CN110857398B - Method and system for processing coking gasoline and heavy raw oil - Google Patents

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 heavy raw oil into a first reaction zone of a main riser reactor to contact with a first catalytic cracking catalyst and carrying out a first catalytic cracking reaction; sending the oil agent obtained by the reaction in the first reaction zone into the second reaction zone to contact with the raffinate oil and carrying out a second catalytic cracking reaction; and introducing the desorbed oil into the secondary riser reactor to contact with the second catalytic cracking catalyst and perform a third catalytic cracking reaction to obtain a second reaction product and a second spent catalyst. 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. The catalytic cracking unit plays a significant role in the production of low-carbon olefins as a second major source of propylene production, and particularly plays a role in bearing an important historical mission of producing more propylene in the future.

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 butylene and propylene 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 heavy raw oil into a first reaction zone of a main riser reactor to contact with a first catalytic cracking catalyst and carrying out a first catalytic cracking reaction; wherein the main riser reactor is also provided with a second reaction zone positioned above the first reaction zone, the second reaction zone is in fluid communication with the first reaction zone, and the inner diameter of the second reaction zone is larger than that of the first reaction zone;

feeding the oil agent obtained by the reaction in the first reaction zone into the second reaction zone to contact with the raffinate oil and perform a second catalytic cracking reaction to obtain a first reaction product and a first catalyst to be generated;

introducing the desorbed oil into the secondary riser reactor to contact with a second catalytic cracking catalyst and perform a third catalytic cracking reaction to obtain a second reaction product and a second spent catalyst;

separating the first reaction product and the second reaction product to obtain at least propylene and butylene;

and feeding the obtained first spent catalyst and second spent catalyst into a regenerator for coke burning regeneration, and returning at least part of the obtained regenerated catalyst to the main riser reactor and the auxiliary riser reactor to be used as the first catalytic cracking catalyst and the second 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 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 of the desorption gas is higherSpace velocity of 100-200 hours^-1。

Optionally, the ratio of the height of the first reaction zone to the height of the second reaction zone is 0.5-2, and the ratio of the inner diameter of the second reaction zone to the inner diameter of the first reaction zone is 1.2-3: 1.

Optionally, the conditions of the first catalytic cracking reaction include: the reaction temperature is 500-650 ℃, 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 490-600 ℃, and the reaction time is 1-10 seconds;

the conditions of the third 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 1-100, and the weight ratio of the water to the oil is (0.01-1): 1.

optionally, the conditions of the first catalytic cracking reaction include: the reaction temperature is 510-600 ℃, the reaction time is 1-5 seconds, and the weight ratio of the catalyst to the oil is 5-30;

the conditions of the second catalytic cracking reaction include: the reaction temperature is 500-580 ℃, and the reaction time is 2-8 seconds;

the conditions of the third catalytic cracking reaction include: the reaction temperature is 580-730 ℃, the reaction time is 1-8 seconds, and the weight ratio of the catalyst to the oil is 10-50.

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 an auxiliary riser reactor and a main riser reactor, wherein the temperatures of the preheated desorption oil and the preheated residual absorption oil are respectively 350-450 ℃.

Optionally, the first catalytic cracking catalyst and the second catalytic cracking catalyst each independently comprise, on a dry basis and based on the total weight of the catalyst, from 1 to 60 weight percent zeolite, from 5 to 99 weight percent inorganic oxide, and from 0 to 70 weight percent 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: the regenerated catalyst from the regenerator is delivered into a degassing tank for degassing and then delivered into the main riser reactor and/or the auxiliary riser reactor, and oxygen-containing gas obtained by degassing in the degassing tank is returned into the regenerator.

The invention also provides a processing system of the coker gasoline and the heavy raw oil, which comprises an adsorption-desorption reactor, a main riser reactor, an auxiliary riser reactor, oil separation equipment and a regenerator, wherein the main riser reactor is provided with a first reaction zone and a second reaction zone positioned above the first reaction zone, the second reaction zone is communicated with the first reaction zone in a fluid manner, and the inner diameter of the second reaction zone is larger than that of the first reaction zone;

