CN117567413A

CN117567413A - Production method of chlorophthalic anhydride

Info

Publication number: CN117567413A
Application number: CN202311490680.5A
Authority: CN
Inventors: 邓兆敬; 杨得岭; 钟吉彬; 武岳
Original assignee: China Chemical Technology Research Institute
Current assignee: China Chemical Technology Research Institute
Priority date: 2023-11-09
Filing date: 2023-11-09
Publication date: 2024-02-20

Abstract

The invention discloses a method for producing chlorophthalic anhydride, which relates to the field of chemical industry and aims to solve the problem that repeated rising, falling and rising and falling are required during the production of the chlorophthalic anhydride. The production method of the chlorophthalic anhydride comprises the following steps: under the action of a catalyst, using chloro-o-xylene and acetic acid as raw materials and using air as an oxidant, and performing oxidation reaction in an oxidation reactor to obtain a gas-phase product and an oxidation reaction liquid, wherein the oxidation reaction liquid at least contains chlorophthalic acid, and the gas-phase product at least contains acetic acid; and (3) carrying out acetic acid separation on the gas-phase product to obtain acetic acid and a purified reaction solution, then dehydrating the purified reaction solution to obtain an anhydride-forming crude product, and then respectively separating at least two chlorophthalic anhydride monomers from the anhydride-forming crude product. The method for producing the chlorophthalic anhydride is used for producing the chlorophthalic anhydride monomer.

Description

Production method of chlorophthalic anhydride

Technical Field

The invention relates to the field of chemical industry, in particular to a production method of chlorophthalic anhydride.

Background

The chlorophthalic anhydride is an important organic chemical raw material, is mainly used for synthesizing polyimide engineering plastics with excellent performance, is also used as a production intermediate of dyes, medicines and pesticides, and is widely applied to the high-tech fields of aviation, aerospace, electronics and the like.

The method for synthesizing the chlorophthalic anhydride is more, wherein o-xylene is used as a raw material, the mono-chloro-o-xylene is obtained after chlorination, and the method for catalyzing and oxidizing the o-xylene into the chlorophthalic anhydride by introducing air under the liquid phase condition is the most economical and most industrial development prospect. The method generally adopts an intermittent oxidation process or an intermittent distillation process, and needs repeated operation of lifting and lowering pressure and lifting temperature of the oxidation tower, so that the equipment is easy to be fatigued and damaged, and the service life of the equipment is shortened. Meanwhile, a great amount of partial oxidation reaction products and side reaction products are generated in incomplete oxidation reaction, and are difficult to separate by a distillation method, and the reflux ratios required for removing different fractions by distillation are greatly different, so that the operation difficulty is high and the energy consumption is high if the same batch distillation device is adopted.

In addition, the presence of halogen as an electron withdrawing substituent on the chloroxylenes benzene ring makes the oxidation of chloroxylenes more difficult under similar conditions than the oxidation of xylenes (e.g., para-xylene), less selective oxidation, and more partially oxidized products and byproducts, which are not cleanly separated from the anhydride even with conventional distillation or recrystallization.

Disclosure of Invention

The invention aims to provide a production method applied to a system for producing chlorophthalic anhydride with full continuous flow and high purity.

In order to achieve the above object, the present invention provides a method for producing chlorophthalic anhydride, comprising:

under the action of a catalyst, performing oxidation reaction in an oxidation reactor by taking chloro-o-xylene and acetic acid as raw materials and taking air as an oxidant to obtain a gas-phase product and an oxidation reaction liquid, wherein the oxidation reaction liquid at least contains chlorophthalic acid, and the gas-phase product at least contains acetic acid;

and (3) separating the gas phase product by acetic acid to obtain acetic acid and a purified reaction solution, dehydrating the purified reaction solution to obtain an anhydride-forming crude product, and separating at least two chlorophthalic anhydride monomers from the anhydride-forming crude product.

Compared with the prior art, in the method for producing the chlorophthalic anhydride, the chloro-o-xylene and the acetic acid are used as raw materials and air is used as an oxidant under the action of the catalyst, and the oxidation reaction is carried out in the oxidation reactor to obtain a gas-phase product and an oxidation reaction liquid, wherein the oxidation reaction liquid at least contains chlorophthalic acid, and the gas-phase product at least contains acetic acid, so that a part of acetic acid and the product can be separated, thereby improving the reaction rate and the reaction yield. Meanwhile, the gas-phase product can be separated by utilizing a post-treatment unit to obtain acetic acid and a purified reaction liquid, and the purified reaction liquid is dehydrated by utilizing an anhydride-forming reactor to obtain an anhydride-forming crude product. The anhydride-forming reaction liquid may be fed to a multistage separation unit, and the corresponding chlorophthalic anhydride monomer may be separated by the multistage separation unit, and the types of the chlorophthalic anhydride monomers to be separated by the respective separation units may be different. Therefore, the embodiment of the invention adopts a fully continuous production method of the chlorophthalic anhydride when the chlorophthalic anhydride is produced, and repeated operation of increasing and decreasing pressure and increasing and decreasing temperature is not needed for the oxidation reactor, so that the production is not easy to damage, and the service life of equipment is prolonged.

