CN115895710B

CN115895710B - Catalytic conversion method and device for producing low-carbon olefin

Info

Publication number: CN115895710B
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: 2024-10-11
Anticipated expiration: 2041-09-30
Also published as: CN115895710A

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 first oil gas and a first to-be-produced agent; introducing a second hydrocarbon raw material into a second riser reactor to contact with a catalyst and perform catalytic cracking to obtain a second oil mixture, and separating the second oil mixture to obtain second oil gas and a second spent agent; introducing the first reaction oil gas and the second spent agent into a countercurrent reactor for contact and catalytic cracking to obtain a third oil agent mixture, and separating the third oil agent mixture to obtain third oil gas and a third spent agent; the third spent catalyst is subjected to a regeneration reaction to obtain a regenerant, and the regenerant is sent into the first riser reactor and the second riser reactor.

Description

Catalytic conversion method and device for producing low-carbon olefin

Technical Field

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

Background

The low-carbon olefin such as ethylene, propylene and the like is a basic chemical raw material, and is mainly sourced from steam cracking, catalytic cracking, olefin preparation from methanol and alkane dehydrogenation devices at present. With the adoption of new light raw materials for steam cracking, the distribution of products is changed, for example, ethane is adopted as the steam cracking raw materials, 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 olefin, and is an effective supplementary measure for preparing ethylene by steam thermal cracking. However, the conventional catalytic cracking process has a low yield of light olefins, which is not more than 15% of the feedstock, and is difficult to meet the market demand, so that it is required to develop a catalytic cracking technology capable of treating heavy feedstock and producing light olefins in a high yield.

A method for the extensive use of shape selective cracking aids in the catalytic cracking of heavy feedstocks is disclosed in US 5997728. The auxiliary agent consists of an amorphous matrix and ZSM-5 zeolite added into the amorphous matrix, wherein the system reserve is at least 10%, so that the proportion of ZSM-5 in the catalyst exceeds 3%. The method can greatly improve propylene and butylene without additionally increasing the aromatic hydrocarbon yield and losing the gasoline yield.

Chinese patent document CN1031834a discloses a catalytic conversion process 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 pentasil 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, the weight hourly space velocity is 0.2-20 hours ^-1, the catalyst-oil ratio is 2-12, and the catalyst after the reaction returns to the reactor for recycling after being burnt and regenerated. The present process is capable of producing more propylene and butene than conventional catalytic cracking and steam cracking.

Chinese patent document CN102690683a discloses a catalytic cracking process for producing propylene. The method adopts a double-riser configuration, wherein 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 shape selective zeolite with the pore diameter smaller than 0.7nm is used. The method adopts two different catalysts, and divides the stripping zone and the regeneration zone into two independent parts through the partition plates respectively, thereby increasing the complexity of the device and being unfavorable for operation.

Chinese patent document CN102206509a discloses a hydrocarbon catalytic conversion process for producing propylene and light aromatic hydrocarbons. The method adopts a combined reactor form of double lifting pipes and a fluidized bed reactor, wherein heavy hydrocarbon and a cracking catalyst containing modified beta zeolite are in contact reaction in a first reactor, C4 hydrocarbon fraction and/or light gasoline fraction and the cracking catalyst containing 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, thereby creating conditions for secondary cracking reaction of the gasoline fraction and further improving the yield of propylene and light aromatic hydrocarbon.

Chinese patent document CN103131464a discloses a hydrocarbon catalytic conversion process for producing propylene and light aromatic hydrocarbons. The method comprises the steps of carrying out contact reaction on petroleum hydrocarbon and a catalytic cracking catalyst in a riser, enabling reaction effluent to enter a fluidized bed reactor without separation, carrying out oligomerization, cracking and aromatization reactions by contacting the reaction effluent with the catalyst which is introduced and is subjected to pore channel modification treatment, separating to obtain products comprising low-carbon olefin and light aromatic hydrocarbon, and separating the products into two parts after stripping and regeneration, wherein one part of the carbon deposition catalyst is recycled in the riser, and the other part of the carbon deposition catalyst is firstly sent to a catalyst pore channel modification area, then is contacted and reacted with a contact agent and is sent to a fluidized bed for use. The method has higher heavy oil conversion capability and high propylene selectivity to heavy hydrocarbon raw materials.

