CN110857396A

CN110857396A - Method and system for processing coking gasoline and heavy raw oil

Info

Publication number: CN110857396A
Application number: CN201810975599.9A
Authority: CN
Inventors: 王迪; 龚剑洪; 魏晓丽
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: 2018-08-24
Filing date: 2018-08-24
Publication date: 2020-03-03
Anticipated expiration: 2038-08-24
Also published as: CN110857396B

Abstract

The invention relates to a method and a system for processing coking gasoline and heavy raw oil, wherein the method comprises the following steps: introducing a coking gasoline raw material into an adsorption and desorption reactor to contact with an n-alkane adsorbent and perform adsorption separation reaction; carrying out desorption treatment on the obtained adsorbent adsorbing the normal alkane by adopting desorption gas; introducing desorption oil and heavy raw oil into a riser reactor to contact with a catalytic cracking catalyst and carrying out a first catalytic cracking reaction; and (3) feeding the oil agent obtained by the reaction of the riser reactor into the fluidized bed reactor to contact with the raffinate oil and carrying out a second catalytic cracking reaction. The method and the system of the invention are used for processing the coker gasoline and the heavy raw oil, and have high propylene and butylene yields.

Description

Method and system for processing coking gasoline and heavy raw oil

Technical Field

The invention relates to a method and a system for processing coking gasoline and heavy raw oil.

Background

Propylene and butenes are important petrochemical industry base stocks. Currently, about 62% of the world's propylene comes from the co-production of ethylene by steam cracking. The steam cracking technology is perfected day by day, and is a process of consuming a large amount of energy, and is limited by the use of high-temperature materials, and the potential of further improvement is very small.

Coking is an important thermal processing process in petroleum refining, and the coking gasoline generated in the coking process is low-quality gasoline which is rich in olefin, has high contents of impurities such as sulfur, nitrogen compounds, diene and the like, has the characteristics of bad odor, poor stability, low octane number and the like, and the service performance of the coking gasoline needs to be improved through secondary processing.

U.S. Pat. No. 5,5685972 discloses a method for improving the octane number of gasoline by first carrying out hydrodesulfurization treatment on coker gasoline and then carrying out aromatization modification on coker gasoline by using a metal modified ZSM-5 molecular sieve catalyst.

Chinese patent CN1715372A discloses a method for reforming coker gasoline, in which hydrogenated coker distillate is cut to obtain reformed raw oil with a suitable distillation range, and then the coker gasoline is reformed under the reaction condition of catalytic reforming to produce high-octane gasoline blending component.

Chinese patent CN 1160746 discloses a catalytic conversion method for increasing gasoline octane number. Injecting low-quality gasoline such as straight-run gasoline and coking gasoline into the lower part of a riser reactor, and preferentially contacting and reacting with a regenerated catalyst; the reaction temperature is 600 ℃ and 700 ℃, and the weight hourly space velocity is 1-180h^-1The ratio of agent to oil is 6-180. The method can improve the octane number of the low-quality gasoline and reduce the olefin content of the gasoline to a certain extent.

Chinese patent CN1069054 adopts a double-riser reactor to perform reaction, low-quality gasoline including catalytic cracking crude gasoline and coker gasoline is injected into the riser reactor, and catalytic modification of the low-quality gasoline is realized by utilizing reaction conditions of high temperature and large catalyst-to-oil ratio, so that the yield of liquefied gas and the octane number of the gasoline are improved.

Chinese patent CN201110420264 provides a method for modifying coker gasoline, the coker gasoline is subjected to selective cracking reaction under the action of selective cracking catalyst, the reaction product is subjected to gas-liquid separation, the separated liquid product is subjected to aromatization reaction, the liquid product of aromatization reaction is subjected to hydrodesulfurization, and all the reaction products are fractionated to obtain hydrofined aromatization oil of low-carbon olefin products.

Among the above methods, the hydrofining method can reduce the content of impurities and olefins, improve the stability, and can still have a great influence on the octane number of a refinery gasoline pool due to the low octane number of the hydrofining method, when used as a gasoline blending component. The coking gasoline can directly enter a catalytic cracking device for cracking, and has the problems that the inactivation of a cracking catalyst is accelerated by high nitrogen content, the product distribution is influenced, the alkane cracking reaction activity in the gasoline is low, and a large amount of unreacted alkane enters a gasoline product to influence the octane number of the gasoline. When the coking gasoline is used as a reforming raw material, the octane number and the reforming hydrogen production rate of the reforming gasoline are influenced by overlarge blending proportion due to low aromatic hydrocarbon potential, and the blending proportion is generally below 30 percent. In recent years, under the conditions that the environmental protection requirement is increasingly strict, the quality of gasoline products is continuously upgraded and the demand of chemical raw materials is vigorous, how to reduce the content of impurities and olefin of the coker gasoline as much as possible and greatly improve the octane number of the coker gasoline, or convert the coker gasoline into the chemical raw materials such as propylene, butylene and the like to the maximum degree is an urgent task.