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 invention and do not constitute a limitation on the invention. In the drawings:

FIG. 1 is a basic block diagram showing a production system of chlorophthalic anhydride in this example;

FIG. 2 is a flow chart showing the production method of chlorophthalic anhydride in this example;

fig. 3 shows a basic block diagram of a post-processing unit in the present embodiment;

fig. 4 shows a flow chart of the production method of the post-treatment stage in the present embodiment;

fig. 5 shows a basic block diagram of an exhaust gas treatment unit in the present embodiment;

FIG. 6 shows a flow chart of the production method of the tail gas treatment stage in the present embodiment;

fig. 7 shows a basic block diagram of the tail gas condensing unit in the present embodiment;

FIG. 8 is a basic block diagram showing an acetic acid separation unit in the present embodiment;

FIG. 9 shows a flow chart of the production method of the acetic acid separation stage in this example;

FIG. 10 shows a basic block diagram of the anhydride-forming apparatus in this example;

fig. 11 shows a basic block diagram of the multistage separation unit of the present embodiment;

FIG. 12 shows a flow chart of the production process of the product separation stage in this example;

FIG. 13 is a schematic view showing the structure of an oxidation reactor of the present embodiment;

FIG. 14 is a schematic view showing the structure of the anhydride-forming apparatus of the present embodiment;

FIG. 15 is a process flow chart showing the production system of chlorophthalic anhydride in this example.

Reference numerals:

a 100-oxidation reactor; 200-a post-processing unit; 210-an exhaust gas treatment unit; 211-a tail gas condensing unit; 2111-a first condenser; 2112-a second condenser; 2113-a third condenser; 212-an exhaust gas absorption unit; 220-acetic acid separation unit; 221-a flash tank; 222-stripper; 230-an acetic acid recovery unit; 300-anhydride forming device; 310-anhydride formation reactor; 311-an oxidation reaction liquid distributor; 312-a first stirrer; 313-a second stirrer; 320-a catalyst separation unit; 321-an evaporator; 322-recovery vessel; 330-crude anhydride condenser; 340-a crude anhydride sump; 350-an anhydride-forming condenser; 400-multistage separation unit; 410-a first separation unit; 420-a second separation unit; 430-a third separation unit; a-a first feed tube; b-a second feed tube.

Detailed Description

In order to make the technical problems, technical schemes and beneficial effects to be solved more clear, the invention is further described in detail below with reference to the accompanying drawings and embodiments. It should be understood that the specific embodiments described herein are for purposes of illustration only and are not intended to limit the scope of the invention.

Furthermore, the terms "first," "second," and the like, are used for descriptive purposes only and are not to be construed as indicating or implying a relative importance or implicitly indicating the number of technical features indicated. Thus, a feature defining "a first" or "a second" may explicitly or implicitly include one or more such feature. In the description of the present invention, the meaning of "a plurality" is two or more, unless explicitly defined otherwise. The meaning of "a number" is one or more than one unless specifically defined otherwise.

The chlorophthalic anhydride is polyimide engineering plastic with excellent performance, and can be prepared into chlorophthalic acid by using chloro-o-xylene and air under the action of a catalyst, and then dehydrating to prepare the chlorophthalic anhydride. The preparation process of the chlorophthalic anhydride adopts an intermittent oxidation process or an intermittent distillation process, repeated operation of lifting and lowering pressure and lifting temperature is required for an oxidation tower, equipment is easy to be fatigued and damaged, and the service life of the equipment is shortened. Meanwhile, a great amount of partial oxidation reaction products and side reaction products are generated in incomplete oxidation reaction, and are difficult to separate by a distillation method, and the reflux ratios required for removing different fractions by distillation are greatly different, so that the operation difficulty is high and the energy consumption is high if the same batch distillation device is adopted.

The traditional process adopts multistage serial reaction kettles with stirring as oxidation reactors, each stage of reaction kettle needs to be provided with an independent cooling system, and the multistage stirring anhydride-forming reaction kettles are adopted as anhydride-forming reactors, so that the process is long and unreasonable, the equipment quantity is large, and the investment is large. Meanwhile, the measure of generating water by oxidation reaction is not considered, the oxidation reaction is weakened when the water content is high, a large amount of byproducts and intermediate products are generated, the system is difficult to control due to slight fluctuation, the process technology is unstable, and the energy consumption for removing acetic acid and water from the oxidized products and separating the acetic acid from the water is huge. Moreover, energy recovery has not been considered yet, and the heat of release during the oxidation of chloro-ortho-xylene is large (the heat of reaction released after oxidation of 3-chloro-ortho-xylene and 4-chloro-ortho-xylene is 1170.32kJ/mol, 1167.3kJ/mol, respectively). Meanwhile, the final product of the method is a mixture of 3-chlorophthalic anhydride and 4-chlorophthalic anhydride, and the commercially required high-purity chlorophthalic anhydride monomer is not obtained.