The above technology promotes the conversion of heavy hydrocarbon feedstock and improves the selectivity of lower olefins by adjusting the catalyst formulation and adopting a combined reactor form of riser combined with fluidized bed, but the yield of lower olefins still remains to be further improved and the formation of methane and coke cannot be suppressed.

Disclosure of Invention

The present disclosure is directed to a catalytic conversion method and apparatus for producing light olefins, thereby improving the yield of light olefins.

To achieve the above object, a first aspect of the present disclosure provides a catalytic conversion method for producing light 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 first reaction oil gas and a first catalyst to be regenerated;

S2, introducing a second hydrocarbon raw material into a second riser reactor to contact with a 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 second reaction oil gas and a second spent catalyst;

S3, introducing the first reaction oil gas and the second spent catalyst into a countercurrent reactor for contact and performing a third catalytic cracking reaction to obtain a third oil agent mixture, and performing third separation on the third oil agent mixture to obtain third reaction oil gas and a third spent catalyst;

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

Optionally, the weight ratio of the second hydrocarbon feedstock to the first hydrocarbon feedstock is from 0.05 to 0.20:1, preferably from 0.08 to 0.15:1; the first hydrocarbon raw material is selected from one or more than one mixture of vacuum wax oil, normal pressure residual oil, vacuum residual oil, coking wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil, distillate oil obtained through 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 ratio of the agent to the oil is 2-25, preferably 3-20; the reaction time is 1 to 15 seconds, preferably 2 to 10 seconds.

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

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

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

Optionally, the method further comprises: and introducing the second oil mixture into a quick-separating 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 a binder; the MFI structure molecular sieve is present in an amount of 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 binder content is 10-40 wt%, preferably 20-35 wt%; the MFI structure molecular sieve is at least one selected from ZRP molecular sieves, phosphorus-containing ZRP molecular sieves, rare earth-containing ZRP molecular sieves, phosphorus-containing ZRP molecular sieves and alkaline earth metal-containing ZRP molecular sieves and phosphorus-containing ZRP molecular sieves and transition metal-containing ZRP zeolite, preferably phosphorus-containing ZRP zeolite and rare earth zeolite; the clay is at least one selected from kaolin, montmorillonite and bentonite; the binder is at least one selected from silica sol, alumina sol and pseudo-boehmite, preferably the binder is a double-alumina binder of alumina sol and pseudo-boehmite.

A second aspect of the present disclosure provides an apparatus for a catalytic conversion process for producing light olefins, the catalytic conversion apparatus comprising a combined reactor, a stripper and a regenerator, the combined reactor comprising a first riser reactor, a second riser reactor and a counter-current 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 countercurrent reactor and is communicated with the countercurrent reactor;

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.

Optionally, 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 1:0.5-5, preferably 1:1-3.

According to the technical scheme, 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 in countercurrent contact reaction, the heavy oil catalytic cracking intermediate product is further converted into the low-carbon olefin, and the yield of the low-carbon olefin is improved. The device provided by the disclosure is used for separating a first mixed oil mixture by arranging a separation baffle at the top of a first riser reactor, wherein an oil gas distributor is arranged for uniformly distributing first reaction oil gas, and a catalyst dipleg is arranged for introducing a first spent catalyst into a stripper. And a quick separating device is arranged at the tail end of the second riser reactor and used for separating the second oil 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 present disclosure will be set forth in the detailed description which follows.

Drawings

The accompanying drawings are included to provide a further understanding of the disclosure, and are incorporated in and constitute a part of this specification, illustrate the disclosure and together with the description serve to explain, but do not limit 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 Countercurrent reactor 4 settler

5 Stripper 6 regenerator

First hydrocarbon feedstock 12 first pre-lift gas line

13 Cycle oil 14 separation baffle

15 Oil gas distributor 16 catalyst dipleg

Second hydrocarbon feedstock 21 second pre-lift gas line 22

24 Catalyst distributor of 23 quick separation device

41 First cyclone 42 second cyclone

43 First plenum 44 split system line

51 Spent feed line 52 stripping gas

53 Stripping baffle 61 windward

62 Second regenerated catalyst line 63 first regenerated catalyst line

64 Third cyclone 65 fourth cyclone

66 Second plenum 67 regeneration flue gas outlet

Detailed Description

The following describes specific embodiments of the present disclosure in detail. It should be understood that the detailed description and specific examples, while indicating and illustrating the disclosure, are not intended to limit the disclosure.