Disclosure of Invention

The invention aims to provide a method and a system for processing coker gasoline and heavy raw oil, which are used for processing the coker gasoline and the heavy raw oil and have high propylene and butylene yields.

In order to achieve the above object, the present invention provides a process for processing coker gasoline and heavy feed oil, comprising:

introducing a coking gasoline raw material into an adsorption and desorption reactor to contact with an n-alkane adsorbent and carry out adsorption separation reaction to obtain the adsorbent adsorbed with n-alkane and absorption residual oil;

carrying out desorption treatment on the obtained adsorbent with adsorbed normal alkane by adopting desorption gas to obtain desorbed adsorbent and desorbed oil;

introducing desorption oil and heavy raw oil into a riser reactor to contact with a catalytic cracking catalyst and carrying out a first catalytic cracking reaction;

feeding the oil agent obtained by the reaction of the riser reactor into a fluidized bed reactor to contact with the raffinate oil and perform a second catalytic cracking reaction to obtain a reaction product and a spent catalyst;

separating the obtained reaction product from the spent catalyst, feeding the obtained spent catalyst into a regenerator for coke burning regeneration, and returning at least part of the obtained regenerated catalyst to the riser reactor for use as the catalytic cracking catalyst.

Optionally, the coking gasoline raw material contains 6-30 wt% of normal paraffin, 20-40 wt% of olefin and 100-1000 microgram/g of nitrogen;

the nitrogen content in the raffinate oil is 60-600 micrograms/gram;

the n-alkane content in the desorption oil is 90-98 wt%, and the nitrogen content is 0-100 micrograms/gram;

the heavy raw oil is one or more selected from vacuum wax oil, atmospheric residue oil, vacuum residue oil, coking wax oil, coking residue oil and Fischer-Tropsch wax oil.

Optionally, the adsorption and desorption reactor is a fixed bed reactor, a moving bed reactor, a simulated moving bed reactor or an expanded bed reactor.

Optionally, the n-alkane adsorbent is one or more selected from activated carbon, activated carbon fibers, carbonized resin silica gel, natural zeolite, synthetic zeolite, molecular sieve and activated alumina.

Optionally, the conditions of the adsorption separation reaction include: the temperature is 250-380 ℃, and the weight hourly space velocity of the coking gasoline raw material is 0.1-20 hours^-1；

The conditions of the desorption treatment include: the temperature is 300-450 ℃, the desorption gas is nitrogen or hydrogen, and the weight hourly space velocity of the desorption gas is 100-200 hours^-1。

Optionally, according to the flow direction of the reaction materials, the desorption oil feed port is positioned at the upstream of the heavy raw oil feed port, and the proportion of the distance between the desorption oil feed port and the heavy raw oil feed port in the total height of the riser reactor is 2-20%.

Optionally, the conditions of the first catalytic cracking reaction include: the reaction temperature is 560 ℃ and 750 ℃, the reaction time is 1-10 seconds, the weight ratio of the catalyst to the oil is 1-50, and the weight ratio of the water to the oil is (0.01-1): 1, the feeding weight ratio of the raffinate oil, the desorption oil and the heavy raw oil is (0.05-0.5): (0.05-0.3): 1;

the conditions of the second catalytic cracking reaction include: the reaction temperature is 550-720 ℃, and the weight hourly space velocity is 0.5-20 h^-1。

Optionally, the conditions of the first catalytic cracking reaction include: the reaction temperature is 580-730 ℃, the reaction time is 2-5 seconds, the weight ratio of the catalyst to the oil is 5-30, and the weight ratio of the water to the oil is (0.05-0.5): 1;

the conditions of the second catalytic cracking reaction include: the reaction temperature is 570-700 ℃, and the weight hourly space velocity is 2-10 h^-1。

Optionally, the method further includes: preheating desorption oil and residual absorption oil, and then respectively introducing the preheated desorption oil and the preheated residual absorption oil into a riser reactor and a fluidized bed reactor, wherein the temperatures of the preheated desorption oil and the preheated residual absorption oil are respectively 350-450 ℃.

Optionally, the catalytic cracking catalyst comprises, on a dry basis and based on the total weight of the catalyst, from 1 to 60 wt% zeolite, from 5 to 99 wt% inorganic oxide, and from 0 to 70 wt% clay;

the zeolite comprises 50 to 100 wt% of a medium pore zeolite and 0 to 50 wt% of a large pore zeolite, on a dry basis and based on the total weight of the zeolite.

Optionally, the zeolite comprises 70 to 100 wt% of a medium pore zeolite and 0 to 30 wt% of a large pore zeolite, on a dry basis and based on the total weight of the zeolite.

Optionally, the medium pore zeolite is a ZSM series zeolite and/or a ZRP zeolite, and the large pore zeolite is one or more selected from rare earth Y, rare earth hydrogen Y, ultrastable Y and high silica Y;

the inorganic oxide is silicon dioxide and/or aluminum oxide;

the clay is kaolin and/or halloysite.

