5 enters the regenerator 6 for coke burning regeneration, and the regenerated flue gas is led out from a top gas collection chamber 66 of the regenerator 6 through a regenerated flue gas outlet 67. The regenerated catalyst is returned to the pre-lifting section of the first riser reactor 1 and the second riser reactor 2 respectively through a first regenerated catalyst pipeline 63 and a second regenerated catalyst pipeline 62 for recycling. The mode of operation and operating conditions of the regenerator described in this disclosure may be with reference to a conventional catalytic cracking regenerator.

According to a specific embodiment of the present disclosure, a lift gas is introduced to the first riser reactor 1 and the second riser reactor 2 through the first pre-lift gas line 12 and the second pre-lift gas line 22, respectively. The lift gas is well known to those skilled in the art and may be selected from one or more of steam, nitrogen, dry gas, preferably steam.

the regenerator is connected with the stripper through a spent agent conveying pipe; the regenerator is respectively connected with the first riser reactor and the second riser reactor through a regenerant conveying pipe.

According to the present disclosure, the reverse-flow reactor may be an equal diameter reactor and/or a variable diameter reactor, and the ratio of the average diameter to the height of the reverse-flow reactor may be 1:0.5 to 5, preferably 1:1-3.

In a specific embodiment of the present disclosure, the catalytic conversion apparatus includes a first riser reactor 1, a second riser reactor 2, a reverse flow reactor 3, a settler 4 and a stripper 5; a separation clapboard 14, an oil-gas distributor 15 and a catalyst dipleg 16 are arranged in the first riser reactor 1; a fast separation device 23 and a catalyst distributor 24 are arranged in the second riser reactor 2; the precipitator 4 is provided with a first cyclone separator 41 and a second cyclone separator 42, the inlets of the first cyclone separator 41 and the second cyclone separator 42 are positioned at the upper part of the precipitator 4, the spent agent outlets of the first cyclone separator 41 and the second cyclone separator 42 are connected with the spent agent inlet of the stripper 5, and the oil gas outlet of the second cyclone separator 42 is communicated with an oil gas separation system.

According to a specific embodiment of the present disclosure, the catalytic conversion apparatus further comprises a regenerator 6, and the regenerator 6 delivers the regenerated catalyst to the bottoms of the second riser reactor 2 and the first riser reactor 1 through a second regenerated catalyst line 62 and a first regenerated catalyst line 63, respectively. Wherein the catalyst delivery rate can be adjusted by means of a valve in the catalyst line.

According to a specific embodiment of the present disclosure, the first oil mixture led out from the outlet of the first riser reactor 1 is separated by a separation partition 14 to obtain a first reaction oil gas and a first catalyst to be generated, the first catalyst to be generated is led into the stripper 5 through a catalyst dipleg 16, and the first reaction oil gas is led into the reverse flow reactor 3 through an oil gas distributor 15. And the second oil mixture led out from the outlet of the second riser reactor 2 is separated by a fast separation device 23 to obtain second reaction oil gas and a second spent catalyst, the second reaction oil gas is led out of the device by a first gas collection chamber 43, and the second spent catalyst is led into the reverse flow reactor 3 by a catalyst distributor 24. And carrying out contact reaction on the first reaction oil gas and the second spent catalyst in the reverse-flow reactor 3. Wherein the first reaction oil gas flows upwards, and the second spent catalyst flows downwards.

The present disclosure is further illustrated by the following examples. The raw materials used in the examples are all available from commercial sources.

In the examples and comparative examples of the present disclosure, the gas product was tested using the petrochemical analysis method RIPP77-90, the coke content was determined using the petrochemical analysis method RIPP 107-90, the organic liquid product composition was determined using the SH/T0558-1993, the cut points of the gasoline and diesel were 221 ℃ and 343 ℃, respectively, and the light aromatics in the gasoline were determined using the petrochemical analysis method RIPP 82-90.

In the following examples, the conversion of the feedstock oil and the yield of cracked products were calculated according to the following formulas:

the RIPP petrochemical analysis method used in the invention is selected from the editions of petrochemical analysis method (RIPP test method), yang Cui, and the like, science publishing company, 1990.

The reagents used below are all chemically pure reagents, unless otherwise specified.

