CN113117725A

CN113117725A - Solid acid alkylation process

Info

Publication number: CN113117725A
Application number: CN201911402571.7A
Authority: CN
Inventors: 李永祥; 张成喜; 胡合新; 赵志海; 罗一斌; 张久顺; 侯栓弟
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: 2019-12-31
Filing date: 2019-12-31
Publication date: 2021-07-16
Anticipated expiration: 2039-12-31
Also published as: CN113117725B

Abstract

A solid acid alkylation process comprising an alkylation reaction unit and an alkylation reaction product separation unit, wherein the alkylation reaction unit is adapted to catalyze an alkylation reaction with a solid acid catalyst having: (1) the specific volume of the macropores is 0.30-0.40 ml/g; (2) the ratio of the specific volume of the macropores to the specific length of the catalyst particles is 1.0-2.5 ml/(g.mm); (3) the ratio of the specific surface area to the length of the particles is 3.40 to 4.50m²Per mm; the macropores refer to pores with the diameter of more than 50 nm; in the alkylation reaction product separation unit, a compressor is arranged at the top of the fractionating tower, a reboiler is arranged at the middle section of the fractionating tower, the temperature and pressure of the gas phase at the top of the fractionating tower are improved after the gas phase is compressed by the compressor, and the gas phase enters the reboiler at the middle section of the fractionating tower and is led out from the middle lower part of the fractionating towerAnd (4) exchanging heat of the liquid phase material flow with lower temperature, and liquefying the gas phase material flow at the top of the tower in a reboiler at the middle section of the tower.

Description

Solid acid alkylation process

Technical Field

The invention relates to a solid acid alkylation method, in particular to a method for producing alkylated gasoline by solid acid alkylation of isoparaffin and olefin, belonging to the field of petrochemical industry.

Background

Isoparaffins of predominantly isobutane with C₃～C₆The alkylation reaction of olefin under the action of acid catalyst is an important process for producing clean and high-octane gasoline component. The alkylated gasoline obtained by the process has low steam pressure, low sensitivity, good anti-knock performance, no aromatic hydrocarbon and olefin, and low sulfur content, and is an ideal blending component of high-octane gasoline.

The alkylation reaction is an acid-catalyzed reaction. The alkylation processes currently used in industry include sulfuric acid and hydrofluoric acid, which are processes for synthesizing alkylate from isoparaffin and olefin using sulfuric acid or hydrofluoric acid as a catalyst. Because of the corrosivity and toxicity of sulfuric acid and hydrofluoric acid and the harm of waste acid discharge in the process to the environment, the pressure of safety and environmental protection for alkylate oil production enterprises is increasing day by day.

The core of the solid acid alkylation process is the development of a solid acid catalyst with excellent performance. The solid acid catalyst has many advantages, such as good stability, no corrosion to equipment, convenient separation from products, little environmental pollution, relatively high safety in the transportation process, etc., and is an ideal catalyst form. The alkylation solid acid catalyst is mainly divided into four types: metal halide, solid super acid, supported heteropoly acid and molecular sieve. Despite decades of development, the progress of industrialization of this process technology is affected due to the rapid deactivation of the solid acid catalyst.

US5986158 discloses an alkylation process using a catalyst comprising a hydrogenation function and a solidThe acid component, alkylatable compound and alkylating agent are reacted to form alkylate, and the catalyst is regenerated intermittently under the condition of saturated hydrocarbon and hydrogen, and the catalyst is regenerated in the active period of the catalyst in which the conversion rate of alkylating agent is greater than 80%, so as to obtain high alkylation yield and high-quality alkylated product. The zeolite catalyst is a porous catalytic material, and impurities in the raw materials can be adsorbed on the active center of the catalyst or macromolecules generated by polymerization of alkadiene in the raw materials can block pore channels, so that the catalyst is inactivated. The reaction process is carried out in a fixed bed reactor, the active period of the catalyst is only less than 4-10 h, the catalyst needs to be regenerated repeatedly, and the examples show that the Research Octane Number (RON) of the alkylate oil is 91.2, the Research Octane Number (RON) of trimethylpentane/dimethylhexane is 2.9, and C is₅-C₇、C₈、C₉₊30.4%, 58.2% and 11.4%, respectively.

EP1286769 discloses a novel alkylation catalyst and its use for the alkylation of hydrocarbons.

EP1392627 discloses a process for the catalytic alkylation of hydrocarbons which comprises (i) reacting an alkylatable compound with an alkylating agent over a solid acid alkylation catalyst to form an alkylate and (ii) regenerating said catalyst under mild regeneration conditions and in the presence of hydrogen and a hydrocarbon, wherein the hydrocarbon comprises at least a portion of the alkylate that has been formed.

Although the solid acid catalysts have certain catalytic performance, the catalytic activity, selectivity and stability of the catalysts still need to be further improved, the regeneration problem of the catalysts is solved, and the quality of the alkylated gasoline is improved. Due to the characteristic of porosity of the solid acid alkylation catalyst, alkylation reaction on an active site of the solid acid catalyst requires that raw materials are diffused to the active site from a fluid main body, products are diffused to the main body from the active site, the adsorption capacity of olefin on the active site is stronger, the alkylation reaction effect can be ensured only by requiring higher external alkane-olefin ratio, and thus, the energy consumption of product separation is higher.

Disclosure of Invention

The invention aims to provide a solid acid alkylation method for improving the performance of a catalyst and the quality of alkylated gasoline, reducing the energy consumption of an alkylation unit and reducing the operation cost of the alkylation unit.