Optionally, the method further includes: feeding the regenerated catalyst from the regenerator into a degassing tank for degassing, and then feeding the regenerated catalyst into a riser reactor for use as the catalytic cracking catalyst, wherein oxygen-containing gas obtained by degassing in the degassing tank is returned to the regenerator.

The invention also provides a processing system of the coker gasoline and the heavy raw oil, which comprises an adsorption and desorption reactor, a riser reactor, a fluidized bed reactor, a stripping section, a settler and a regenerator, wherein the settler, the fluidized bed reactor and the stripping section are arranged from top to bottom and are in fluid communication, and the riser reactor penetrates through the stripping section from bottom to top to enter the fluidized bed reactor;

the adsorption and desorption reactor is provided with a coker gasoline raw material inlet, a desorption gas inlet, an absorbed residual oil outlet and a desorbed oil outlet, the riser reactor is provided with a desorbed oil feed inlet, a heavy raw oil feed inlet, a catalyst inlet and an oil agent outlet, the fluidized bed reactor is provided with an absorbed residual oil feed inlet, the stripping section is provided with a catalyst outlet, the settler is provided with an oil gas outlet, and the regenerator is provided with a catalyst inlet and a catalyst outlet;

the absorption desorption reactor inhale the residual oil export with fluidized bed reactor inhale residual oil feed inlet fluid intercommunication, the desorption oil export of absorption desorption reactor with the desorption oil feed inlet fluid intercommunication of riser reactor, the finish outlet of riser reactor is located in fluidized bed reactor, the catalyst entry of riser reactor with the catalyst outlet fluid intercommunication of regenerator, the catalyst entry of regenerator with the catalyst outlet fluid intercommunication of strip section.

Optionally, the system further comprises a degassing tank through which the catalyst inlet of the riser reactor is in fluid communication with the catalyst outlet of the regenerator; and/or

The system also comprises a preheating device for preheating the desorption oil and/or the residual absorption oil.

The normal alkane component with lower reaction activity (namely desorption oil) and the non-normal alkane component with higher reaction activity (namely absorption residual oil) in the coker gasoline are separated, and the normal alkane component independently enters the reactor to be in contact reaction with the regenerated catalyst, so that the competitive reaction of the non-normal alkane component on the active center of the catalyst is reduced, the catalytic cracking reaction performance of the normal alkane component is improved, and the yield of propylene and butylene in the coker gasoline catalytic cracking process is improved.

The invention can contact the heavy raw oil and normal alkane with the regenerated catalyst first to do cracking reaction, to make the regenerated catalyst have little carbon deposit, adjust the catalyst property reasonably, reduce the strong acid center quantity, provide proper catalyst acidity for non normal alkane with high reaction activity, especially olefin, and improve the yield and selectivity of propylene and butylene.

The invention can partially remove nitrogen compounds in the coking gasoline in the adsorption separation process, and reduces the toxic action of the adsorption of the nitrogen compounds on the catalytic cracking catalyst on the active center of the catalyst.

The invention not only solves the problem of reasonable and efficient utilization of the coking gasoline, but also solves the problem of shortage of petrochemical raw materials, and improves the economic benefit and the social benefit of the petrochemical industry.

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

Drawings

The accompanying drawings, which are included to provide a further understanding of the invention and are incorporated in and constitute a part of this specification, illustrate embodiments of the invention and together with the description serve to explain the principles of the invention and not to limit the invention. In the drawings:

FIG. 1 includes a schematic flow diagram of one embodiment of the method of the present invention and also includes a schematic structural diagram of one embodiment of the system of the present invention.

Description of the reference numerals

1 riser reactor 2 regenerator 3 settler

4 stripping section 5 degassing tank 6 cyclone separator

7 gas collection chamber 8 spent inclined tube 9 spent slide valve

10 line 11 line 12 regenerative chute

13 regenerative slide valve 14 line 15 line

16 line 17 line 18 line

19 line 20 large oil and gas line 21 line

22 air distributor 23 line 24 cyclone

25 flue gas duct 26 fluidized bed reactor 27 line

28 line 29 line 30 line

31 adsorption-desorption reactor 32 heating furnace 33 pipeline

34 pipeline

Detailed Description

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

The invention provides a method for processing coking gasoline and heavy raw oil, which comprises the following steps:

feeding the oil agent obtained by the reaction of the riser reactor into the fluidized bed reactor to contact with the raffinate oil and perform a second catalytic cracking reaction to obtain a reaction product and a spent catalyst;

According to the present invention, Coker gasoline (english name: Coker naptha), also known as Coker Naphtha, is a gasoline fraction produced by a delayed coking process. The normal paraffin content in the coking gasoline raw material can be 6-30 wt%, the olefin content can be 20-40 wt%, and the nitrogen content can be 100-1000 microgram/g.