The MFI structure molecular sieve is produced by Qilu catalyst factories and has the industrial grades as follows:

ZRP-1: wherein SiO is ₂ /Al ₂ O ₃ ＝30，Na ₂ O content 0.17 wt%, rare earth oxide RE ₂ O ₃ Is 1.4 wt%, wherein the lanthanum oxide is 0.84 wt%, cerium oxide is 0.18 wt%, and the other rare earth oxides are 0.38 wt%.

The catalysts used in the examples and comparative examples were self-made catalysts, designated as CAT, having an active component of ZRP molecular sieve with specific properties as shown in Table 1.

The specific preparation process of the catalyst CAT comprises the following steps:

uniformly mixing the ZRP molecular sieve, adding deionized water, pulping, and uniformly stirring to obtain molecular sieve slurry with the solid content of 20-40 wt%; mixing clay, binder and deionized water, pulping, and stirring to obtain carrier slurry with solid content of 15-25 wt%; and mixing and pulping the homogenized molecular sieve slurry and the homogenized carrier slurry, and then sequentially carrying out spray drying, washing, filtering and drying to obtain the catalyst CAT. The catalyst CAT was aged at 790 ℃ under 100% steam for 14 hours prior to testing.

TABLE 1

Examples 1-2 and comparative examples 1-2 in this disclosure are used to illustrate that the counter-current reaction can promote the intermediate product of the heavy oil catalytic cracking reaction to be further converted into lower olefins such as ethylene and propylene.

The feedstock used in examples 1-2 and comparative examples 1-2 was a light gasoline fraction produced by a catalytic cracking unit, and the specific properties are shown in Table 2.

TABLE 2

Examples 1 to 2

The tests of examples 1-2 were carried out on a counter-current reaction pilot plant. Wherein, the internal diameter of the reverse-flow reactor is 64mm, and the height is 500mm. Introducing light gasoline into the bottom of the reverse flow reactor, introducing a catalyst CAT into the top of the reverse flow reactor, carrying out a reaction after the catalyst CAT and the reverse flow reactor are in countercurrent contact, separating an oil agent mixture after the reaction through a cyclone separator to obtain a spent catalyst and reaction oil gas, introducing the spent catalyst into a stripper and then into a regenerator for regeneration to obtain a regenerated catalyst, returning the regenerated catalyst to the riser reactor for recycling, and introducing the reaction oil gas into a fractionation system for separation. The reaction conditions and results are shown in Table 3.

Comparative examples 1 to 2

The tests of comparative examples 1-2 were carried out on a pilot plant of the fluidized bed reaction type. The fluidized bed reactor had an inner diameter of 64mm and a height of 500mm. Light gasoline and catalyst CAT are simultaneously introduced into a fluidized bed reactor for reaction, an oil agent mixture after the reaction is separated through a cyclone separator to obtain a spent catalyst and reaction oil gas, the spent catalyst enters a stripper and then enters a regenerator for regeneration to obtain a regenerated catalyst, the regenerated catalyst is returned to a riser reactor for recycling, and the reaction oil gas is introduced into a fractionation system for separation. The reaction conditions and results are shown in Table 3.

TABLE 3

Examples 3 to 4 and comparative examples 3 to 4 are provided to illustrate that the embodiment shown in fig. 1 can improve the heavy oil conversion rate and the low carbon olefin yield.

The starting materials used in examples 3-4 and comparative examples 3-4 were wax oils, the specific properties of which are shown in Table 4.

TABLE 4

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Example 3

The test of this example was carried out on a medium-sized test apparatus as shown in FIG. 1. The apparatus comprises two riser reactors and a reverse-flow reactor. The first riser reactor 1 has an inner diameter of 16mm and a length of 3200mm, the second riser reactor 2 has an inner diameter of 16mm and a height of 3800mm, and the reverse flow reactor 3 has an inner diameter of 64mm and a height of 500mm.

Introducing wax oil into the bottom of a first riser reactor 1, contacting with a regenerated catalyst CAT from a regenerator 6, and carrying out a first catalytic cracking reaction to obtain a first oil mixture, separating the first oil mixture by a separation baffle to obtain first reaction oil gas and a first catalyst to be regenerated, and introducing the obtained first reaction oil gas into a reverse flow reactor 3;

introducing C4 hydrocarbon into the bottom of the second riser reactor 2, contacting with the regenerated catalyst CAT from the regenerator 6, and carrying out a second catalytic cracking reaction to obtain a second oil mixture, separating the second oil mixture by a fast separation device to obtain second reaction oil gas and a second spent catalyst, and introducing the obtained second spent catalyst into the reverse flow reactor 3;