the adsorption and desorption reactor is provided with a coking gasoline raw material inlet, a desorption gas inlet, an oil absorption residual outlet and a desorption oil outlet, the first reaction zone is provided with a heavy raw oil feed inlet and a catalyst inlet, the second reaction zone is provided with an oil absorption residual oil feed inlet and an oil agent outlet, the auxiliary riser reactor is provided with a desorption oil feed inlet, a catalyst inlet and an oil agent outlet, the oil agent separation equipment is provided with an oil agent inlet, a catalyst outlet and 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 the second reaction zone inhale residual oil feed inlet fluid intercommunication, the desorption oil export of absorption desorption reactor with the desorption oil feed inlet fluid intercommunication of vice riser reactor, the finish outlet of second reaction zone and vice riser reactor's finish export with finish splitter's finish entry fluid intercommunication, the catalyst entry of first reaction zone and vice riser reactor's catalyst entry with the catalyst outlet fluid intercommunication of regenerator, the catalyst entry of regenerator with finish splitter's catalyst outlet fluid intercommunication.

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;

the ratio of the height of the first reaction zone to the height of the second reaction zone is (0.5-2): 1, the ratio of the inner diameters of the second reaction zone to the first reaction zone is (1.2-3): 1.

optionally, the system further comprises a degassing tank through which the catalyst inlet of the first reaction zone and/or the catalyst inlet of the secondary 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 carry out cracking reaction by preferentially contacting the heavy raw oil with the regenerated catalyst, so that a small amount of carbon is deposited on the regenerated catalyst, the property of the catalyst is reasonably adjusted, the number of strong acid centers on the catalyst is reduced, proper catalyst acidity is provided for non-normal alkanes with higher reaction activity, especially olefins, and the yield and the selectivity of propylene and butylene are improved.

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 first reaction zone 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 regeneration pipe chute

13 regenerative spool valve 14 line 15 line

16 line 17 line 18 line

19 line 20 large oil-gas line 21 line

22 air distributor 23 line 24 cyclone

25 flue gas duct 26 second reaction zone 27 line

28 line 29 line 30 line

31 adsorption and desorption reactor 32 heating furnace 33 pipeline

34 pairs of riser reactors 35 line 36 line

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:

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 heavy raw oil into a first reaction zone of a main riser reactor to contact with a first catalytic cracking catalyst and carrying out a first catalytic cracking reaction; wherein the main riser reactor is also provided with a second reaction zone positioned above the first reaction zone, the second reaction zone is in fluid communication with the first reaction zone, and the inner diameter of the second reaction zone is larger than that of the first reaction zone;

feeding the oil agent obtained by the reaction in the first reaction zone into the second reaction zone to contact with the raffinate oil and perform a second catalytic cracking reaction to obtain a first reaction product and a first catalyst to be generated;

introducing the desorbed oil into the secondary riser reactor to contact with a second catalytic cracking catalyst and perform a third catalytic cracking reaction to obtain a second reaction product and a second spent catalyst;

separating the first reaction product and the second reaction product to obtain at least propylene and butylene;

and feeding the obtained first spent catalyst and second spent catalyst into a regenerator for coke burning regeneration, and returning at least part of the obtained regenerated catalyst to the main riser reactor and the auxiliary riser reactor to be used as the first catalytic cracking catalyst and the second catalytic cracking 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, the riser reactor is well known to those skilled in the art, the main riser reactor in the present invention is a variable diameter riser, such as the one described in patent application CN1237477A, the ratio of the height of said first reaction zone to the height of the second reaction zone is preferably (0.5-2): 1, the ratio of the internal diameter of the second reaction zone to the internal diameter of the first reaction zone being (1.2-3) to 1, more preferably (1.5-2): 1, the heavy raw oil and the raffinate may be fed in the lower portions of the first reaction zone and the second reaction zone, respectively.

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, the same below) is 500-650 ℃, the reaction time is 1-10 seconds, the reaction pressure (gauge pressure, the same below) is 0.05-1 MPa, 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 first catalytic cracking reaction preferably include: the reaction temperature is 510-600 ℃, the reaction time is 1-5 seconds, and the weight ratio of the catalyst to the oil is 5-30; the conditions of the second catalytic cracking reaction include: the reaction temperature is 490-600 ℃, the reaction time is 1-10 seconds, and the 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 500-580 ℃, and the reaction time is 2-8 seconds; the conditions of the third catalytic cracking reaction may include: the reaction temperature is 560 ℃ and 750 ℃, the reaction pressure is 0.05-1 MPa, the reaction time is 1-10 seconds, the weight ratio of the catalyst to the oil is 1-100, and the weight ratio of the water to the oil is (0.01-1): 1; the conditions of the third catalytic cracking reaction preferably include: the reaction temperature is 580-730 ℃, the reaction time is 1-8 seconds, and the weight ratio of the catalyst to the oil is 10-50.