The solid acid alkylation method provided by the invention comprises an alkylation reaction unit and an alkylation reaction product separation unit, and is characterized in that:

in the alkylation reaction unit, a solid acid catalyst is used for catalyzing alkylation reaction, and the solid acid catalyst is provided with:

(1) the specific volume of the macropores is 0.30-0.40 ml/g;

(2) the ratio of the specific volume of the macropores to the specific length of the catalyst particles is 1.0-2.5 ml/(g.mm);

(3) the ratio of the specific surface area to the length of the particles is 3.40 to 4.50m²Per mm; and the macropores refer to pores with the diameter of more than 50 nm;

in the alkylation reaction product separation unit, a compressor is arranged at the top of the fractionating tower, a reboiler is arranged at the middle section of the fractionating tower, the temperature and the pressure of the gas phase at the top of the fractionating tower are improved after the gas phase at the top of the fractionating tower is compressed by the compressor, the gas phase at the top of the fractionating tower enters the reboiler at the middle section of the fractionating tower to exchange heat with the liquid phase material flow with lower temperature led out from the middle lower.

The invention has the following advantages:

(1) the solid acid catalyst of the invention has high catalytic activity, can further improve the selectivity, solve the regeneration problem of the catalyst and improve the quality of the alkylated gasoline.

(2) By adopting the alkylation product separation unit provided by the invention, the gas phase at the top of the fractionating tower is pressurized by the gas compressor, so that the temperature and the pressure of the gas phase at the top of the fractionating tower are improved, the gas phase at the top of the fractionating tower can be used as a heat source of a reboiler at the middle section of the fractionating tower, the phase change heat during liquefaction of the gas phase at the top of the fractionating tower is fully utilized, the total energy consumption of an alkylation device can be effectively reduced, and the aim of reducing the operation cost of the alkylation device is fulfilled.

Drawings

FIG. 1 is a schematic process flow diagram of a solid acid alkylation process;

FIG. 2 is a schematic process flow diagram of another solid acid alkylation process.

FIG. 3 is an SEM and energy spectrum surface scanning of a solid acid catalyst sample to characterize the morphology and element distribution of the solid acid catalyst.

In the figure, 1 is alkylation raw material, 2 a-2 n are reactors, 3 is reacted material, 4 and 18 are fractionating towers, 5 is fractionating tower top outlet material, 6 is a stabilization tank, 7 is fractionating tower top material, 8 is a compressor, 9 is pressurized material, 10, 16 and 25 are reboilers, 11 is condensed material, 12 and 22 are tower top reflux tanks, 13 and 23 are reflux materials, 14 and 24 are light hydrocarbon fraction byproducts at the top of the fractionating tower, 15 is circulating isobutane, 17 is tower bottom material of the fractionating tower, 26 is alkylated product, 19 is fractionating tower top material, 20 is a cooler, 21 is tower top condensation cooling material, 27 a-27 m is segmented entering reactor material, 28 is reactor outlet material, 29 is circulating pump, and 30 is reactor outlet circulating material.

Detailed Description

A solid acid alkylation method comprises an alkylation reaction unit and an alkylation reaction product separation unit, and is characterized in that,

(1) the specific volume of the macropores is 0.30-0.40 ml/g;

The alkylation reaction unit adopts a solid acid catalyst with special physical and chemical properties, and comprises solid acid catalyst particles and a regeneration auxiliary agent with a hydrogenation function, so that alkylation reaction and catalyst regeneration can be realized.

The International Union of Pure and Applied Chemistry (IUPAC) specifies that pores with a diameter greater than 50nm are denoted by "macropore", and the volume in the pores of such pores is denoted by "macropore volume". The macropore specific volume refers to the volume of macropores per unit mass of catalyst particles. The specific volume of the macropores is 0.30-0.40 ml/g, preferably at least 0.35 ml/g. Catalyst particle specific length refers to the ratio of the geometric volume to the geometric surface of the solid portion of the catalyst particle. Methods for determining geometric volumes and geometric surfaces are well known to the person skilled in the art and can be determined, for example, as described in DE 2354558. It is to be noted that the specific length of the catalyst particles is different from the diameter of the catalyst particles. For example, for cylindrical catalyst particles, the diameter ratio of the particles is four to six times greater than the length (depending on the diameter and length of the particles), and for spherical catalyst particles, the diameter ratio of the particles is six times greater than the length. The solid acid catalyst has a particle specific length of preferably 0.15 to 0.4mm, more preferably 0.18 to 0.36mm, and most preferably 0.20 to 0.32 mm. The ratio of the specific volume of the macropores to the specific length of the catalyst particles is 1.0 to 2.5 ml/g.mm, preferably 1.1 to 1.8 ml/g.mm. The total pore volume refers to the total pore volume per unit mass of catalyst particles of the solid acid catalyst, which is at least 0.40ml/g, preferably at least 0.45 ml/g. The invention is based on the Washbum equation, and the volume of the macropore and the total pore volume are measured by a mercury intrusion method: d (-4 γ cos θ)/p where D is the pore diameter, p is the pressure applied during the measurement, γ is the surface tension, taking 485 dynes/cm, θ is the contact angle, taking 130 °.

The catalyst particles can have many different shapes including spherical, cylindrical, toroidal, and symmetrical or asymmetrical multi-lobed shapes (e.g., butterfly, trilobe, quadralobe). The average diameter of the catalyst particles is preferably at least 1.0mm, and its upper limit value is preferably 5.0 mm. The average diameter of the catalyst particles refers to the longest line segment among line segments connecting any two points on the cross section of one catalyst particle, and can be measured by a conventional measuring means such as a vernier caliper.

The solid acid component of the catalyst is preferably a molecular sieve. The molecular sieve may be selected from a variety of molecular sieves, for example, may be one or more selected from Y-type molecular sieves, beta, MOR, MCM-22 and MCM-36. The unit cell size of the Y-type molecular sieve is 2.430-2.470nm, the preferable unit cell size is 2.440-2.460 nm, and the molar ratio of silicon dioxide to aluminum oxide is 5-15. If desired, the solid acid component may also include non-zeolitic solid acids such as heteropolyacids, silica-aluminas, sulfated oxides such as sulfated oxides of zirconium, titanium or tin, mixed oxides of zirconium, molybdenum, tungsten, phosphorus, or the like, chlorinated aluminas or clays, and the like.