Meanwhile, since halogen as an electron withdrawing substituent exists on the chloroxylenes benzene ring, the oxidation reaction of chloroxylenes under similar conditions is more difficult than that of xylenes (such as paraxylene), the selectivity of the oxidation reaction is lower, and partial oxidation and byproducts including four kinds of chloro2-benzo [ C ] furanones, four kinds of chlorobenzoic acids and phthalic acid are more abundant, and these partial oxidation and byproduct impurities cannot be separated cleanly from the acid anhydride even by conventional distillation or recrystallization.

In view of the above problems, the embodiments of the present invention provide a method for producing chlorophthalic anhydride, which can provide a method for producing chlorophthalic anhydride that is fully continuous, has high product purity, few intermediates and byproducts, high acetic acid recovery rate, and little loss to equipment, and is applied to a production system of chlorophthalic anhydride that includes an oxidation reactor, a post-treatment unit, an anhydride-forming reactor, and a multistage separation unit.

FIG. 1 shows a basic block diagram of a system for producing chlorophthalic anhydride in an embodiment of the present invention. As shown in fig. 1, the system for producing chlorophthalic anhydride according to the embodiment of the present invention comprises: the oxidation reactor 100, the post-treatment unit 200, the anhydride-forming device 300 and the multi-stage separation unit 400 are sequentially communicated with each other, and the oxidation reactor 100, the post-treatment unit 200, the anhydride-forming device 300 and the multi-stage separation unit 400.

As shown in fig. 1, the oxidation reactor 100 is in communication with a first feed pipe a for transporting chloro-ortho-xylene, acetic acid, and catalyst, and a second feed pipe b for transporting air, respectively. In this case, the first feed pipe a may be used to feed o-xylene, acetic acid, and a catalyst to the oxidation reactor 100, and the second feed pipe b may be used to feed air to the oxidation reactor 100, and at this time, the o-xylene and air undergo an oxidation reaction in the oxidation reactor 100 under the action of the catalyst to obtain an oxidation reaction liquid containing chlorophthalic acid and a gas-phase product.

As shown in fig. 1, the oxidation reactor 100 is connected to the post-treatment unit 200, the post-treatment unit 200 is connected to the anhydride-forming device 300, the post-treatment unit 200 is used to post-treat the oxidation reaction liquid and then convey the post-treated reaction liquid to the anhydride-forming device 300, in the anhydride-forming device 300, the chlorophthalic acid can be dehydrated and converted into at least two types of chlorophthalic anhydride monomers, the oxidation reaction liquid outlet of the anhydride-forming device 300 is connected to the inlet of the multistage separation unit 400, the multistage separation unit 400 comprises a plurality of separation units connected in series, each separation unit can separate a corresponding type of chlorophthalic anhydride monomer, and the types of the chlorophthalic anhydride monomers used for separation by each separation unit are different.

FIG. 2 shows a flow chart of a method for producing chlorophthalic anhydride in an embodiment of the present invention. As shown in FIG. 2, the method for producing the chlorophthalic anhydride provided by the invention comprises the following steps:

step 201: under the action of a catalyst, chloro-o-xylene and acetic acid are used as raw materials, air is used as an oxidant, and oxidation reaction is carried out in an oxidation reactor to obtain a gas-phase product and oxidation reaction liquid, wherein the oxidation reaction liquid at least contains chlorophthalic acid, and the gas-phase product at least contains acetic acid.

In specific implementation, the first feeding pipe a and the second feeding pipe b can convey chloro-o-xylene, acetic acid and a catalyst to the oxidation reactor 100 according to actual reaction molar ratio, so that the chloro-o-xylene and oxygen perform oxidation reaction in the oxidation reactor 100 at a certain temperature, the chloro-o-xylene is almost completely converted into chlorophthalic acid, oxidation reaction liquid and gas phase products are obtained, the oxidation reaction liquid at least contains chlorophthalic acid, and the gas phase products at least contain acetic acid.

For example: the air can be compressed air with the oxygen content of 21v%, the catalyst can be at least one of cobalt acetate, manganese acetate and tetrabromoethane, and the ion mass ratio of chloro-o-xylene, acetic acid, cobalt ions, manganese ions and bromide ions can be 1: (3-6): (0.002-0.006): (0.001-0.004): (0.0005-0.004), and the mass ratio of air (the mole content of oxygen is 21%) to chloro-o-xylene is (3.6-5): 1, the reaction temperature is 185 ℃ to 230 ℃, and the reaction pressure is 1.2MPa (G) to 2.3MPa (G). The oxidation reaction liquid can also contain chloro-o-xylene, acetic acid, a catalyst and an intermediate product, the mass content of water in the oxidation reaction liquid is 5-10wt%, the gas phase product can at least contain gaseous acetic acid and water vapor, and can also contain waste gases such as oxygen, nitrogen, carbon monoxide, carbon dioxide and the like, and the volume content of oxygen in the gas phase product is 3-7%. The acetic acid used in the present invention is pure acetic acid, preferably acetic acid having a purity of more than 98wt%, for example, acetic acid having a purity of 98.5wt% and 99 wt%. Acetic acid on the one hand acts as a reactant and on the other hand also as a solvent for the reaction.