and (3) enabling the first reaction oil gas from the first riser reactor 1 and the second spent catalyst from the second riser reactor 2 to contact in a reverse-flow reactor 3 and carrying out a third catalytic cracking reaction to obtain a third oil mixture, and separating the third oil mixture through a cyclone separator to obtain a third reaction oil gas and a third spent catalyst. And feeding the third spent catalyst into a stripper 5, feeding the third spent catalyst into a regenerator 6 for a regeneration reaction to obtain a regenerated catalyst, returning the regenerated catalyst to the first riser reactor and the second riser reactor for recycling, and introducing oil gas of the third reaction into a fractionation system for separation.

Wherein the mass ratio of the C4 hydrocarbon to the wax oil is 0.08:1. the reaction conditions and results are shown in Table 5.

Example 4

The method of this example is the same as example 3, except that: in addition to the C4 fraction being introduced into the second riser reactor 2, a light gasoline fraction (distillation range 40-80 ℃, olefin content 65 wt%) is also introduced into the second riser reactor 2.

Wherein the mass ratio of the C4 hydrocarbon to the light gasoline fraction to the wax oil is 0.05:0.05:1. the reaction conditions and results are shown in Table 5.

Comparative example 3

The test of this comparative example was performed on a medium-sized test apparatus. The apparatus comprises a riser reactor and a fluidized bed reactor. The riser reactor 1 has an inner diameter of 16mm and a length of 3200mm, and the fluidized bed reactor 2 has an inner diameter of 64mm and a height of 500mm.

Wax oil is introduced into the bottom of a riser reactor 1, contacts with a regenerated catalyst CAT from a regenerator 6 and undergoes a catalytic cracking reaction to obtain a first oil mixture, the first oil mixture is introduced into a fluidized bed reactor 2 to continue to react to obtain a second oil mixture, the second oil mixture is separated by a cyclone separator to obtain reaction oil gas and a spent catalyst, the spent catalyst enters a stripper 5 and then enters the regenerator 6 for regeneration, the regenerated catalyst returns to the riser reactor for recycling, and the reaction oil gas is introduced into a fractionation system for separation. The reaction conditions and results are shown in Table 5.

Comparative example 4

The test of this comparative example was conducted on a medium-sized test apparatus. The apparatus comprises two riser reactors and a fluidized bed reactor. The inner diameter of the first riser reactor 1 was 16mm, the length was 3200mm, the inner diameter of the second riser reactor 2 was 16mm, the height was 3800mm, and the inner diameter of the fluidized bed reactor 3 was 64mm, and the height was 500mm.

Introducing wax oil into the bottom of a first riser reactor 1, contacting with a regenerated catalyst CAT from a regenerator 6, and carrying out a first catalytic cracking reaction to obtain a first oil mixture, and introducing the first oil mixture into a fluidized bed reactor 3; introducing C4 hydrocarbon into the bottom of the second riser reactor 2, contacting with a regenerated catalyst CAT from a regenerator 6, and carrying out a second catalytic cracking reaction to obtain a second oil mixture, and introducing the second oil mixture into the fluidized bed reactor 3; and (2) carrying out a third catalytic cracking reaction on the first oil mixture from the first riser reactor 1 and the second oil mixture from the second riser reactor 2 in a fluidized bed reactor 3 to obtain a third oil mixture, separating the third oil mixture by a cyclone separator to obtain reaction oil gas and a spent catalyst, introducing the spent catalyst into a stripper 5 and then into a regenerator 6 for regeneration to obtain a regenerated catalyst, returning the regenerated catalyst to the first riser reactor and the second riser reactor for recycling, and introducing the reaction oil gas into a fractionation system for separation.

TABLE 5

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As can be seen from tables 3 and 5, higher hydrocarbon conversion capacity and higher yield of lower olefins can be achieved by using the method and apparatus provided by the present disclosure.

The preferred embodiments of the present disclosure have been described in detail above, however, the present disclosure is not limited to the specific details of the above embodiments, and various simple modifications may be made to the technical solution of the present disclosure within the technical idea of the present disclosure, and these simple modifications all fall within the protection scope of the present disclosure.

It should be noted that the various features described in the above embodiments may be combined in any suitable manner without departing from the scope of the invention. In order to avoid unnecessary repetition, various possible combinations will not be separately described in this disclosure.