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

According to the method, the countercurrent contact reaction is carried out between the first reaction oil gas obtained by the first riser reactor and the second spent catalyst obtained by the second riser reactor through the countercurrent reactor, so that the heavy oil catalytic cracking intermediate product is 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 hydrocarbon feedstock to the first hydrocarbon feedstock may be from 0.05 to 0.20:1, preferably from 0.08 to 0.15:1; the first hydrocarbon raw material can be selected from one or more than one mixture of vacuum wax oil, atmospheric residuum, vacuum residuum, 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 is primarily subjected to a cracking reaction in the first riser reactor from macromolecular reactants to small molecular products.

According to the present disclosure, the second hydrocarbon feedstock may 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 the form of gas at normal temperature and normal pressure, and the low molecular hydrocarbon comprises alkane, alkene and alkyne. The mixture of C4-C8 hydrocarbons includes mixtures of C4-C8 hydrocarbons produced by the apparatus of the invention, and may also include mixtures of C4-C8 hydrocarbons produced by other apparatus processes, with mixtures of C4-C8 hydrocarbons produced by the apparatus of the invention being preferred. The mixture of C4-C8 hydrocarbons is preferably a mixture of olefin-rich C4-C8 hydrocarbons, and the content of olefins in the mixture of C4-C8 hydrocarbons may be greater than 50 wt.%, preferably greater than 60 wt.%.

According to the present disclosure, the reaction temperature of the first riser reactor may be 520-620 ℃, preferably 540-600 ℃; the ratio of the agent to the oil 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 ℃, and 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 countercurrent reactor from the bottom of the countercurrent reactor, and the second spent catalyst may be introduced into the countercurrent reactor from the top of the countercurrent reactor; the reaction temperature of the countercurrent reactor may be 540-640 ℃, preferably 560-620 ℃; the catalyst density may be 50-400kg/m ³, preferably 150-250kg/m ³; the hydrocarbon residence time may be from 0.5 to 10 seconds, preferably from 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 plate at the top of the first riser reactor to obtain first reaction oil gas and a first to-be-regenerated catalyst, and introducing the first to-be-regenerated catalyst into a regenerator for regeneration reaction after steam stripping to obtain the first regenerated catalyst. The present disclosure provides for the separation of a first mixed oil mixture by providing a separation baffle at the top of a first riser reactor.

According to the present disclosure, the method may further include: and introducing the second oil mixture into a quick-separating 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 agent mixture in a settler to obtain 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 180-340 ℃; the second hydrocarbon feedstock is preferably a preheated second hydrocarbon feedstock, and more preferably the temperature of the preheated second hydrocarbon feedstock is from 100 to 150 ℃.

According to the present disclosure, the catalyst may include the MFI structure molecular sieve, clay, and a binder; the MFI structure molecular sieve may be present in an amount of 20 to 60 wt%, preferably 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 a ZRP molecular sieve, a phosphorous-containing ZRP molecular sieve, a rare earth-containing ZRP molecular sieve, a phosphorous-and alkaline earth-containing ZRP molecular sieve, and a phosphorous-and transition metal-containing ZRP molecular sieve, preferably a phosphorous-and rare earth-containing ZRP zeolite; the clay can be at least one selected from kaolin, montmorillonite and bentonite; the binder may be selected from at least one of silica sol, alumina sol and pseudo-boehmite, preferably the binder is a double alumina binder of alumina sol and pseudo-boehmite.