The catalyst also contains a matrix material. The content of the base material is 2-98 wt%, and the preferable content is 10-70 wt%. The catalyst preferably comprises 2 to 98 wt% of the solid acid component and 2 to 98 wt% of the base material, further preferably comprises 5 to 95 wt% of the solid acid component and 5 to 95 wt% of the base material, more preferably comprises 15 to 85 wt% of the solid acid component and 15 to 85 wt% of the base material, may comprise 20 to 80 wt% of the solid acid component and 20 to 80 wt% of the base material, or may comprise 60 to 80 wt% of the solid acid component and 20 to 40 wt% of the base material, based on the total weight of the solid acid component and the base material present in the catalyst. Wherein the matrix material comprises alumina, and the precursor of the alumina is at least partially from alumina sol with the granularity of 20-400 nm.

The specific surface area of the catalyst is not less than 500m²(ii) in terms of/g. The solid acid component is highly dispersed in the base material in micron level, and the specific surface area of the solid acid component is not less than 650m²The specific surface area of the base material is not more than 400m²(ii) in terms of/g. After the solid acid component is dispersed in the substrate material in a micron-level height, the specific surface area of the catalyst particles per unit length is required to fluctuate within a narrow range, and the large change caused by the large difference between the specific surface area of the solid acid component and the specific surface area of the substrate material is avoided. The ratio of the specific surface area of the catalyst to the length of the particles is 3.40 to 4.50m²And/mm. The particle length is obtained by randomly selecting 1g of catalyst particles, measuring the length of each of the 1g of catalyst particles, and adding the lengths of each of the particles. For spherical particles, the particle length is the diameter of the sphere; for a particle in the form of a rod (including butterfly, trilobe, and quadralobe cross-sections, among others), the length of the particle is the average rod of the particleLength; for annular particles, the particle length is the outer diameter of the annulus.

The preparation method of the catalyst comprises the steps of mixing and stirring slurry containing solid acid components and aluminum sol uniformly, drying, mixing with extrusion assistant and peptizing agent, and forming, wherein the particle size of the aluminum sol is 20-400 nm. In the preparation method, the particle size of the aluminum sol is 20-400nm, preferably 20-300 nm. The extrusion aid is well known to those skilled in the art, and the commonly used extrusion aid is selected from sesbania powder, oxalic acid, tartaric acid, citric acid and the like, preferably sesbania powder; the peptizing agent is also well known to those skilled in the art, and commonly used peptizing agents are selected from the group consisting of nitric acid, hydrochloric acid, acetic acid, formic acid, citric acid, trichloroacetic acid, and the like, preferably nitric acid.

After the regeneration auxiliary agent with the hydrogenation function is loaded, the catalyst can be regenerated under the condition of inactivation and the hydrogen and proper conditions, so that the repeated regeneration and the recycling of the catalyst are realized. Therefore, the catalyst adopted by the alkylation reaction unit of the invention comprises a regeneration auxiliary agent component formed by metal with hydrogenation function. Suitable hydrogenation-functional metals are mainly group VIII metals, preferably group VIII noble metals. More preferably, the group VIII noble metal is one or more of rhodium, palladium and platinum. The content of the metal having a hydrogenation function is 0.01 to 10 wt%, preferably 0.1 to 1 wt%, in terms of metal, based on the weight of the alkylation catalyst. Typical preparation steps include impregnation of the particles by a solution containing the hydrogenation function metal and/or addition of the hydrogenation function metal to the solid acid catalyst described above by ion exchange; a typical preparation procedure may also be to add a precursor of the metal having the hydrogenation function to a liquid phase mixture comprising the solid acid component and an aluminum sol having a particle size of 20 to 400nm, dry the resulting mixture and shape it.

The alkylation reaction unit of the invention is filled with the catalyst. The alkylation reactor may be in the form of various types of reactors such as fluidized bed reactors, slurry bed reactors, loop reactors, and fixed bed reactors. The process can also be carried out in single and multiple reactors.

In the alkylation reaction, the catalyst is filled in the reactor, particularly the alkylation reaction is carried out in a fixed bed reactor, the reaction can be carried out at the temperature of 5-200 ℃, preferably 20-150 ℃, more preferably 30-100 ℃, the pressure of 0.5-6.0 MPa, preferably 1.0-3.0 MPa, and the olefin feeding mass space velocity of the alkylation reactor is 0.02-2.0 h^-1Preferably 0.05 to 0.5 hour^-1The molar ratio of isoparaffin to olefin in the total feed is 5-50, preferably 7-30. The pretreated raw materials enter the alkylation reactors in multiple sections, the alkylation reactors are switched in a reaction-regeneration mode, and a part of materials at the outlet of the reactors can be internally circulated back to the inlet of the reactors. In the method, the material entering the reactor can be divided into a plurality of sections and enter the reactor, and the materials are fed in a plurality of points, so that the total alkane-alkene ratio can be reduced, and the energy consumption of the fractionating tower can be reduced. The segmentation is at least 2 segments and at most 20 segments, and too many segments make the internal structure of the reactor more complicated, and 3 to 12 segments are preferred. Part of materials at the outlet of the reactor are circulated, so that the high internal alkane-alkene ratio can be kept, the total alkane-alkene ratio can be kept unchanged, and the total fractionation energy consumption is reduced.

The invention is provided with a plurality of reactors for reaction-regeneration switching, and the number of the reactors is 2-10, preferably 2-6. The regeneration of the catalyst related by the invention is on-line in-situ regeneration in an alkylation reactor, and the regeneration conditions of the catalyst are as follows: the temperature is 30-500 ℃, the pressure is 0.5-5.0 MPa, and the time is 0.5-15 h; preferably, the temperature is 150-300 ℃, the pressure is 0.5-3.0 MPa, and the time is 0.5-6 h.