In addition, any combination of various embodiments of the present disclosure may be made, and the same should be considered as the disclosure of the present disclosure, as long as it does not depart from the spirit of the present disclosure.

Claims

1. A catalytic conversion process for producing lower olefins, the process comprising:

and S4, introducing the third spent catalyst into a regenerator after steam stripping for regeneration reaction to obtain a regenerated catalyst, and feeding the regenerated catalyst into the first riser reactor and the second riser reactor.

2. A catalytic conversion process according to claim 1,

the weight ratio of the second hydrocarbonaceous feedstock to the first hydrocarbonaceous feedstock is in the range from 0.05 to 0.20, preferably from 0.08 to 0.15;

the first hydrocarbon raw material is selected from one or a mixture of more than one of vacuum wax oil, atmospheric residue oil, vacuum residue oil, coking wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil, distillate oil obtained by F-T synthesis and animal and vegetable oil;

the second hydrocarbon feedstock is a mixture of C4-C8 hydrocarbons; the mixture of C4-C8 hydrocarbons has an olefin content of more than 50% by weight, preferably more than 60% by weight.

3. A catalytic conversion process according to claim 1,

the reaction temperature of the first riser reactor is 520-620 ℃, and preferably 540-600 ℃; the agent-oil ratio is 2-25, preferably 3-20; the reaction time is from 1 to 15 seconds, preferably from 2 to 10 seconds.

4. A catalytic conversion process according to claim 1,

the reaction temperature of the second riser reactor is 560-660 ℃, preferably 580-640 ℃, the agent-oil ratio is 3-40, preferably 5-30, and the reaction time is 0.5-10 seconds, preferably 1-5 seconds.

5. A catalytic conversion process according to claim 1,

the first reaction oil gas is introduced into the reverse-flow reactor from the bottom of the reverse-flow reactor, and the second spent catalyst is introduced into the reverse-flow reactor from the top of the reverse-flow reactor;

the reaction temperature of the reverse-flow reactor is 540-640 ℃, preferably 560-620 ℃; the density of the catalyst is 50-400kg/m ³ Preferably 150 to 250kg/m ³ (ii) a The oil gas residence time is 0.5-10 seconds, preferably 2-5 seconds.

6. A catalytic conversion process according to claim 1, wherein the process further comprises:

and carrying out first separation on the first oil mixture through a separation baffle at the top of the first riser reactor to obtain first reaction oil gas and a first catalyst to be regenerated, and introducing the first catalyst to be regenerated into a regenerator after steam stripping to obtain a first regenerated catalyst.

7. A catalytic conversion process according to claim 1, wherein the process further comprises:

and introducing the second oil agent mixture into a quick separation device connected with the tail end of the second riser reactor for second separation to obtain second reaction oil gas and a second spent catalyst, and leading the second reaction oil gas out of a gas collection chamber.

8. A catalytic conversion process according to claim 1,

the catalyst comprises the MFI structure molecular sieve, clay and a binder;

the content of the MFI structure molecular sieve is 20 to 60 wt%, preferably 30 to 50 wt%, based on the total weight of the catalyst; the clay content is 10-70 wt%, preferably 15-45 wt%; the content of the binder is 10-40 wt%, preferably 20-35 wt%;

the MFI structure molecular sieve is selected from at least one of ZRP molecular sieves, ZRP molecular sieves containing phosphorus, ZRP molecular sieves containing rare earth, ZRP molecular sieves containing phosphorus and alkaline earth metal and ZRP molecular sieves containing phosphorus and transition metal, and the ZRP zeolite containing phosphorus and rare earth is preferred;

the clay is at least one of kaolin, montmorillonite and bentonite;

the binder is at least one selected from silica sol, aluminum sol and pseudo-boehmite, and preferably, the binder is double-aluminum binder of the aluminum sol and the pseudo-boehmite.

9. An apparatus suitable for the catalytic conversion process for producing lower olefins according to any of claims 1 to 8, wherein the catalytic conversion apparatus comprises a combined reactor, a stripper and a regenerator, the combined reactor comprises a first riser reactor, a second riser reactor and a reverse-flow reactor;

10. The apparatus of claim 9, wherein,

the countercurrent reactor is an equal-diameter reactor and/or a variable-diameter reactor, and the ratio of the average diameter to the height of the countercurrent reactor is 1:0.5 to 5, preferably 1:1-3.