In one 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 a regenerated catalyst entering the bottom of the first riser reactor 1 through a first regenerated catalyst line 63, thereby obtaining 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 spent catalyst, the separated first spent catalyst is introduced into the stripper 5 through a catalyst dipleg 16, and the first reaction oil gas is introduced into the countercurrent reactor 3 through an oil gas distributor 15. After the second hydrocarbon feedstock 21 is preheated to 100-150 ℃, it is sprayed into the second riser reactor 2 through a nozzle and subjected to a second catalytic cracking reaction with a regenerated catalyst entering the bottom of the second riser reactor 2 through a second regenerated catalyst line 62, to obtain a second oil mixture. The second oil mixture is separated by the rapid separation device 23 connected with the tail 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 countercurrent reactor 3 by the catalyst distributor 24. The first reaction oil gas and the second spent catalyst are contacted in a countercurrent reactor 3 and subjected to a third catalytic cracking reaction to obtain a third oil agent mixture, the third oil agent mixture is separated in a 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 a stripper 5, and the stripped spent catalyst is led into a regenerator 6 for recycling after being regenerated through a spent agent conveying pipeline 51. The third reaction oil and gas enters the subsequent product separation system through separation system line 44. In the product separation system, the catalytic cracking products are separated into dry gas, cracked gas, gasoline, light oil, slurry oil and other products. The product separation system may be any of a variety of separation systems known in the art, and there are no particular requirements for the present disclosure. The cracked gas can be separated and refined to obtain a mixture of a polymerization grade propylene product and C4-C8 hydrocarbon. The mixture of C4-C8 hydrocarbons may be partially or fully returned to the second riser reactor 2 for reaction. The spent agent separated by the first cyclone 41 and the second cyclone 42 enters the stripper 5 for stripping. The stripping steam in the stripper 5 can directly enter the settler 4, and after being separated together with other oil and gas by the first cyclone 41 and the second cyclone 42, the reaction oil and gas is led out of the reactor through a separation system line 44. The catalyst stripped in the stripper 5 enters the regenerator 6 for 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-riser sections of the first riser reactor 1 and the second riser reactor 2 for recycling via the first regenerated catalyst line 63 and the second regenerated catalyst line 62, respectively. The manner of operation and operating conditions of the regenerator described in this disclosure may be referenced to a conventional catalytic cracking regenerator.

According to a specific embodiment of the present disclosure, lift gas is introduced into 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.

According to the present disclosure, the countercurrent 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 countercurrent reactor may be 1:0.5-5, preferably 1:1-3.

In one specific embodiment of the present disclosure, the catalytic conversion device comprises a first riser reactor 1, a second riser reactor 2, a countercurrent reactor 3, a settler 4 and a stripper 5; the first riser reactor 1 is provided with a separation baffle 14, an oil-gas distributor 15 and a catalyst dipleg 16; a fast separating device 23 and a catalyst distributor 24 are arranged in the second riser reactor 2; the settler 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 settler 4, the positions of 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 the oil-gas separation system.

According to a specific embodiment of the present disclosure, the catalytic conversion device further comprises a regenerator 6, the regenerator 6 delivering regenerated catalyst to the bottoms of the second riser reactor 2 and the first riser reactor 1 via a second regenerated catalyst line 62 and a first regenerated catalyst line 63, respectively. Wherein the catalyst delivery rate can be regulated by a valve on 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 spent catalyst, the first spent catalyst is introduced into the stripper 5 through a catalyst dipleg 16, and the first reaction oil gas is introduced into the countercurrent reactor 3 through an oil gas distributor 15. The second oil agent mixture led out from the outlet of the second riser reactor 2 is separated by a quick separating device 23 to obtain second reaction oil gas and second spent catalyst, the second reaction oil gas is led out from the first gas collection chamber 43, and the second spent catalyst is led into the countercurrent reactor 3 by a catalyst distributor 24. The first reaction oil gas and the second spent catalyst are contacted and reacted in the countercurrent reactor 3. Wherein the first reaction oil gas flows upwards, and the second spent catalyst flows downwards.

The present disclosure is further illustrated in detail by the following examples. The starting materials used in the examples are all available commercially.

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

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

The RIPP petrochemical analysis method used in the present invention is selected from the group consisting of "petrochemical analysis method (RIPP test method)", code Yang Cuiding, et al, science Press, 1990.

The reagents used hereinafter are all chemically pure reagents unless otherwise specified.

The MFI structure molecular sieves are all produced by Qilu catalyst factories and have the industrial marks:

ZRP-1: wherein the content of SiO ₂/Al₂O₃＝30,Na₂ O is 0.17 wt%, the content of rare earth oxide RE ₂O₃ is 1.4 wt%, the content of lanthanum oxide is 0.84 wt%, the content of cerium oxide is 0.18 wt%, and the content of other rare earth oxides is 0.38 wt%.

The catalysts used in the examples and comparative examples were homemade catalysts, denoted CAT, and the active component of the catalysts was ZRP molecular sieves, the specific properties of which are shown in table 1.

The catalyst CAT is prepared by the following steps:

Uniformly mixing ZRP molecular sieves, 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 performing spray drying, washing, filtering and drying to obtain the catalyst CAT. The catalyst CAT was aged at 790℃for 14 hours under 100% steam conditions before the test.