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 auxiliary riser reactor and the main riser 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.

According to the present invention, catalytic cracking catalysts are well known to those skilled in the art, and in particular for the purposes of the present invention, the first catalytic cracking catalyst and the second catalytic cracking catalyst may each independently comprise1-60 wt% zeolite, 5-99 wt% inorganic oxide and 0-70 wt% clay; 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 can be 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 the main riser reactor and/or the auxiliary riser reactor, oxygen-containing gas obtained by degassing in the degassing tank is returned to the regenerator, and impurities such as oxygen and the like in the regenerated catalyst in the degassing tank are removed 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.

The invention also provides a processing system of the coker gasoline and the heavy raw oil, which comprises an adsorption-desorption reactor, a main riser reactor, an auxiliary riser reactor, oil separation equipment and a regenerator, wherein the main riser reactor is provided with a first reaction zone and a second reaction zone positioned above the first reaction zone, the second reaction zone is communicated with the first reaction zone in a fluid manner, and the inner diameter of the second reaction zone is larger than that of the first reaction zone;

the adsorption and desorption reactor is provided with a coking gasoline raw material inlet, a desorption gas inlet, an oil absorption residual outlet and a desorption oil outlet, the first reaction zone is provided with a heavy raw oil feed inlet and a catalyst inlet, the second reaction zone is provided with an oil absorption residual oil feed inlet and an oil agent outlet, the auxiliary riser reactor is provided with a desorption oil feed inlet, a catalyst inlet and an oil agent outlet, the oil agent separation equipment is provided with an oil agent inlet, a catalyst outlet and 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 the second reaction zone inhale residual oil feed inlet fluid intercommunication, the desorption oil export of absorption desorption reactor with the desorption oil feed inlet fluid intercommunication of vice riser reactor, the finish outlet of second reaction zone and vice riser reactor's finish export with finish splitter's finish entry fluid intercommunication, the catalyst entry of first reaction zone and vice riser reactor's catalyst entry with the catalyst outlet fluid intercommunication of regenerator, the catalyst entry of regenerator with finish splitter's catalyst outlet fluid intercommunication.

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 are well known to those skilled in the art, the main riser reactor in the present invention is a variable diameter riser, and the ratio of the height of the first reaction zone to the height of the second reaction zone is preferably (0.5-2): 1, more preferably (0.2-0.8): 1, the ratio of the internal diameter of the second reaction zone to the internal diameter of the first reaction zone being (1.2-3) to 1, more preferably (1.5-2): 1, the heavy raw oil and the raffinate may be fed in the lower portions of the first reaction zone and the second reaction zone, respectively. It is well known to those skilled in the art that a pre-lift section may also be provided below the first reaction zone and an outlet section may also be provided above the second reaction zone, the pre-lift section having an internal diameter generally smaller than the internal diameter of the first reaction zone and the outlet section having an internal diameter generally smaller than the internal diameter of the second reaction zone.

According to the present invention, the system may further comprise a degassing tank, through which the catalyst inlet of the first reaction zone and/or the catalyst inlet of the secondary riser reactor is in fluid communication with the catalyst outlet of the regenerator, 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 of which is in communication with the catalyst outlet of the regenerator, the catalyst outlet being in communication with the catalyst inlet of the reactor, the stripping gas inlet being for introducing a stripping gas such as water vapor, the stripping gas outlet being in communication with the regenerator for feeding an 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-desorption reactor 31 through a pipeline 27, and after adsorption and separation by an adsorbent, 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 31 through the pipeline 28 to make the molecular sieve desorb the adsorbed normal paraffin components, and the produced desorption oil enters the heating furnace 32 through the pipeline 29 to be preheated.