According to the invention, the adsorption separation reaction and the desorption treatment can remove part of nitrogen in the coker gasoline and can also enable the desorption oil to be rich in normal paraffin, for example, the nitrogen content in the raffinate oil can be 60-600 micrograms/gram; the n-alkane content in the desorption oil can be 90-98 wt%, the nitrogen content can be 0-100 micrograms/gram, and the adsorption and desorption reactor can be a fixed bed reactor, a moving bed reactor, a simulated moving bed reactor or an expanded bed reactor, and is preferably a fixed bed reactor. The n-alkane adsorbent may be one or more selected from the group consisting of activated carbon, activated carbon fiber, carbonized resin silica gel, natural zeolite, synthetic zeolite, and activated alumina, and when carried out in a fixed bed reactor, zeolite is preferably used as the n-alkane adsorbent. The conditions of the adsorptive separation reaction may include: the temperature is 250-380 ℃, and the weight hourly space velocity (the weight ratio of the coking gasoline raw material feed to the adsorbent in the reactor in unit time) of the coking gasoline raw material is 0.1-20 hours^-1(ii) a The conditions of the desorption treatment may include: the temperature is 300-450 ℃, the desorption gas is nitrogen or hydrogen, the desorption gas can be used as carrier gas for conveying desorption oil and is sent into the catalytic cracking reactor for reaction, the weight hourly space velocity (the weight ratio of the desorption gas feed to the adsorbent in the reactor in unit time) of the desorption gas is 100-200 h^-1。

According to the present invention, the heavy feedstock oil is well known to those skilled in the art, for example, the heavy feedstock oil may be one or more selected from vacuum wax oil, atmospheric residue, vacuum residue, coker wax oil, coker residue, and fischer-tropsch wax oil.

According to the present invention, a riser reactor and a fluidized bed reactor are well known to those skilled in the art, the riser reactor may be an equal-diameter riser reactor or a variable-diameter riser reactor, preferably an equal-diameter riser reactor, the variable-diameter riser reactor is, for example, an equal linear velocity riser reactor, the riser reactor may be provided with a plurality of feed inlets, the feed ratio of each feed inlet may be the same or different, the number of the feed inlets may be two or more, preferably two, the riser reactor may include a pre-lifting section and at least one reaction zone from bottom to top, in order to enable the raw oil to fully react, and the number of the reaction zones may be 2 to 8, preferably 2 to 3, according to different target product quality requirements. Further preferably, according to the flow direction of the reaction materials, the desorption oil feed port is positioned at the upstream of the heavy raw oil feed port, and the distance between the desorption oil feed port and the heavy raw oil feed port accounts for 2-20% of the total height of the riser reactor.

Catalytic cracking reactions according to the present invention are well known to those skilled in the art, and in particular for the purposes of the present invention, the conditions of the first catalytic cracking reaction may include: the reaction temperature (outlet temperature) is 560-: 1, the feeding weight ratio of the raffinate oil, the desorption oil and the heavy raw oil is (0.05-0.5): (0.05-0.3): 1; the conditions of the first catalytic cracking reaction preferably include: reaction temperatureThe degree is 580-730 ℃, more preferably 600-700 ℃, the reaction time is 2-5 seconds, the oil-water weight ratio is 5-30, and the water-oil weight ratio is (0.05-0.5): 1; the conditions of the second catalytic cracking reaction may include: the reaction temperature (bed layer temperature) is 550-720 ℃, and the weight hourly space velocity is 0.5-20 h^-1The weight ratio of water to oil is (0.01-1): 1; the conditions of the second catalytic cracking reaction preferably include: the reaction temperature is 570-700 ℃, and the weight hourly space velocity is 2-10 h^-1。

According to the invention, the method may further comprise: the desorption oil and the residual absorption oil are preheated and then respectively introduced into the riser reactor and the fluidized bed reactor, the temperature of the preheated desorption oil and the preheated residual absorption oil is respectively and independently 350-450 ℃, preferably 380-420 ℃, and the desorption oil and the residual absorption oil are both in a gas state, so that the contact efficiency of the oil agent is improved, and a heat source can be provided for the reaction.

Catalytic cracking catalysts are well known to those skilled in the art in accordance with the present invention, and in particular for the purposes of the present invention, the catalytic cracking catalyst may comprise from 1 to 60 wt% zeolite, from 5 to 99 wt% inorganic oxide, and from 0 to 70 wt% clay, on a dry basis and based on the total weight of the catalyst; the zeolite may comprise as an active component a medium pore zeolite and/or a large pore zeolite, which may comprise, on a dry basis and based on the total weight of the zeolite, from 50 to 100% by weight of medium pore zeolite and from 0 to 50% by weight of large pore zeolite, preferably from 70 to 100% by weight of medium pore zeolite and from 0 to 30% by weight of large pore zeolite. The medium and large pore zeolites are defined as conventional in the art, i.e., the medium pore zeolite has an average pore size of 0.5 to 0.6nm and the large pore zeolite has an average pore size of 0.7 to 1.0 nm. The medium pore zeolite may be a zeolite having an MFI structure, such as a ZSM-series zeolite and/or a ZRP zeolite, which may also be modified with a nonmetallic element such as phosphorus and/or a transition metal element such as iron, cobalt, nickel, as described in more detail in U.S. Pat. No. 5,232,675, and the ZSM-series zeolite is selected from one or more mixtures of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48 and other zeolites of similar structure, as described in more detail in US3,702,886. The large pore zeolite may beIs one or more selected from Rare Earth Y (REY), Rare Earth Hydrogen Y (REHY), ultrastable Y and high silicon Y; the inorganic oxide may be silicon dioxide (SiO) as a binder₂) And/or aluminum oxide (Al)₂O₃) (ii) a The clay as a matrix (carrier) may be kaolin and/or halloysite.