TABLE 1

Examples 1-2 and comparative examples 1-2 in this disclosure are presented to illustrate that the countercurrent reaction can promote further conversion of heavy oil catalytic cracking reaction intermediates to 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 cracker, the specific properties of which are shown in Table 2.

TABLE 2

Examples 1 to 2

The tests of examples 1-2 were performed on a medium-sized counter-current reaction test apparatus. Wherein the inner diameter of the countercurrent reactor is 64mm and the height is 500mm. Introducing light gasoline into the bottom of a countercurrent reactor, introducing a catalyst CAT into the top of the countercurrent reactor, carrying out countercurrent contact between the light gasoline and the catalyst CAT, then carrying out reaction, separating the reacted oil mixture through a cyclone separator to obtain a spent catalyst and reaction oil gas, introducing the spent catalyst into a stripper, then introducing the spent catalyst into a regenerator for regeneration to obtain a regenerated catalyst, returning the regenerated catalyst to a 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 to 2 were carried out on a medium-sized fluidized bed reaction test apparatus. The inside diameter of the fluidized bed reactor was 64mm and the height was 500mm. Light gasoline and catalyst CAT are simultaneously introduced into a fluidized bed reactor for reaction, the reacted oil mixture is separated by 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 Table 3

Examples 3-4 and comparative examples 3-4 are provided to illustrate the ability to increase heavy oil conversion and low olefin yields using the embodiment shown in fig. 1.

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

TABLE 4 Table 4

Example 3

The test of this example was performed on a medium-sized test apparatus as shown in FIG. 1. The apparatus comprises two riser reactors and a countercurrent 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 countercurrent 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 regenerated catalyst CAT from a regenerator 6 and performing 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 countercurrent reactor 3;

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

The first reaction oil gas from the first riser reactor 1 and the second spent catalyst from the second riser reactor 2 are contacted in a countercurrent reactor 3 and subjected to a third catalytic cracking reaction to obtain a third oil mixture, and the third oil mixture is separated by a cyclone separator to obtain third reaction oil gas and a third spent catalyst. And sending the third spent catalyst into a stripper 5, then sending the third spent catalyst into a regenerator 6 for 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 the third reaction oil gas into a fractionation system for separation.

Wherein the mass ratio of C4 hydrocarbon to 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 that of example 3, except that: in addition to the introduction of the C4 cut into the second riser reactor 2, a light gasoline cut (40-80 ℃ C. With 65% olefins by weight) was introduced into the second riser reactor 2.

Wherein, the mass ratio of C4 hydrocarbon, light gasoline fraction and 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 inner diameter of the riser reactor 1 was 16mm, the length was 3200mm, the inner diameter of the fluidized bed reactor 2 was 64mm, and the height was 500mm.

Wax oil is introduced into the bottom of the riser reactor 1, contacts with regenerated catalyst CAT from the regenerator 6 and undergoes catalytic cracking reaction to obtain a first oil mixture, the first oil mixture is introduced into the 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 the 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 performed on a medium-sized test apparatus. The apparatus comprises two riser reactors and a fluidized bed reactor. The first riser reactor 1 had an inner diameter of 16mm and a length of 3200mm, the second riser reactor 2 had an inner diameter of 16mm and a height of 3800mm, and the fluidized bed reactor 3 had an inner diameter of 64mm and a height of 500mm.

Introducing wax oil into the bottom of a first riser reactor 1, contacting with regenerated catalyst CAT from a regenerator 6 and performing a first catalytic cracking reaction to obtain a first oiling agent mixture, and introducing the first oiling agent mixture into a fluidized bed reactor 3; introducing C4 hydrocarbon into the bottom of the second riser reactor 2, contacting with regenerated catalyst CAT from a regenerator 6 and performing a second catalytic cracking reaction to obtain a second oil mixture, and introducing the second oil mixture into the fluidized bed reactor 3; and (3) 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 the fluidized bed reactor 3 to obtain a third oil mixture, separating the third oil mixture through a cyclone separator to obtain reaction oil gas and a spent catalyst, feeding the spent catalyst into a stripper 5, then feeding the spent catalyst into a regenerator 6 to regenerate the 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.

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

TABLE 5

As can be seen from tables 3 and 5, higher hydrocarbon conversion capacity and higher low carbon olefin yields can be achieved using the methods and apparatus provided by the present disclosure.