Step 202: and (3) carrying out acetic acid separation on the gas-phase product to obtain acetic acid and a purified reaction solution, then dehydrating the purified reaction solution to obtain an anhydride-forming crude product, and then respectively separating at least two chlorophthalic anhydride monomers from the anhydride-forming crude product.

In the specific implementation, the gas-phase product can enter a post-treatment unit to separate acetic acid and waste gas, the oxidation reaction liquid can enter the post-treatment unit to separate acetic acid, the separated oxidation reaction liquid enters an anhydride forming device, chlorophthalic acid contained in the oxidation reaction liquid is dehydrated and converted into at least two chlorophthalic anhydride monomers by the anhydride forming device, and then different chlorophthalic anhydride monomers are separated by a multi-stage separation unit. At this time, various types of chlorophthalic anhydride monomers, for example, can be obtained: 4-chlorophthalic anhydride and 3-chlorophthalic anhydride.

Therefore, in the method for producing the chlorophthalic anhydride, the oxidation reactor, the post-treatment unit, the anhydride forming device and the multistage separation unit in the production system are sequentially communicated. Under the action of a catalyst, chloro-o-xylene and acetic acid are used as raw materials, air is used as an oxidant, and oxidation reaction is carried out in an oxidation reactor to obtain a gas-phase product and oxidation reaction liquid, wherein the oxidation reaction liquid at least contains chlorophthalic acid, the gas-phase product at least contains acetic acid, and a part of acetic acid and the product can be separated, so that the reaction rate and the reaction yield are improved. Meanwhile, the oxidation reactor is communicated with the post-treatment unit, the post-treatment unit is communicated with the anhydride forming device, the post-treatment unit can be used for separating the gas phase product into acetic acid to obtain acetic acid and purified reaction liquid, and the anhydride forming reactor is used for dehydrating the purified reaction liquid to obtain a crude anhydride product. In addition, since the anhydride forming device is communicated with the multi-stage separation unit, the anhydride forming reaction liquid can be conveyed to the multi-stage separation unit, the multi-stage separation unit is used for separating corresponding types of chlorophthalic anhydride monomers, and the types of the chlorophthalic anhydride monomers separated by the separation units are different. Therefore, the embodiment of the invention adopts a fully continuous production method of the chlorophthalic anhydride when the chlorophthalic anhydride is produced, and repeated operation of increasing and decreasing pressure and increasing and decreasing temperature is not needed for the oxidation reactor, so that the production is not easy to damage, and the service life of equipment is prolonged.

Fig. 3 shows a basic block diagram of a post-processing unit according to an embodiment of the present invention, and as shown in fig. 3, the post-processing unit 200 includes: a tail gas treatment unit 210, an acetic acid separation unit 220, and an acetic acid recovery unit 230. The gas phase component discharge port of the oxidation reactor 100 is communicated with the inlet of the tail gas treatment unit 210, the oxidation reaction liquid outlet of the oxidation reactor 100 is communicated with the inlet of the acetic acid separation unit 220, and the oxidation reaction liquid outlet of the acetic acid separation unit 220 is communicated with the inlet of the anhydride-forming device 300. The oxidation reaction liquid outlet of the tail gas treatment unit 210 and the gas outlet of the acetic acid separation unit 220 are respectively communicated with the inlet of the acetic acid recovery unit 230, so that the acetic acid solution in the gas-phase product can be recovered to the acetic acid recovery unit 230, the acetic acid outlet of the acetic acid recovery unit 230 is communicated with the first feed pipe a of the oxidation reactor 100, and the oxidation reaction liquid outlet of the tail gas treatment unit 210 is also communicated with the reflux port of the oxidation reactor 100.

Fig. 4 shows a flow chart of the production method of the post-treatment stage in this embodiment. As shown in fig. 4, acetic acid separation is performed on the gas phase product to obtain acetic acid and a purified reaction liquid, which comprises:

step 401: condensing the gas-phase product to obtain acetic acid solution, and separating acetic acid from the acetic acid solution and sending the acetic acid solution into the oxidation reactor.

Step 402: acetic acid separation is carried out on the oxidation reaction liquid to obtain a purified reaction liquid and a gas phase component containing acetic acid, and then the acetic acid is separated from the gas phase component.

In a specific implementation, the oxidation reactor 100 may be a bubbling oxidation tower, the acetic acid recovery unit 230 may be an acetic acid recovery tower, the gas phase product discharged from the oxidation reactor 100 enters the tail gas treatment unit 210 to separate the waste gas from the acetic acid gas, the waste gas may be discharged into the waste gas treatment system, most of the acetic acid gas is refluxed into the oxidation reactor 100 through valve control after condensation, and the small part of the acetic acid gas is recovered into the acetic acid recovery unit 230 to control the mass content of water in the oxidation reactor 100 to be 5wt% -10 wt%, and the reflux ratio to be (20-40): 1. meanwhile, the oxidation reaction liquid of the oxidation reactor 100 enters the acetic acid separation unit 220 to separate acetic acid, the acetic acid in the oxidation reaction liquid is recovered to the acetic acid recovery unit 230, the acetic acid is treated in the acetic acid recovery unit 230 to obtain purified acetic acid, and then the purified acetic acid is recovered to the oxidation reactor 100 to continue the reaction. According to this, acetic acid volatilized during the reaction and acetic acid discharged from the reaction solution can be recovered and reused, and the amount of acetic acid used can be reduced, thereby reducing the production cost.