According to the present invention, the spent catalyst and the reaction product are generally separated to obtain the spent catalyst and the reaction product, then the obtained reaction product is subjected to a subsequent separation system to separate fractions such as dry gas, liquefied gas, pyrolysis gasoline and pyrolysis diesel oil, then the liquefied gas is further separated by a gas separation device to obtain propylene, butylene and the like, and the method for separating propylene and butylene from the reaction product is similar to the conventional technical method in the art, and the present invention is not limited thereto, and is not described in detail herein.

According to the present invention, the regeneration of the spent catalyst is well known to those skilled in the art, all or at least part of the catalytic cracking catalyst can be from the regenerated catalyst, during the regeneration process, an oxygen-containing gas is generally introduced from the bottom of the regenerator, the oxygen-containing gas can be, for example, air, and then the spent catalyst is contacted with oxygen for coke burning regeneration, the gas-solid separation is performed on the upper part of the regenerator on the flue gas generated after the catalyst is burned and regenerated, and the flue gas enters the subsequent energy recovery system. The method of the present invention preferably further comprises: the regenerated catalyst from the regenerator is sent to a degassing tank for degassing and then sent to a riser reactor to be used as the catalytic cracking catalyst, oxygen-containing gas obtained by degassing in the degassing tank is returned to the regenerator, and impurities such as oxygen are removed from the regenerated catalyst in the degassing tank by stripping gas such as water vapor. The conditions for regeneration may include: the regeneration temperature is 550-750 ℃, preferably 600-730 ℃, and more preferably 650-700 ℃; the gas superficial linear velocity is 0.5 to 3 m/s, preferably 0.8 to 2.5 m/s, more preferably 1 to 2 m/s, and the average residence time of the spent catalyst is 0.6 to 3 minutes, preferably 0.8 to 2.5 minutes, more preferably 1 to 2 minutes.

In the invention, the adsorption and desorption reactor can be a fixed bed reactor, a moving bed reactor, a simulated moving bed reactor or an expanded bed reactor, and is preferably a fixed bed reactor; riser reactors and fluidized bed reactors are well known to those skilled in the art, a riser reactor may be an equal diameter riser reactor or a variable diameter riser reactor, preferably an equal diameter riser reactor, a variable diameter riser reactor is, for example, an equal linear velocity riser reactor, a plurality of feed inlets may be provided in the riser reactor, the feed ratio of each feed inlet may be the same or different, the number of feed inlets may be two or more, preferably two, the riser reactor may include a pre-lifting section and at least one reaction zone from bottom to top, in order to enable the raw oil to react sufficiently, and according to different target product quality requirements, the number of the reaction zones may be 2 to 8, preferably 2 to 3. Further preferably, according to the flow direction of the reaction materials, the desorption oil feed port is positioned at the upstream of the heavy raw oil feed port, and the distance between the desorption oil feed port and the heavy raw oil feed port accounts for 2-20% of the total height of the riser reactor.

According to the invention, the system may further comprise a degassing tank, the catalyst inlet of the riser reactor being in fluid communication with the catalyst outlet of the regenerator via the degassing tank, the degassing tank may be provided with a catalyst inlet, a catalyst outlet, a stripping gas inlet and a stripping gas outlet, the catalyst inlet being in communication with the regenerator catalyst outlet, the catalyst outlet being in communication with the reactor catalyst inlet, the stripping gas inlet being for introduction of stripping gas such as water vapour, the stripping gas outlet being in communication with the regenerator for feeding oxygen-containing stripped gas into the regenerator; and/or the system can also comprise a preheating device for preheating the desorption oil and/or the residual absorption oil, wherein the preheating device can be a heating furnace and the like.

The invention will be further illustrated by means of specific embodiments in the following description with reference to the drawings, without being restricted thereto.

As shown in fig. 1, the coker gasoline enters the top of an adsorption and desorption reactor 31 through a pipeline 27, and after adsorption and separation by a molecular sieve, the absorption residual oil enters a heating furnace 32 through a pipeline 30 for preheating; nitrogen is injected into the bottom of the adsorption and desorption reactor through a pipeline 28 to cause the molecular sieve to desorb the adsorbed normal paraffin components, and the generated desorption oil enters a heating furnace 32 through a pipeline 29 to be preheated.