The alkylation reaction product separation unit of the process of the present invention may be operated with one or two fractionation columns.

For a fractionating tower, an alkylation reaction product from an alkylation reaction unit enters from the middle of the fractionating tower, after an overhead steam outlet is communicated with an inlet of a gas compressor for pressurization, the temperature and the pressure of a material at an outlet of the gas compressor are both increased and used as a heat source of a middle reboiler of the fractionating tower, the material is subjected to phase change in the middle reboiler of the fractionating tower and exchanges heat with a liquid phase material flow with lower temperature led out from the middle lower part of the fractionating tower and then returns to the fractionating tower, the liquid phase after phase change enters a top reflux tank, one part of the liquid phase is used as reflux of the fractionating tower, the other part of the liquid phase is used as light fraction at the top of the fractionating tower, a bottom reboiler is arranged at the bottom of the fractionating.

In the case of two fractionation columns,

(1) introducing an alkylation reaction product from the alkylation reaction unit into a first fractionating tower for fractionation and separation, introducing a gas phase substance at the top of the first fractionating tower, which is used as a heat source of a middle-section reboiler of the first fractionating tower after the gas phase substance flows through a gas compressor and is pressurized, introducing the gas phase substance into a top reflux tank after heat exchange and condensation, and returning a part of the gas phase substance as top reflux to the first fractionating tower;

(2) introducing the liquid phase material flow at the bottom of the first fractionating tower into a second fractionating tower, condensing and cooling the gas phase material flow led out from the top of the second fractionating tower, returning one part of the gas phase material flow to the fractionating tower as the reflux of the top of the fractionating tower, obtaining the light fraction material flow at the other part of the gas phase material flow, and taking the liquid phase material flow at the bottom of the second fractionating tower as the alkylated gasoline product.

More specifically, the alkylation reaction product separation device comprises a first fractionating tower and a second fractionating tower which are sequentially communicated, wherein an alkylation reaction product inlet is formed in the middle of the first fractionating tower, a steam outlet at the top of the first fractionating tower is communicated with an inlet of a gas compressor, and an outlet of the gas compressor is communicated with a reflux inlet at the top of the first fractionating tower through a middle reboiler at the middle of the first fractionating tower and a reflux tank at the top of the first fractionating tower; the middle section reboiler is in the position on the liquid fraction outlet return first fractionating tower middle part, first fractionating tower set up the reboiler at the bottom of the tower, material export intercommunication at the bottom of the tower the raw materials entry at second fractionating tower middle part, second fractionating tower top of the tower set up top of the tower condenser, top of the tower reflux drum to set up the export of light fraction at top of the tower reflux drum, second fractionating tower bottom of the tower set up the reboiler at the bottom of the tower to set up the export of alkylate product.

Preferably, the reflux drum at the top of the first fractionating tower is also provided with a light fraction outlet. Preferably, the middle liquid phase fraction outlet of the first fractionating tower is arranged at the position of 20-98% of the first fractionating tower from top to bottom, and preferably 40-80%. Preferably, the middle liquid phase fraction outlet of the first fractionating tower is arranged on the reducing section of the first fractionating tower, the height ratio of the upper expanding section to the lower fractionating tower is 0.25-49: 1, preferably 0.66-4: 1, and the diameter ratio is 1-6: 1. Preferably 2-4: 1. Preferably, the first fractionating tower is an isobutane removing tower, and the second fractionating tower is an n-butane removing tower.

In the method, an alkylation reaction product outlet of the solid acid alkylation reaction unit is communicated with an alkylation reaction product inlet of the alkylation reaction product separation unit.

In the method, a supercharging device is required to be arranged at the top of the deisobutanizer to supercharge the isobutane fraction, the compression ratio of the supercharging device is 1.3-4.5: 1, and the absolute pressure of an outlet is 1.0-3.2 MPa.

The separation in the process of the present invention is carried out by a separation method different from the conventional separation method. The conventional separation method is to directly condense and cool the gas phase material flow at the top of the fractionating tower, and the alkylation reaction needs to adopt a higher external alkane-olefin ratio, so that the flow of the circulating isobutane is multiple times of the flow of the alkylation raw material, the phase change heat of the gas phase material flow at the top of the fractionating tower in the condensation process is considerable, but the temperature of the material flow at the top of the fractionating tower is lower and cannot be used as a heat source, and the heat required by the separation process is completely provided by a reboiler at the bottom of the fractionating tower, so that the heat load of the reboiler at the bottom of the fractionating tower. The separation method adopted by the invention is that the gas compressor is used for pressurizing the overhead gas phase flow of the fractionating tower, the temperature of the pressurized overhead gas phase flow is increased, the pressurized overhead gas phase flow can be used as a heat source of a reboiler at the middle section of the fractionating tower, the overhead gas phase flow is liquefied in the reboiler at the middle section, the phase change heat of the overhead gas phase flow is fully utilized, most of heat required for separating the iso-butane fraction from the alkylation reaction product is provided, and a small amount of heat required for separation is provided by the reboiler at the bottom of the fractionating tower.

The gas compressor is an important device in separation, and applies work to the gas phase at the top of the tower through the gas compressor, so that the pressure and the temperature of the gas phase at the top of the tower are improved, and the gas phase at the top of the tower can meet the requirement of serving as a reboiling heat source at the middle section of the fractionating tower after being pressurized by the compressor.

The content of the second fractional fraction in the alkylation reaction product is small, so that n-butane which does not participate in the reaction in the raw material is mainly removed, the n-butane removing tower adopts a conventional separation method to separate the n-butane in the alkylation reaction product from the alkylated gasoline, a byproduct n-butane fraction is obtained at the tower top, and an alkylated gasoline product is obtained at the tower bottom.