Fig. 5 is a basic block diagram of an exhaust gas treatment unit according to an embodiment of the present invention, and as shown in fig. 5, a method for producing chlorophthalic anhydride according to an embodiment of the present invention is provided, where an exhaust gas treatment unit 210 in a production system of chlorophthalic anhydride includes: the device comprises a tail gas condensing unit 211 and a tail gas absorbing unit 212, wherein a tail gas discharge port of the oxidation reactor 100 is communicated with an inlet of the tail gas condensing unit 211, a tail gas discharge port of the tail gas condensing unit 211 is communicated with an inlet of the tail gas absorbing unit 212, an oxidation reaction liquid outlet of the tail gas absorbing unit 212 is communicated with an inlet of the acetic acid recovery unit 230, an oxidation reaction liquid outlet of the tail gas condensing unit 211 is divided into two paths, one path is communicated with a reflux port of the oxidation reactor 100, and the other path is communicated with an inlet of the acetic acid recovery unit 230.

Fig. 6 shows a flow chart of the production method of the tail gas treatment stage in the present embodiment. As shown in fig. 6, the vapor phase product is subjected to condensation treatment to obtain an acetic acid solution, which comprises:

step 601: condensing the gas-phase product to obtain a first acetic acid solution and tail gas;

step 602: absorbing the tail gas and separating the residual second acetic acid solution in the tail gas by using a tail gas absorption unit;

Step 603: separating acetic acid from the first acetic acid solution and the second acetic acid solution.

In specific implementation, the gas phase component discharged from the oxidation reactor 100 enters the tail gas condensing unit 211 to perform condensation treatment, after part of acetic acid gas in the gas phase component is condensed into a first acetic acid solution, most of the condensed acetic acid solution and water vapor are refluxed into the oxidation reactor 100 through valve control, and a small part of the condensed acetic acid solution and water vapor are sent to the tail gas absorbing unit 212, and uncondensed gas in the gas phase component is discharged into the tail gas absorbing unit 212. The tail gas absorption unit 212 can be a high-pressure tail gas absorption tower, and is internally provided with a filler and desalted water, so that residual acetic acid gas in uncondensed gas can be converted into acetic acid solution, and the converted acetic acid solution is discharged to an acetic acid recovery unit for recovery, thereby greatly reducing acetic acid consumption and production cost.

For example, fig. 7 shows a basic block diagram of an exhaust gas condensing unit according to an embodiment of the present invention, and as shown in fig. 7, in the method for producing chlorophthalic anhydride according to the embodiment of the present invention, the exhaust gas condensing unit 211 includes a first condenser 2111, a second condenser 2112 and a third condenser 2113, and the first condenser 2111, the second condenser 2112 and the third condenser 2113 are sequentially connected.

The method for producing the chlorophthalic anhydride provided by the embodiment of the invention comprises the steps of condensing a gas-phase product to obtain an acetic acid solution, and separating the acetic acid from the acetic acid solution by an acetic acid recovery unit: dividing the acetic acid solution into a first path of acetic acid solution and a second path of acetic acid solution, refluxing the first path of acetic acid solution to the oxidation reactor, and sending the second path of acetic acid solution to the acetic acid recovery unit.

The outlets of the first condenser 2111, the second condenser 2112 and the third condenser 2113 are sequentially connected and divided into two paths, one path is connected to the reflux port of the oxidation reactor 100, and the other path is connected to the inlet of the acetic acid recovery unit 230. It should be understood that the number of the tail gas condensing units may be plural or one, and three are exemplified herein.

In practical applications, shell passes of the first condenser 2111 and the second condenser 2112 are connected to a boiler water supply pipe and a steam pipe network, shell passes of the third condenser 2113 are connected to desalted water, tube pass gas phase inlets and outlets of the first condenser 2111, the second condenser 2112 and the third condenser 2113 are sequentially connected, and tube pass gas phase inlets and outlets of the third condenser 2113 are also connected to the tail gas absorbing unit 212. The outlets of the oxidation reaction liquid of the first condenser 2111, the second condenser 2112 and the third condenser 2113 are connected with the gas phase component discharging pipe and the condensate return pipe of the oxidation reactor 100, the upper part of the tail gas absorbing unit 212 is connected with the desalted water inlet, the top outlet of the tail gas absorbing unit 212 is connected with the tail gas treatment device of the whole plant, and the bottom outlet of the tail gas absorbing unit 212 is connected with the acetic acid recovery unit 230.