The pre-lifting medium enters from the bottom of the main riser reactor through a pipeline 14, the regenerated catalyst from a regenerated inclined pipe 12 enters into a first reaction zone 1 of the main riser reactor after being regulated by a regeneration slide valve 13, and moves upwards in an accelerated manner along the main riser reactor under the lifting action of the pre-lifting medium, the preheated heavy raw oil is injected into the first reaction zone 1 of the main riser reactor through a pipeline 36 together with the atomized steam from a pipeline 35, and is mixed with the existing material flow of the main riser reactor, the raw oil is subjected to catalytic cracking reaction on the hot catalyst and moves upwards in an accelerated manner and enters into a second reaction zone 26. The preheated raffinate oil is injected into a second reaction zone 26 of the main riser reactor together with the atomized steam from the pipeline 17 through the pipeline 18, is mixed with the existing material flow of the main riser reactor and generates catalytic cracking reaction; preheated desorption oil is injected into the bottom of the secondary riser reactor 34 together with atomized steam from the pipeline 15 through the pipeline 16, and is contacted with the regenerated catalyst injected through the pipeline 33 to carry out catalytic cracking reaction, a reaction product generated by the reaction and the inactivated spent catalyst enter the cyclone separator 6 in the settler 3 to realize the separation of the spent catalyst and the reaction product, the reaction product enters the gas collection chamber 7, and catalyst fine powder 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 main riser reactor 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 raw oils used in the examples and comparative examples were Daqing atmospheric residue and coker gasoline, and their properties are shown in tables 1 and 2.

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

Comparative example 1

The test is carried out on a medium-sized device of a riser reactor, heavy raw oil is Daqing atmospheric residue, preheated Daqing atmospheric residue enters the bottom of a main riser reactor, tar gasoline enters the bottom of an auxiliary riser reactor to carry out catalytic cracking reaction respectively, the feeding weight ratio of coking gasoline to Daqing atmospheric residue in unit time is 0.15:1, and the height ratio of a first reaction zone to a second reaction zone of the main riser reactor is 1: 1, the ratio of the inner diameter of the second reaction zone to the inner diameter of the first reaction zone is 1.5: 1. and the reaction product and the spent catalyst enter a closed cyclone separator from the outlet of the reactor, the reaction product and the spent catalyst are quickly separated, and the reaction product is cut in a separation system according to the distillation range, so that fractions such as propylene, butylene, pyrolysis gasoline and the like are obtained.

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 11.65 wt% and 13.84 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, and the height ratio of the first reaction zone to the second reaction zone is 1: 1, the ratio of the inner diameter of the second reaction zone to the inner diameter of the first reaction zone is 1.5: 1. the preheated heavy raw oil and the residual oil respectively enter a first reaction zone and a second reaction zone of the main riser reactor to carry out a first catalytic cracking reaction and a second catalytic cracking reaction, and the preheated desorption oil enters the bottom of the auxiliary riser reactor to carry out a third catalytic cracking reaction; the feed weight ratio of the desorption oil, the absorption residual oil and the heavy oil raw oil in unit time is 0.05: 0.1: 1. cutting the reaction product in a separation system according to the distillation range to obtain fractions such as butylene, propylene, 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 18.94 wt% and 16.53 wt%, respectively.

Comparative example 2

The same as example 1 except that: the feed inlets 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 auxiliary riser reactor, and the desorbed oil was fed from the second reaction zone of the main riser 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, heavyThe content of
		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
				Main riser reactor
Outlet temperature of the first reaction zone,. degree.C	530	530	530
				First reaction zone reaction time, second	2	2	2
Water-to-oil weight ratio in the first reaction zone	0.1	0.1	0.1
				First reaction zone catalyst oil weight ratio	8	8	8
Outlet temperature of the second reaction zone,. degree.C	515	515	515
				Second reaction zone reaction time, second	2.5	2.5	2.5
Secondary riser reactor
				Outlet temperature of	635	635	635
Reaction time in seconds	1.8	1.8	1.8
				Water to oil weight ratio	0.2	0.2	0.2
Weight ratio of solvent to oil	25	25	25
				Product distribution, weight%
Dry gas	4.08	4.48	4.25
				LPG	31.19	39.33	37.37
Gasoline (gasoline)	39.63	32.13	33.2
				Diesel oil	12.07	11.39	11.96
Heavy oil	4.28	4.24	4.37
				Coke	8.75	8.43	8.85
Total up to	100	100.00	100
				Propylene (PA)	11.65	18.94	17.05
Butene (butylene)	13.84	16.53	14.88
				Conversion, wt.%	83.65	84.37	83.67

Claims (16)