According to the alkylation reaction product separation unit, the top of the fractionating tower is provided with the compressor, the middle section of the fractionating tower is provided with the reboiler, the temperature and the pressure of the gas phase at the top of the tower are improved after the gas phase at the top of the fractionating tower is compressed by the compressor, the gas phase at the top of the tower enters the middle section reboiler to exchange heat with the liquid phase material flow with lower temperature led out from the middle lower part of the fractionating tower, the gas phase material flow at the top of the tower is liquefied in the middle section reboiler of the.

The process of the present invention will be further described with reference to the accompanying drawings, in which only the main equipment and piping are shown, illustrating the main features of the process of the present invention, but not limiting the invention thereto.

Figure 1 illustrates one embodiment of the present invention.

In FIG. 1, after the alkylation raw material 1 is mixed with the circulating isobutane 15 at the top of the fractionating tower 4, the reactor material 27 a-27 m enters the reactors 2 a-2 n in sections (the number of the sections can be a-m, m is generally less than 20, the number of the reactors can be a-n, and n is generally less than 10 according to specific requirements), and the reactors 2 a-2 n are filled with the catalyst of the invention. The method comprises the steps of reacting under alkylation reaction conditions, introducing a material 3 after the reaction into a fractionating tower 4 for fractionation and separation, introducing a gas phase material 5 at the top of the fractionating tower 4 into a gas compressor inlet stabilizing tank 6, introducing a material 7 at the top of the fractionating tower after pressure stabilization into a compressor 8, increasing the temperature and pressure of a material 9 after pressurization, introducing the material into the fractionating tower as a heat source of a middle section reboiler 10, exchanging heat with a liquid phase material flow with lower temperature led out from the middle lower part of the fractionating tower 4, liquefying the gas phase material flow at the top of the fractionating tower in the middle section reboiler 10, fully utilizing the phase change heat of the gas phase material flow at the top of the fractionating tower, introducing a condensed material 11 into a fractionating tower top reflux tank 12, introducing one part of the condensed material into the top of the fractionating tower 4 as a reflux material 13, delivering the other. Insufficient heat in the separation process of the fractionating tower 4 is provided by a fractionating tower bottom reboiler 16, tower bottom materials 17 of the fractionating tower at the bottom of the fractionating tower 4 are introduced into the fractionating tower 18, the fractionating tower 18 adopts a conventional separation method, gas-phase materials 19 of the fractionating tower top materials are condensed and cooled by a cooler 20, then tower top condensed and cooled materials 21 are introduced into a tower top reflux tank 22 of the fractionating tower 18, one part of the gas-phase materials are introduced into the top of the fractionating tower 18 as reflux materials 23, the other part of the gas-phase materials are obtained as light hydrocarbon fraction byproducts 24, tower bottom liquid-phase materials are alkylated products 26, and heat required by the separation of the fractionating tower 18 is provided by a.

Figure 2 illustrates another embodiment of the present invention.

In fig. 2, the difference from fig. 1 is that a portion of the reactor outlet feed 28 is recycled back to the reactor inlet via a recycle pump 29 and another portion of the reacted feed 3 is fed to the fractionation column 4 for product separation.

The method can improve the performance of the solid acid alkylation catalyst and the quality of the alkylated gasoline, reduce the energy consumption of an alkylation unit and reduce the operation cost of the alkylation unit.

In the examples, the physicochemical parameter characterization method of the solid acid catalyst particles was as follows:

the macropore volume and total pore volume were determined by mercury intrusion based on the Washbum equation. D (-4 γ cos θ)/p where D is the pore diameter, p is the pressure applied during the measurement, γ is the surface tension, taking 485 dynes/cm, θ is the contact angle, taking 130 °.

Measurement of the average diameter of the catalyst particles: and measuring the longest side distance of the cross section of the particle by using a vernier caliper to obtain the average diameter of the particle.

Measurement of specific surface area: the specific surface area of the catalyst is measured by adopting a nitrogen low-temperature adsorption method, and the specific surface area is calculated by using a BET formula.

Measurement of particle length: randomly selecting 1g of catalyst particles, measuring the length of each particle in the 1g of catalyst particles, and adding the lengths of the particles; the length of each particle was measured using a vernier caliper.

The alkylation reaction performance reaction evaluation analysis method is as follows:

weighing quartz sand (20-40 meshes) and filling the quartz sand into a non-constant temperature section at the lower end of a tubular reactor, compacting, then filling into a three-layer nickel screen, filling and compacting 100g of catalyst, filling into the three-layer nickel screen, filling the quartz sand with 20-40 meshes into the non-constant temperature section at the upper layer of the reactor, and compacting. Finally, proper quartz cotton and nickel net are filled in sequence.

The reactor is connected into a pipeline, after the airtightness and the smoothness of the pipeline are detected, air in the nitrogen replacing device is replaced for more than three times, and then hydrogen is used for replacing for three times. Setting the hydrogen flow rate to be 300mL/min, the back pressure to be 3.0MPa, opening a heating source, setting the heating speed to be 1 ℃/min, heating to 200 ℃ and keeping for 1 h; then the temperature is raised to 450 ℃ at 1 ℃/min and kept for 3 h. After the pretreatment, the catalyst was cooled to the reaction temperature in the examples, the hydrogen in the nitrogen displacement device was displaced three times or more, and after the displacement, the catalyst was fed at a certain feed rate and reacted under the reaction conditions described in the examples.

The product is distributed through the Al-containing₂O₃And Agilent 7890A gas chromatography using PONA column and high pressure sampler. Sampling after a back pressure valve and before the exhaust gas is exhausted, sampling once every two hours, and dividing the sample into two parts at a sample inlet, namely a low-boiling-point mixture (C) for 0.01-0.1 min₄The following hydrocarbons) into Al₂O₃Column, high boiling point material (C) for 0.2 to 9.5 minutes₅The above hydrocarbons) is blown into the PONA column by a carrier gas. The obtained spectrogram is identified and the percentage content of each component is calculated by gasoline analysis software developed by petrochemical engineering scientific research institute.