In specific implementation, the gas phase component discharged from the top of the oxidation reactor 100 sequentially passes through the first condenser 2111, the second condenser 2112 and the third condenser 2113, and divides the outlet of the oxidation reaction liquid into two paths, one path is returned to the oxidation reactor 100, and the other path is sent to the acetic acid recovery unit 230. It should be understood that the oxidation reaction liquid in the oxidation reaction liquid outlet is acetic acid solution, and the mass content of water in the oxidation reactor 100 is 5wt% -10 wt% through valve control, and the reflux ratio is (20-40): 1. the first condenser 2111 of the embodiment of the invention produces steam of 0.4MPa (G) to 0.6MPa (G), the second condenser 2112 produces steam of 0.1MPa (G) to 0.3MPa (G), the third condenser 2113 cools the gas phase component by desalted water, and the temperatures of the gas phase components coming out from the tube passes of the first condenser 2111, the second condenser 2112 and the third condenser 2113 are 165 ℃ to 180 ℃, 130 ℃ to 155 ℃ and 50 ℃ to 110 ℃ in sequence. Based on the above, because the gasification points of different components are different, the components reaching the gasification point under a certain pressure can be gasified, different gas phase components can be gasified respectively through the first condenser, the second condenser and the third condenser, and the oxidized gas phase components coming out from the top of the oxidation reactor can be subjected to gradual heat recovery through the first condenser, the second condenser and the third condenser in sequence, and the third condenser adopts a method of preheating desalted water, so that the reaction heat is effectively and reasonably utilized in a cascade manner.

The gas phase component from the third condenser still contains more acetic acid, and if the part of acetic acid is not recovered, the waste of resources and the increase of the product cost are caused, so that the high-pressure tail gas absorption tower is arranged, the desalted water contained in the high-pressure tail gas absorption tower is used for absorbing the acetic acid in the gas phase component, and the obtained acetic acid aqueous solution is directly sent to a subsequent acetic acid recovery device for recovery, so that the consumption and the production cost of the acetic acid are greatly reduced.

In an alternative manner, fig. 8 shows a basic block diagram of an acetic acid separation unit according to an embodiment of the present invention, and as shown in fig. 8, in a method for producing chlorophthalic anhydride according to an embodiment of the present invention, an acetic acid separation unit 220 in a production system for chlorophthalic anhydride is used, including: a flash tank 221 and a stripper 222. The oxidation reaction liquid outlet of the oxidation reactor 100 is communicated with the inlet of the flash tank 221, the oxidation reaction liquid outlet of the flash tank 221 and the exhaust port of the anhydride forming device 300 are communicated with the inlet of the stripping tower 222, the gas outlet of the flash tank 221 and the gas outlet of the stripping tower 222 are communicated with the inlet of the acetic acid recovery unit 230, and the oxidation reaction liquid outlet of the stripping tower 222 is communicated with the inlet of the anhydride forming device 300.

Fig. 9 shows a flow chart of the production method of the acetic acid separation stage in this example. As shown in fig. 9, acetic acid separation is performed on the oxidation reaction liquid by an acetic acid separation unit to obtain a purified reaction liquid and a gas phase component containing acetic acid, comprising:

step 901: acetic acid separation is carried out on the oxidation reaction liquid, and a first purified reaction liquid and a first gas phase component containing acetic acid are obtained;

step 902: acetic acid separation is carried out on the first purification reaction liquid, so that a second purification reaction liquid and a second gas phase component containing acetic acid are obtained;

step 903: and sending the second purified reaction liquid into an anhydride-forming reactor for anhydride-forming reaction.

In specific implementation, the oxidation reaction liquid of the oxidation reactor 100 is sent to the flash tank 221 for normal pressure flash evaporation, the high temperature and pressure energy of the oxidation product are fully utilized, the acetic acid accounting for about 40% of the total amount of the acetic acid is vaporized, the temperature of the feed liquid after flash evaporation is 115-130 ℃, and the flash evaporated acetic acid gas directly enters the acetic acid recovery unit 230 for rectification, so that the steam consumption of the subsequent stripping tower 222 and the steam consumption of the acetic acid recovery unit 230 are saved. Meanwhile, the flashed oxidation reaction liquid material enters a stripping tower 222, the bottom of the stripping tower 222 is connected with a stripping tower reboiler, the stripping tower reboiler adopts an external forced circulation heating mode to heat and gasify most of the residual acetic acid and water, a built-in condenser is further arranged at a position, close to the top, in the stripping tower, of the stripping tower, the gasified acetic acid steam and the gasified water steam are partially condensed by the built-in condenser to obtain purified acetic acid steam and purified water steam, the purified acetic acid steam and the purified water steam enter an acetic acid recovery unit 230 to be rectified, and an oxidation reaction liquid inlet of an anhydride forming device 300 is connected with a pump outlet of the stripping tower 222, so that the oxidation reaction liquid in the stripping tower 222 enters the anhydride forming device 300. Wherein the operation pressure of the stripping tower is normal pressure, the temperature of the tower kettle is 160-200 ℃, and the reflux ratio is 0.08-0.3; acetic acid recovery tower rectifying separation acetic acid and water, operating pressure is normal pressure, reflux ratio is 2-6, concentration of acetic acid separated at tower bottom is 98-99.9 wt%, acetic acid mass content in waste water separated at tower top is 0.1-2 wt% and fed into waste water treatment device, non-condensable gas is fed into tail gas treatment device of factory and is discharged after reaching standard, acetic acid recovery tower can be plate tower or composite tower combining tower plate and filler, theoretical tower plate number is 27-45.