1. A process for processing coker gasoline and heavy feed oil, the process 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 heavy raw oil into a first reaction zone of a main riser reactor to contact with a first catalytic cracking catalyst and carrying out a first catalytic cracking reaction; wherein the main riser reactor is also provided with a second reaction zone positioned above the first reaction zone, the second reaction zone is in fluid communication with the first reaction zone, and the inner diameter of the second reaction zone is larger than that of the first reaction zone;

feeding the oil agent obtained by the reaction in the first reaction zone into the second reaction zone to contact with the raffinate oil and perform a second catalytic cracking reaction to obtain a first reaction product and a first catalyst to be generated;

introducing the desorbed oil into the secondary riser reactor to contact with a second catalytic cracking catalyst and perform a third catalytic cracking reaction to obtain a second reaction product and a second spent catalyst;

separating the first reaction product and the second reaction product to obtain at least propylene and butylene;

feeding the first spent catalyst and the second spent catalyst into a regenerator for coke burning regeneration, and returning at least part of the regenerated catalyst to the main riser reactor and the auxiliary riser reactor to be used as the first catalytic cracking catalyst and the second catalytic cracking catalyst;

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

2. The method of claim 1, wherein,

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.

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, 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；

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。

6. The process of claim 1, wherein the ratio of the height of the first reaction zone to the second reaction zone is from 0.5 to 2 and the ratio of the inner diameter of the second reaction zone to the first reaction zone is from 1.2 to 3: 1.

7. The method of claim 1, wherein the conditions of the first catalytic cracking reaction comprise: the reaction temperature is 500-650 ℃, 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 490-600 ℃, and the reaction time is 1-10 seconds;

the conditions of the third 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 1-100, and the weight ratio of the water to the oil is (0.01-1): 1.

8. the method of claim 1, wherein the conditions of the first catalytic cracking reaction comprise: the reaction temperature is 510-600 ℃, the reaction time is 1-5 seconds, and the weight ratio of the catalyst to the oil is 5-30;

the conditions of the second catalytic cracking reaction include: the reaction temperature is 500-580 ℃, and the reaction time is 2-8 seconds;

the conditions of the third catalytic cracking reaction include: the reaction temperature is 580-730 ℃, the reaction time is 1-8 seconds, and the weight ratio of the catalyst to the oil is 10-50.

9. The method of claim 1, wherein the method further comprises: preheating desorption oil and residual absorption oil, and then respectively introducing the preheated desorption oil and the preheated residual absorption oil into an auxiliary riser reactor and a main riser 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 first and second catalytic cracking catalysts each independently comprise, 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.

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: the regenerated catalyst from the regenerator is delivered into a degassing tank for degassing and then delivered into the main riser reactor and/or the auxiliary riser reactor, and oxygen-containing gas obtained by degassing in the degassing tank is returned into the regenerator.

14. A processing system for coker gasoline and heavy raw oil comprises an adsorption-desorption reactor, a main riser reactor, an auxiliary riser reactor, oil separation equipment and a regenerator, wherein the main riser reactor is provided with a first reaction zone and a second reaction zone positioned above the first reaction zone, the second reaction zone is communicated with the first reaction zone in a fluid manner, and the inner diameter of the second reaction zone is larger than that of the first reaction zone;

the adsorption and desorption reactor is provided with a coking gasoline raw material inlet, a desorption gas inlet, an oil absorption residual outlet and a desorption oil outlet, the first reaction zone is provided with a heavy raw oil feed inlet and a catalyst inlet, the second reaction zone is provided with an oil absorption residual oil feed inlet and an oil agent outlet, the auxiliary riser reactor is provided with a desorption oil feed inlet, a catalyst inlet and an oil agent outlet, the oil agent separation equipment is provided with an oil agent inlet, a catalyst outlet and 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 the second reaction zone inhale residual oil feed inlet fluid intercommunication, the desorption oil export of absorption desorption reactor with the desorption oil feed inlet fluid intercommunication of vice riser reactor, the finish outlet of second reaction zone and vice riser reactor's finish export with finish splitter's finish entry fluid intercommunication, the catalyst entry of first reaction zone and vice riser reactor's catalyst entry with the catalyst outlet fluid intercommunication of regenerator, the catalyst entry of regenerator with finish splitter's catalyst outlet fluid intercommunication.

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;

the ratio of the height of the first reaction zone to the height of the second reaction zone is (0.5-2): 1, the ratio of the inner diameters of the second reaction zone to the first reaction zone is (1.2-3): 1.

16. the system of claim 14, wherein the system further comprises a degassing tank through which a catalyst inlet of the first reaction zone and/or a catalyst inlet of a secondary riser reactor is in fluid communication with a 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.

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