Starting materials used in examples or comparative examples:

1. y-type molecular sieve (China petrochemical catalyst Co., Ltd.) with specific surface area of 680m²Pore volume 0.36mL/g, unit cell constant 2.457nm, m (SiO)₂/Al₂O₃) No. 9, No. Ya.

2. Several nano-alumina sols (china petrochemical catalyst division):

code Al 1: the alumina concentration was 5% and the average particle size was 20 nm.

Code Al 2: the alumina concentration was 15% and the average particle size was 150 nm.

Code Al 3: the alumina concentration was 20% and the average particle size was 300 nm.

3、Al₂O₃Binder powder: specific surface area 280m²G, pore volume 0.98 mL/g.

Example 1

This example illustrates the alkylation catalyst employed in the present invention.

Adding water into the Y-shaped molecular sieve numbered Ya and pulping to obtain the product with the solid content of 200kg/m³Adding alumina sol numbered as Al1 into the molecular sieve slurry according to the weight ratio of Ya to Al1 of 60:40 on a dry basis, stirring for 4 hours, uniformly mixing, adding 3 weight percent (based on the weight of the molecular sieve and the alumina sol calcined at 600 ℃) of nitric acid and sesbania powder into the dried mixed powder, adding water to ensure that the water-powder ratio is 0.8, uniformly kneading, extruding, drying and calcining the obtained wet strip to obtain a molded solid acid catalyst sample named as 60A1, wherein the properties are shown in Table 1.

The solid acid catalyst sample 60A1 of example 1 was separately charged with hydrogenation metal-containing Pt (H) under vacuum₂PtCl₆·6H₂O is a precursor) and impregnation liquid with a liquid-solid ratio of 2:1, after the addition is finished, the impregnation is finished under normal pressure for no more than 10 hours, the vacuum pumping is carried out at the temperature of no more than 80 ℃, the moisture in the catalyst is evaporated until the weight of the catalyst is 1.2-1.5 times of that of a solid acid catalyst precursor, and the catalyst is taken out after the evaporation, dried and roasted.

The resulting alkylation catalyst was numbered C1 and had a Pt content of 0.25 wt%.

Example 2

Adding water into the Y-shaped molecular sieve numbered Ya and pulping to obtain the product with the solid content of 200kg/m³Adding alumina sol numbered as Al2 into the molecular sieve slurry according to the weight ratio of Ya to Al2 of 80:20 on a dry basis, stirring for 4 hours, uniformly mixing, adding 3 weight percent (based on the weight of the molecular sieve and the alumina sol calcined at 600 ℃) of nitric acid and sesbania powder into the dried mixed powder, adding water to ensure that the water-powder ratio is 0.8, uniformly kneading, extruding, drying and calcining the obtained wet strip to obtain a molded solid acid catalyst sample named as 80A2, wherein the properties are shown in Table 1.

The solid acid catalyst sample 80A2 of example 2 was separately charged with hydrogenation metal-containing Pt (H) under vacuum₂PtCl₆·6H₂O is a precursor) and impregnation liquid with a liquid-solid ratio of 2:1, after the addition is finished, the impregnation is finished under normal pressure for no more than 10 hours, the vacuum pumping is carried out at the temperature of no more than 80 ℃, the moisture in the catalyst is evaporated until the weight of the catalyst is 1.2-1.5 times of that of a solid acid catalyst precursor, and the catalyst is taken out after the evaporation, dried and roasted.

The resulting alkylation catalyst was numbered C2 and had a Pt content of 0.25 wt%.

Example 3

Adding water into the Y-shaped molecular sieve numbered Ya and pulping to obtain the product with the solid content of 200kg/m³Adding alumina sol numbered Al3 into the molecular sieve slurry according to the weight ratio of Ya to Al3 of 95:5 in dry basis, stirring for 4 hours, uniformly mixing, adding 3 weight percent (based on the weight of the molecular sieve and the alumina sol calcined at 600 ℃) of nitric acid and sesbania powder into the dried mixed powder, adding water to ensure that the water-powder ratio is 0.8, uniformly kneading, extruding, drying and calcining the obtained wet strip to obtain a molded solid acid catalyst sample named 95A3, wherein the properties are shown in Table 1.

The solid acid catalyst samples 95A3 from example 3 were each charged with hydrogenation metal-containing Pt (H) under vacuum₂PtCl₆·6H₂O is a precursor), liquid and solidAnd (3) adding the impregnation liquid in a ratio of 2:1, then impregnating at normal pressure for no more than 10 hours, finishing impregnation, vacuumizing at the temperature of no more than 80 ℃, evaporating water in the catalyst until the weight of the catalyst is 1.2-1.5 times of that of a solid acid catalyst matrix, taking out the catalyst after evaporation, drying and roasting.

The resulting alkylation catalyst was numbered C3 and had a Pt content of 0.25 wt%.

Comparative examples 1 to 3

Comparative examples 1-3 illustrate the use of Y-type molecular sieves and Al₂O₃Binder powder solid phase mixing molding procedure and comparative solid acid catalyst samples obtained.

Mixing the Y-type molecular sieve with Al₂O₃Mixing the binder powders at a weight ratio of 60:40, 80:20 and 95:5, and adding 3 wt% (based on molecular sieve and Al)₂O₃The dry basis weight of the binder powder after being roasted at 600 ℃) is determined), water and powder ratio of the final mixed powder is ensured to be 0.8 by adding water, the mixture is extruded after being uniformly kneaded, and the obtained wet strip is dried and roasted to obtain a formed comparative solid acid catalyst sample.

Comparative solid acid catalyst samples were designated 60A, 80A and 95A, respectively, with properties as shown in table 1.

The comparative solid acid catalysts 60A, 80A and 95A were loaded with hydrogenation metal Pt to give (Pt content of 0.25 wt%) alkylation catalysts, which were numbered DB1, DB2 and DB3, respectively.