According to the embodiment of the invention, the acetic acid steam from the top of the stripping tower is not required to be totally condensed into the oxidation reaction liquid, a part of gas directly enters the acetic acid recovery tower for rectification, so that the steam consumption of a reboiler of the acetic acid recovery tower is saved, meanwhile, the top of the stripping tower is provided with the built-in condenser, a small amount of acetic acid steam can be condensed to wash and purify organic matters in the acetic acid steam, and the recovered acetic acid has high purity.

In one example, fig. 10 shows a basic block diagram of an anhydride-forming apparatus according to an embodiment of the present invention, and as shown in fig. 10, in the method for producing chlorophthalic anhydride according to the embodiment of the present invention, an anhydride-forming apparatus 300 in a chlorophthalic anhydride production system is applied, which includes an anhydride-forming reactor 310, a catalyst separation unit 320, and a crude anhydride condenser 330, and an oxidation reaction liquid outlet of the anhydride-forming reactor 310 is sequentially communicated with a multistage separation unit 400 through the catalyst separation unit 320 and the crude anhydride condenser 330. The catalyst separation unit 320 includes: the oxidation reaction liquid outlet of the anhydride-forming reactor 310 is communicated with the inlet of the evaporator 321, the gas outlet of the evaporator 321 is communicated with the multistage separation unit 400 through the crude anhydride condenser 330, and the gas outlet of the evaporator 321 is communicated with the recovery container 322. Wherein, the recovery vessel 322 may be a raffinate tank for storing separated catalyst and heavy component raffinate, and the evaporator 321 may be a wiped film evaporator.

The method for producing the chlorophthalic anhydride provided by the embodiment of the invention comprises the steps of dehydrating the purified reaction liquid by using an anhydride-forming reactor, and when obtaining an anhydride-forming crude product, further comprising: dehydrating the second purified reaction solution to obtain an anhydride-forming crude product containing at least acetic acid and chlorophthalic anhydride; and feeding acetic acid separated from the crude anhydride-forming product into a post-treatment unit. Wherein the second purifying reaction liquid is the reaction liquid after separating acetic acid in the stripping tower.

In specific implementation, the oxidation reaction liquid from the bottom of the stripping tower 222 enters an anhydride forming reactor 310, the oxidation reaction liquid in the anhydride forming reactor 310 is evaporated and dehydrated to convert chlorophthalic acid into chlorophthalic anhydride, the crude anhydride from the anhydride forming reactor 310 enters an evaporator 321 of a catalyst separation unit 320, and as the oxidation reaction liquid outlet of the anhydride forming reactor 310 is sequentially communicated with a multistage separation unit 400 through the catalyst separation unit 320 and a crude anhydride condenser 330, and the crude anhydride condenser 330 is communicated with the multistage separation unit 400 through a crude anhydride storage tank 340, so that the crude anhydride is continuously vaporized, the vaporized crude anhydride is liquefied through the crude anhydride condenser 330 and then enters the crude anhydride storage tank 340, so that the catalyst and the multi-ring heavy component raffinate are separated and stored in a recovery container 322, and then the crude anhydride is pumped into the multistage separation unit 400 for separating various phthalic anhydride monomers. It should be understood that the multi-stage separation unit 40 may be a plurality of chlorophthalic anhydride separation columns in series. Wherein the operating pressure of the evaporator 321 is 0kPa (A) to 10kPa (A), and the temperature is 170 ℃ to 250 ℃.

Fig. 11 is a basic block diagram of a multistage separation unit according to an embodiment of the present invention, and as shown in fig. 11, a method for producing chlorophthalic anhydride according to an embodiment of the present invention is provided, in which a multistage separation unit 400 in a production system of chlorophthalic anhydride is applied, comprising: the crude anhydride sump contained in the anhydride-forming apparatus 300 is in communication with the first, second and third separation units 410, 420 and 430, respectively.

Fig. 12 shows a flow chart of the production method of the product separation stage in this example. As shown in fig. 12, at least two types of chlorophthalic anhydride monomers are separated from the crude anhydride-forming product by a multi-stage separation unit, comprising:

step 1201: separating light components from the crude anhydride-forming product by using a first separation unit;

step 1202: separating 4-chlorophthalic anhydride from the crude anhydride-forming product by a second separation unit;

step 1203: and separating 3-chlorophthalic anhydride from the crude anhydride product by a third separation unit.

For example: the multi-stage separation unit 400 may be three chlorophthalic anhydride separation towers, namely a phthalic anhydride separation tower, a 4-chlorophthalic anhydride separation tower and a 3-chlorophthalic anhydride separation tower, wherein the theoretical plate number of the phthalic anhydride separation tower is 35-50, the operation pressure is 2kPa (A) to 7kPa (A), the tower top temperature is 178-185 ℃, and the reflux ratio is 220-300. The theoretical plate number of the 4-chlorophthalic anhydride separation tower is 30-45, the operating pressure is 0kPa (A) -10 kPa (A), the tower top temperature is 170-205 ℃, and the reflux ratio is 2-10. The theoretical plate number of the 3-chlorophthalic anhydride separation tower is 20-35, the operating pressure is 0kPa (A) to 10kPa (A), the tower top temperature is 200-230 ℃, and the reflux ratio is 1-10. The phthalic anhydride separation tower, the 4-chlorophthalic anhydride separation tower and the 3-chlorophthalic anhydride separation tower all comprise reboilers, and the reboilers can be falling film reboilers or reboilers of other forms, and are not limited herein.