The pre-lifting medium enters from the bottom of the riser reactor 1 through a pipeline 14, the regenerated catalyst from a regenerated inclined pipe 12 enters the riser reactor 1 after being regulated by a regeneration slide valve 13, and moves upwards in an accelerated manner along the riser under the lifting action of the pre-lifting medium, the preheated desorption oil is injected into the lower part of the riser reactor 1 through a pipeline 16 together with the atomized steam from a pipeline 15, and is mixed with the existing material flow of the riser reactor, and the raw oil is subjected to catalytic cracking reaction on the hot catalyst and moves upwards in an accelerated manner. The preheated heavy raw oil is injected into the middle lower part of the riser reactor 1 through a pipeline 34 and atomized steam from a pipeline 33, and is mixed with the existing material flow of the riser reactor, the raw oil is subjected to catalytic cracking reaction on a hot catalyst and moves upwards in an accelerated manner; the preheated raffinate is injected into the fluidized bed reactor 26 through the pipeline 18 together with the atomized steam from the pipeline 17, and is mixed with the material flow from the riser reactor, and the raffinate undergoes catalytic cracking reaction on the catalyst and moves upward in an accelerated manner. The generated reaction product and the inactivated spent catalyst enter a cyclone separator 6 in a settler 3 to realize the separation of the spent catalyst and the reaction product, the reaction product enters an air collection chamber 7, and the fine powder of the catalyst returns to the settler. Spent catalyst in the settler flows to the stripping section 4 where it is contacted with steam from line 19. The reaction product stripped from the spent catalyst enters the gas collection chamber 7 after passing through the cyclone separator. The stripped spent catalyst enters the regenerator 2 after being regulated by a spent slide valve 9 through a spent inclined pipe 8, air from a pipeline 21 enters the regenerator 2 after being distributed by an air distributor 22, coke on the spent catalyst in a dense bed layer at the bottom of the regenerator 2 is burned off to regenerate the inactivated spent catalyst, and flue gas enters a subsequent energy recovery system through an upper flue gas pipeline 25 of a cyclone separator 24. Wherein the pre-lifting medium may be dry gas, water vapor or a mixture thereof.

The regenerated catalyst enters a degassing tank 5 through a pipeline 10 communicated with a catalyst outlet of a regenerator 2, and is contacted with a stripping medium from a pipeline 23 at the bottom of the degassing tank 5 to remove flue gas (namely oxygen-containing gas) carried by the regenerated catalyst, the degassed regenerated catalyst circulates to the bottom of a riser reactor 1 through a regeneration inclined tube 12, the catalyst circulation amount can be controlled through a regeneration slide valve 13, the flue gas returns to the regenerator 2 through a pipeline 11, and reaction products in a gas collection chamber 7 enter a subsequent separation system through a large oil-gas pipeline 20.

The following examples further illustrate the invention but are not intended to limit the invention thereto.

The feedstock oils used in the examples and comparative examples were coker gasoline and Daqing atmospheric residue, the properties of which are shown in tables 1 and 2, respectively.

The catalytic cracking catalyst used in the examples and comparative examples was sold under the trade name RAG-6.

Comparative example 1

Performing a test on a medium-sized device of a riser reactor, wherein raw oil is coker gasoline and heavy raw oil, preheated coker gasoline enters the lower part of the riser reactor to perform a catalytic cracking reaction, the preheated heavy raw oil enters the middle lower part of the riser reactor to perform the catalytic cracking reaction together with materials from the lower part, the feed weight ratio of the coker gasoline to Daqing atmospheric residue per unit time is 0.15:1, the ratio of the feed inlet of the coker gasoline to the feed inlet of the heavy raw oil to the total height of the riser reactor is 5%, a reaction oil gas and catalyst mixture continuously ascends to enter a fluidized bed reactor to continue the reaction, a reaction product and a spent catalyst enter a closed cyclone separator from the outlet of the reactor, the reaction product and the spent catalyst are rapidly separated, and the reaction product is cut according to the distillation range in a separation system to obtain, butene and pyrolysis gasoline.

The spent catalyst enters a stripping section under the action of gravity, hydrocarbon products adsorbed on the spent catalyst are stripped by steam, and the stripped spent catalyst enters a regenerator and is in contact with air for regeneration; the regenerated catalyst enters a degassing tank to remove oxygen-containing gas adsorbed and carried by the regenerated catalyst; the degassed regenerated catalyst returns to the riser reactor for recycling; the operating conditions and the product distribution are listed in Table 3.

As can be seen from the results of table 3, the yields of propylene and butene were 11.78 wt% and 14.66 wt%, respectively.