TABLE 1

The solid acid catalyst sample numbered 80A2 and the comparative solid acid catalyst sample numbered 80A were characterized using SEM and energy spectral surface scanning, and the morphology and element distribution results are shown in fig. 3. As can be seen from FIG. 3, the Si/Al distribution of the solid acid catalyst sample numbered 80A2 is more uniform, which indicates that the Y-type molecular sieve and Al are mixed in the liquid phase₂O₃The particle size distribution is uniform, and the acid site dispersibility is good.

Examples 4 to 5

Examples 4-5 illustrate the alkylation catalyst employed in the present invention.

The alkylation catalyst samples numbered C4-C5, having a Pt content of 0.1 wt% and 0.7 wt%, respectively, were obtained by supporting a hydrogenation metal on the basis of the solid acid catalyst sample 80a2 of example 2.

Example 6

The alkylation catalyst, code C6, was obtained by supporting the hydrogenation metal on the solid acid catalyst 80a2 of example 2, except that the hydrogenation metal was Pd (palladium nitrate was the precursor) and the Pd content was 0.5 wt%.

Example 7

This example illustrates the process of the present invention, the flow chart being shown in FIG. 1.

The main components of the carbon four raw material used for the alkylation reaction comprise isobutane by weight: 47.49%, n-butane: 14.62%, butene: 37.56 percent, and the balance of impurities.

The reactor is a fixed bed reactor, the number n of the reactors is 2, and the number m of the sections is 5. The loaded solid acid alkylation catalyst was alkylation catalyst sample C2.

Alkylation reaction conditions: the reaction temperature is 70 ℃, the pressure is 3.0MPa, the mole ratio of isoparaffin to olefin in the total feeding is 25, and the mass space velocity of olefin feeding is 0.15h^-1The catalyst was alkylation catalyst C2 prepared in example 2. Wherein cycle life is defined as the catalyst single pass run time at less than 99% butene conversion; and (3) adopting gas chromatography to analyze and detect the content of olefin at the outlet of the reactor, and carrying out octane number determination on the collected alkylated product, namely alkylated gasoline.

The reacted materials are separated by two fractionating towers to obtain the alkylated product, and the fractionation process is shown in figure 1. The operation conditions of the fractionating tower 4 are that the tower top temperature is 53 ℃, the tower bottom temperature is 130 ℃, and the tower top operation pressure is 0.7 MPa; the operation conditions of the fractionating tower 18 are that the tower top temperature is 53 ℃, the tower bottom temperature is 158 ℃, and the tower top operation pressure is 0.5 MPa. The compression ratio of the compressor is 2.3:1, and the outlet pressure of the compressor is 1.7 MPa.

The alkylation reaction results are shown in table 2.

The material balance data for the alkylation reaction product separation process are shown in table 3.

The energy consumption for the alkylation reaction product separation process using this example is shown in Table 4.

Example 8

The difference from example 7 is that the scheme is shown in FIG. 2.

The alkylation reaction results are shown in table 2.

Example 9

The difference from example 7 is that catalyst C1 from example 1 was used.

The alkylation reaction results are shown in table 2.

Example 10

The difference from example 7 is that catalyst C3 from example 3 was used.

The alkylation reaction results are shown in table 2.

Example 11

The difference from example 7 is that catalyst C4 from example 4 was used.

The alkylation reaction results are shown in table 2.

Example 12

The difference from example 7 is that catalyst C5 from example 5 was used.

The alkylation reaction results are shown in table 2.

Example 13

The difference from example 7 is that catalyst C6 from example 6 was used.

The alkylation reaction results are shown in table 2.

Comparative examples 4 to 6

Comparative examples 4-6 illustrate the use of comparative alkylation catalysts.

The feed, pretreatment and alkylation reaction conditions and product isolation procedures of example 7 were followed except that the alkylation catalysts used were the comparative alkylation catalysts DB1, DB2, DB3 prepared in comparative examples 1, 2, 3.

The alkylation reaction results are shown in table 2.

TABLE 2

Comparative example 7

This comparative example differs from example 7 only in the product separation step. Namely: and the gas phase material flow at the top of the fractionating tower enters a reflux tank after being condensed and cooled, the heat of the phase change process of the gas phase material flow at the top of the fractionating tower is not recycled, and the heat required by the separation process is completely provided by a reboiler at the bottom of the fractionating tower. The alkylation reaction product is derived from the alkylation reaction unit of the solid acid alkylation technology.

The material balance data for the alkylation reaction product separation process of this comparative example is shown in table 3.

The energy consumption for the alkylation reaction product separation process using this comparative example is shown in table 4.

Comparative example 8

This comparative example differs from example 8 only in the product separation step. Namely: the gas phase material flow at the top of the fractionating tower enters a reflux tank after being condensed and cooled, and the heat required by the separation process is provided by a reboiler at the bottom of the tower. The alkylation reaction product is derived from the alkylation reaction unit of the solid acid alkylation technology.

TABLE 3

	Examples 7 and 8	Comparative examples 7 and 8
			Alkylation reaction product, kg/h	288.5	288.5
Isobutane was circulated in kg/h	263.7	263.7
			Isobutane fraction, kg/h	2.0	2.0
N-butane fraction, kg/h	3.3	3.3
			Alkylated gasoline, kg/h	19.5	19.5

TABLE 4

	Example 7	Comparative example 7	Example 8	Comparative example 8
					Conversion of electricity consumption to energy consumption, MJ/t alkane oil	2239.6	568.1	2478.2	702.2
Conversion of steam usage to energy consumption, MJ/t alkane oil	4460.1	7674.7	1362.2	4798.7
					Conversion of circulating water consumption to energy consumption, MJ/t alkane oil	525.7	987.3	149.6	353.3
Total energy consumption, MJ/t alkane oil	7225.4	9230.1	3990.1	5854.2

As can be seen from Table 4: the combined energy consumption of example 7 is less than about 21.7% for comparative example 7 and the combined energy consumption of example 8 is less than about 31.8% for comparative example 8, demonstrating that the energy consumption level of the alkylation unit can be substantially reduced using the process of the present invention including the product separation step.