The preferred embodiments of the present disclosure have been described in detail above, but the present disclosure is not limited to the specific details of the above embodiments, and various simple modifications may be made to the technical solutions of the present disclosure within the scope of the technical concept of the present disclosure, and all the simple modifications belong to the protection scope of the present disclosure.

In addition, the specific features described in the above embodiments may be combined in any suitable manner without contradiction. The various possible combinations are not described further in this disclosure in order to avoid unnecessary repetition.

Moreover, any combination between the various embodiments of the present disclosure is possible as long as it does not depart from the spirit of the present disclosure, which should also be construed as the disclosure of the present disclosure.

Claims

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

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 a first riser reactor and a second riser reactor;

The first hydrocarbon raw material is selected from one or more than two mixtures of vacuum wax oil, normal pressure residual oil, vacuum residual oil, coking wax oil, deasphalted oil, furfural refined raffinate oil, coal liquefied oil, oil sand oil, shale oil, distillate oil obtained through 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 having an olefin content of greater than 50 wt%;

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

the first oil mixture is subjected to first separation through a separation baffle plate at the top of the first riser reactor to obtain first reaction oil gas and a first to-be-regenerated catalyst, and the first to-be-regenerated catalyst is introduced into a regenerator for regeneration reaction after steam stripping;

And introducing the second oil mixture into a quick-separating 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.

2. The catalytic conversion process of claim 1, wherein the weight ratio of the second hydrocarbon feedstock to the first hydrocarbon feedstock is from 0.05 to 0.20:1;

the mixture of C4-C8 hydrocarbons has an olefin content of greater than 60 wt.%.

3. The catalytic conversion process of claim 2, wherein the weight ratio of the second hydrocarbon feedstock to the first hydrocarbon feedstock is from 0.08 to 0.15:1.

4. The catalytic conversion process of claim 1, wherein the reaction temperature of the first riser reactor is 520-620 ℃; the ratio of the agent to the oil is 2-25; the reaction time is 1-15 seconds.

5. The catalytic conversion process of claim 4, wherein the reaction temperature of the first riser reactor is from 540 to 600 ℃; the agent-oil ratio is 3-20; the reaction time is 2-10 seconds.

6. The catalytic conversion process according to claim 1, wherein the second riser reactor has a reaction temperature of 560-660 ℃, a catalyst to oil ratio of 3-40, and a reaction time of 0.5-10 seconds.

7. The catalytic conversion process according to claim 6, wherein the reaction temperature of the second riser reactor is 580-640 ℃, the catalyst to oil ratio is 5-30, and the reaction time is 1-5 seconds.

8. The catalytic conversion process of claim 1, wherein the reaction temperature of the countercurrent reactor is 540-640 ℃; the density of the catalyst is 50-400kg/m ³; the residence time of the oil gas is 0.5-10 seconds.

9. The catalytic conversion process of claim 8, wherein the reaction temperature of the countercurrent reactor is 560-620 ℃; the density of the catalyst is 150-250kg/m ³; the residence time of the oil gas is 2-5 seconds.

10. The catalytic conversion process of claim 1, wherein the catalyst comprises an MFI structure molecular sieve, clay, and a binder;

The MFI structure molecular sieve is present in an amount of 20 to 60 wt% based on the total weight of the catalyst; the clay content is 10-70 wt%; the content of the binder is 10-40 wt%;

the MFI structure molecular sieve is selected from ZRP molecular sieves;

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

the binder is at least one selected from silica sol, alumina sol and pseudo-boehmite.

11. The catalytic conversion process of claim 10, wherein the MFI structure molecular sieve is present in an amount of 30-50 wt%, based on the total weight of the catalyst; the clay content is 15-45 wt%; the content of the binder is 20-35 wt%;

the binder is a double-aluminum binder of aluminum sol and pseudo-boehmite.

12. An apparatus suitable for use in the catalytic conversion process for producing light olefins according to any of the claims 1-11, characterized in that the apparatus comprises a combined reactor, a stripper and a regenerator, the combined reactor comprising a first riser reactor, a second riser reactor and a counter-current reactor;

13. The apparatus of claim 12, wherein the countercurrent reactor is an equal diameter reactor and/or a variable diameter reactor, the ratio of average diameter to height of the countercurrent reactor being 1:0.5-5.

14. The apparatus of claim 13, wherein the countercurrent reactor has an average diameter to height ratio of 1:1-3.