Example 1

The test is carried out according to the flow of fig. 1, the raw oil is a coker gasoline raw material, the adsorption separation reaction is carried out on a fixed bed adsorption and desorption reactor, the normal alkane adsorbent is a 5A molecular sieve, and the adsorption separation reaction conditions are as follows: the temperature is 300 ℃, and the weight hourly space velocity of the coking gasoline raw material is 1.0 hour^-1(ii) a Desorption treatment conditions: the desorption gas is nitrogen, and the weight hourly space velocity of the desorption gas is 150 hours^-1And the temperature is 360 ℃. The nitrogen content in the raffinate oil is 350 microgram/g, the normal alkane content in the desorbed oil is 95 weight percent, and the nitrogen content is 10 microgram/g. The desorbed oil and the residual oil generated by the adsorption separation of the coker gasoline enter a heating furnace and are heated to 420 ℃.

The catalytic cracking reaction is carried out on a medium-sized device of a riser reactor, preheated desorption oil enters the lower part of the riser reactor, preheated heavy raw oil enters the middle lower part of the riser reactor to carry out first catalytic cracking reaction, preheated raffinate oil enters a fluidized bed reactor to be mixed with material flow from the riser reactor and carries out second catalytic cracking reaction on a small amount of carbon-deposited catalyst, the proportion of the distance between a desorption oil feed inlet and a heavy raw oil feed inlet to the total height of the riser reactor is 5%, and the feed weight ratio of the desorption oil, the raffinate oil and the heavy raw oil in unit time is 0.05: 0.1: 1, enabling a reaction product and a spent catalyst to enter a closed cyclone separator from an outlet of the reactor, quickly separating the reaction product from the spent catalyst, and cutting the reaction product in a separation system according to a distillation range to obtain fractions such as propylene, butylene, pyrolysis gasoline and the like; the spent catalyst enters a stripping section under the action of gravity, hydrocarbon products adsorbed on the spent catalyst are stripped by steam, and the stripped catalyst enters a regenerator and is in contact with air for regeneration; the regenerated catalyst enters a degassing tank to remove oxygen-containing gas adsorbed and carried by the regenerated catalyst; the degassed regenerated catalyst returns to the riser reactor for recycling; the operating conditions and the product distribution are listed in Table 3.

As can be seen from the results of table 3, the yields of propylene and butene were 18.65 wt% and 17.32 wt%, respectively.

Comparative example 2

The same as example 1 except that: the feed ports for the absorbed residual oil and desorbed oil were changed in the weight ratio per unit time, the absorbed residual oil was fed from the riser reactor, and the desorbed oil was fed from the fluidized bed reactor, and the other steps and conditions were the same as in example 1, and the operating conditions and product distribution are shown in Table 3.

As can be seen from the results of table 3, the process and system of the present invention have the advantage of high propylene and butene yields.

The preferred embodiments of the present invention have been described in detail, however, the present invention 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 invention within the technical idea of the present invention, and these simple modifications are within the protective scope of the present invention.

It should be noted that the various technical features described in the above embodiments can be combined in any suitable manner without contradiction, and the invention is not described in any way for the possible combinations in order to avoid unnecessary repetition.

In addition, any combination of the various embodiments of the present invention is also possible, and the same should be considered as the content of the present invention as long as it does not depart from the gist of the present invention.

TABLE 1

Properties of crude oil
		Density (20 deg.C), g/cm³	0.7373
Vapour pressure, kilopascal	50.0
		Nitrogen content, microgram/gram	400
Group composition, weight%
		N-alkanes	28.2
Isoalkanes	18.1
		Olefins	35.8
Cycloalkanes	7.9
		Aromatic hydrocarbons	10.0
Distillation range, deg.C
		Initial boiling point	44
10% by volume	65
		30% by volume	90
50% by volume	122
		70% by volume	153
90% by volume	184
		95% by volume	195
End point of distillation	208

TABLE 2

Properties of crude oil	Daqing atmospheric residue
		Density (20 deg.C), g/cm³	0.8974
Viscosity (100 ℃), mm³Second/second	30.02
		Carbon residue, by weight%	4.5
Group composition, weight%
		Saturated hydrocarbons	56.4
Aromatic hydrocarbons	25.9
		Glue	17.6
Asphaltenes	0.1
		Carbon content, wt.%	86.26
Hydrogen content, wt.%	12.91
		Distillation range, deg.C
Initial boiling point	324
		10% by volume	408
30% by volume	486
		50% by volume	555

TABLE 3

	Comparative example 1	Example 1	Comparative example 2
				Adsorption separation reaction and desorption treatment
Adsorption temperature of low DEG C	/	300	300
				Weight hourly space velocity of raw materials, hours^-1	/	1.0	1.0
Desorption temperature,. degree.C	/	360	360
				Weight hourly space velocity, hour of desorbed gas^-1	/	150	150
Catalytic cracking unit
				Riser reactor
Outlet temperature of	635	635	635
				Reaction time in seconds	1.8	1.8	1.8
Water to oil weight ratio	0.20	0.20	0.20
				Weight ratio of solvent to oil	25	25	25
Fluidized bed reactor
				Bed reaction temperature of deg.C	620	620	620
Weight hourly space velocity, hours^-1	5	5	5
				Product distribution, weight%
Dry gas	4.11	4.47	4.24
				LPG	31.02	39.33	37.37
Gasoline (gasoline)	40.13	32.10	33.59
				Diesel oil	12.56	12.32	12.64
Heavy oil	4.98	4.67	4.80
				Coke	7.20	7.11	7.36
Total up to	100.00	100.00	100
				Propylene (PA)	11.78	18.65	16.79
Butene (butylene)	14.66	17.32	15.59
				Conversion, wt.%	82.46	83.01	82.56