Claims

1. A solid acid alkylation method comprises an alkylation reaction unit and an alkylation reaction product separation unit, and is characterized in that,

(1) the specific volume of the macropores is 0.30-0.40 ml/g;

(3) the ratio of the specific surface area to the length of the particles is 3.40 to 4.50m²/mm；

The macropores refer to pores with the diameter of more than 50 nm;

2. The process of claim 1, wherein the ratio of the specific volume of macropores to the specific length of particles is 1.1 to 1.8 ml/(g-mm).

3. The process of claim 1 wherein said solid acid catalyst has a macropore specific volume of at least 0.35 ml/g.

4. The process according to claim 1, wherein the solid acid catalyst has a particle specific length of 0.15 to 0.4mm, preferably 0.18 to 0.36mm, more preferably 0.20 to 0.32 mm.

5. The process according to claim 1, wherein the solid acid catalyst has a total pore volume of at least 0.40ml/g, preferably at least 0.45 ml/g.

6. According to the claimsThe process of claim 1, wherein the solid acid catalyst has a specific surface area of not less than 500m²/g。

7. The process of claim 1 wherein the solid acid catalyst and the matrix material is alumina.

8. The method according to claim 7, wherein the precursor of the alumina is an alumina sol having a particle size of 20 to 400 nm.

9. The process according to claim 8, wherein the alumina sol is present in an amount of 2 to 98% by weight of the catalyst, preferably 10 to 70% by weight of the catalyst, based on the alumina.

10. The process of claim 1 wherein said solid acid catalyst has a molecular sieve as its solid acid component.

11. The process of claim 10 wherein said molecular sieve is selected from one or more of Y-type molecular sieve, beta, MCM-22 and MOR.

12. The method of claim 11, wherein the Y-type molecular sieve has a unit cell size of 2.430 to 2.470nm, preferably 2.440 to 2.460 nm.

13. The process of claim 1 wherein in said alkylation reaction unit, C₄～C₆With C₃～C₅The single-bond olefin is subjected to contact reaction in a fixed bed reactor.

14. The process according to claim 13, characterized in that the molar ratio of isoparaffin to olefin in the total feed is in the range of 5 to 50:1, preferably 7 to 30: 1.

15. the process according to claim 1, wherein the alkylation reaction temperature is 5 to 200 ℃, preferably 30 to 100 ℃, and the alkylation reaction pressure is 0.5 to 6.0MPa, preferably 1.0 to 3.0 MPa.

16. The method of claim 1, wherein the alkylation reaction unit has an olefin feed mass space velocity of 0.02-2.0 h^-1Preferably 0.05 to 0.5 hour^-1。

17. The process according to claim 1, wherein the alkylation reaction product separation unit is provided with a middle liquid phase fraction outlet of the first fractionating tower at a position 20-98%, preferably 40-80% of the first fractionating tower from top to bottom.

18. The method according to claim 1, wherein the alkylation reaction product separation unit is provided with a liquid phase fraction outlet in the middle of the fractionating tower arranged in the reducing section of the fractionating tower, the height ratio of the upper expanding section to the lower expanding section is 0.25: 1-49: 1, preferably 0.66: 1-4: 1, and the diameter ratio is 1: 1-6: 1, preferably 2: 1-4: 1.

19. The process of claim 1 wherein the alkylation reaction product outlet of said solid acid alkylation reaction unit is in communication with the alkylation reaction product inlet of said alkylation reaction product separation unit.

20. The process of claim 1 wherein the alkylation reaction product separation unit comprises a step of introducing the alkylation reaction product into a fractionation column for fractionation, wherein a gas phase stream at the top of the fractionation column is pressurized by a gas compressor, is used as a heat source for a middle reboiler of the fractionation column, is subjected to heat exchange and condensation, is introduced into an overhead reflux tank, and is partially or completely returned to the fractionation column as overhead reflux, and a liquid phase stream at the bottom of the fractionation column is used as an alkylation product.

21. The process of claim 1, the alkylation reaction product separation unit being provided with two fractionation columns, characterized by comprising the steps of:

(1) introducing an alkylation reaction product from the alkylation reaction unit into a fractionating tower of the product separation unit, pressurizing a gas phase material flow led out from the top of the fractionating tower by a gas compressor, taking the gas phase material flow as a heat source of a middle-section reboiler of the fractionating tower, returning one part of the tower top material flow subjected to heat exchange and condensation to the top of the fractionating tower as reflux of the fractionating tower, and obtaining isobutane fraction from the other part of the tower top material flow;

(2) introducing the liquid phase material flow at the bottom of the fractionating tower into a n-butane removing tower, condensing and cooling the gas phase material flow led out from the top of the n-butane removing tower, returning one part of the gas phase material flow to the top of the n-butane removing tower as reflux of the n-butane removing tower, obtaining n-butane fraction from the other part of the gas phase material flow, and taking the liquid phase material flow at the bottom of the n-butane removing tower as.

22. The method according to claim 1, characterized in that a pressurizing device is arranged at the top of the fractionating tower to pressurize the isobutane fraction, the compression ratio of the pressurizing device is 1.3-4.5: 1, and the absolute pressure of an outlet is 1.0-3.2 MPa.

23. The method according to claim 1, characterized in that the fractionating tower is provided with a middle section reboiler, the middle section material flow of the fractionating tower is heated by the isobutane fraction pressurized and heated at the top of the fractionating tower in the middle section reboiler, and the bottom of the fractionating tower is also provided with a bottom reboiler for supplementing the residual heat required for separating isobutane.