Claims

1. A process for processing coker gasoline and heavy feed oil, the process comprising:

2. The process as claimed in claim 1, wherein the coker gasoline feedstock has an n-paraffin content of 6-30 wt%, an olefin content of 20-40 wt%, and a nitrogen content of 100-;

the nitrogen content in the raffinate oil is 60-600 micrograms/gram;

3. The process of claim 1, wherein the adsorption-desorption reactor is a fixed bed reactor, a moving bed reactor, a simulated moving bed reactor, or an expanded bed reactor.

4. The method according to claim 1, wherein the n-alkane adsorbent is one or more selected from the group consisting of activated carbon, activated carbon fiber, carbonized resin silica gel, natural zeolite, synthetic zeolite, molecular sieve and activated alumina.

5. The method of claim 1, wherein the conditions of the adsorptive separation reaction comprise: the temperature is 250-380 ℃, and the weight hourly space velocity of the coking gasoline raw material is 0.1-20 hours^-1；

6. The process according to claim 1, wherein the desorption oil feed port is located upstream of the heavy raw oil feed port in terms of the flow direction of the reaction material, and the gap between the desorption oil feed port and the heavy raw oil feed port accounts for 2-20% of the total height of the riser reactor.

7. The method of claim 1, wherein the conditions of the first catalytic cracking reaction comprise: the reaction temperature is 560 ℃ and 750 ℃, the reaction time is 1-10 seconds, the weight ratio of the catalyst to the oil is 1-50, and the weight ratio of the water to the oil is (0.01-1): 1, the feeding weight ratio of the raffinate oil, the desorption oil and the heavy raw oil is (0.05-0.5): (0.05-0.3): 1;

8. The method of claim 1, wherein the conditions of the first catalytic cracking reaction comprise: the reaction temperature is 580-730 ℃, the reaction time is 2-5 seconds, the weight ratio of the catalyst to the oil is 5-30, and the weight ratio of the water to the oil is (0.05-0.5): 1;

9. The method of claim 1, further comprising: preheating desorption oil and residual absorption oil, and then respectively introducing the preheated desorption oil and the preheated residual absorption oil into a riser reactor and a fluidized bed reactor, wherein the temperatures of the preheated desorption oil and the preheated residual absorption oil are respectively 350-450 ℃.

10. The process of claim 1 wherein the catalytic cracking catalyst comprises, on a dry basis and based on the total weight of the catalyst, from 1 to 60 wt% zeolite, from 5 to 99 wt% inorganic oxide, and from 0 to 70 wt% clay;

11. The process of claim 10 wherein the zeolite comprises from 70 to 100 wt% of a medium pore zeolite and from 0 to 30 wt% of a large pore zeolite, on a dry basis and based on the total weight of the zeolite.

12. The process of claim 10 wherein the medium pore zeolite is a ZSM-series zeolite and/or a ZRP zeolite and the large pore zeolite is one or more selected from rare earth Y, rare earth hydrogen Y, ultrastable Y and high silica Y;

the inorganic oxide is silicon dioxide and/or aluminum oxide;

the clay is kaolin and/or halloysite.

13. The method of claim 1, further comprising: feeding the regenerated catalyst from the regenerator into a degassing tank for degassing, and then feeding the regenerated catalyst into a riser reactor for use as the catalytic cracking catalyst, wherein oxygen-containing gas obtained by degassing in the degassing tank is returned to the regenerator.

14. A processing system for coker gasoline and heavy raw oil comprises an adsorption-desorption reactor, a riser reactor, a fluidized bed reactor, a stripping section, a settler and a regenerator, wherein the settler, the fluidized bed reactor and the stripping section are arranged from top to bottom and are in fluid communication with each other, and the riser reactor penetrates through the stripping section from bottom to top and enters the fluidized bed reactor;

15. The system of claim 14, wherein the adsorption-desorption reactor is a fixed bed reactor, a moving bed reactor, a simulated moving bed reactor, or an expanded bed reactor.

16. The system according to claim 14, wherein the desorption oil feed port is located upstream of the heavy feed oil feed port according to the flow direction of the reaction material, and the distance between the desorption oil feed port and the heavy feed oil feed port accounts for 2-20% of the total height of the riser reactor.

17. The system of claim 14, further comprising a degassing tank through which a catalyst inlet of the riser reactor is in fluid communication with a catalyst outlet of the regenerator; and/or