In practice, the crude anhydride is continuously subjected to product separation in three separation columns. Firstly removing light components such as phthalic anhydride in a phthalic anhydride separation tower, then separating 4-chlorophthalic anhydride in a second tower, namely a 4-chlorophthalic anhydride separation tower, wherein the purity of the 4-chlorophthalic anhydride is more than 99.6wt%, and finally separating 3-chlorophthalic anhydride in a 3-chlorophthalic anhydride separation tower, wherein the purity of the 3-chlorophthalic anhydride is more than 99.8wt%. According to the embodiment of the invention, through the design of the multistage separation unit and the optimization of rectification process parameters, the high-purity chlorophthalic anhydride monomer product can be continuously obtained with low energy consumption.

In an alternative manner, fig. 13 shows a schematic structural diagram of an oxidation reactor according to an embodiment of the present invention, and as shown in fig. 13, in the method for producing chlorophthalic anhydride according to the embodiment of the present invention, an oxidation reactor 100 in a production system of chlorophthalic anhydride is used, which comprises an outer cylinder 110, an inner cylinder 120, a filler 130, and a gas distributor 140, wherein the filler 130 comprises a first filler 131 and a second filler 132, and the gas distributor 140 comprises a first gas distributor 141 and a second gas distributor 142. The first filler 131 is located at a position of the outer cylinder 110 near the top of the inner cylinder 120, the first gas distributor 141 is annularly arranged at the outer bottom of the inner cylinder 120, and the second filler 132 and the second gas distributor 142 are both located at the inner bottom of the inner cylinder 120.

As shown in FIG. 13, the oxidation reactor 100 may be a straight cylinder, the diameter of the inner cylinder 120 is 0.4-0.8 times of the diameter of the outer cylinder 110, the height of the inner cylinder 120 is 0.5-0.9 times of the liquid level height of the outer cylinder 110, the outer cylinder 110 is divided into a water concentrating section, a gas-liquid separating section and a gas-liquid reaction section from top to bottom, the water concentrating section accounts for 20-40% of the total height of the oxidation reactor, the gas-liquid separating section accounts for 15-25% of the total height of the oxidation reactor, and the gas-liquid reaction section accounts for 40-70% of the total height of the oxidation reactor; the first filler 131 is arranged on the water concentration section, the first gas distributor 141 is arranged on the lower portion of the gas-liquid reaction section, the first gas distributor 141 can be a ring-pipe type gas distributor or a disc type gas distributor, the gas distributor is connected with a compressed air inlet of the oxidation reactor, the first feeding pipe a is arranged on the gas-liquid reaction section above the gas distributor, the second feeding pipe b is arranged above the gas distributor, and a tail gas discharging pipe c and a condensate return pipe d are arranged on the top of the oxidation reactor 100. The inner cylinder 120 is a gas-liquid reaction section, the lower part of the inner cylinder 120 is provided with a second gas distributor 142, the second gas distributor 142 can be a ring pipe type gas distributor or a disc type gas distributor, the middle upper part of the inner cylinder 120 is provided with a second filler 132 which can be a silk screen filler, and the bottom of the inner cylinder 120 is provided with a discharge pipe. Wherein, the theoretical stage number of the packing or the tower plate of the water concentration section is 1-4, the liquid phase reaction residence time of the outer cylinder 110 is 60-90 minutes, and the liquid phase reaction residence time of the inner cylinder 120 (residence time after oxidation, namely reaction time of secondary oxidation) is not less than 50 minutes. The production system of the embodiment of the invention is provided with the filler at the middle upper part of the outer cylinder and the inner cylinder which are contained in the oxidation reactor to increase the gas-liquid contact area, so that the chloro-o-xylene and intermediate products thereof are fully oxidized, thereby achieving higher conversion rate and yield, and greatly reducing partial oxidation and the generation of byproduct impurities.

In particular, the outer cylinder 110 is a main reaction occurrence area of raw material chloro-ortho-xylene and air, raw material chloro-ortho-xylene, catalyst and acetic acid are fed into the outer cylinder 110 through a first feed pipe a contained in the oxidation reactor 100, and compressed air is fed into the lower part of a first gas distributor 141 contained in the outer cylinder 110 through a second feed pipe b contained in the oxidation reactor 100, so that the air fed from the second feed pipe b is uniformly distributed to the outer cylinder 110 by the first gas distributor 141. At this time, most of the chloro-o-xylene is converted into chlorophthalic acid in the outer cylinder 110, the reacted feed liquid flows downwards from the upper part of the inner cylinder 120 in a manner similar to a plug flow, and then compressed air is fed under a second gas distributor 142 contained in the inner cylinder through a second feed pipe b contained in the oxidation reactor 100, so that the air conveyed by the second feed pipe b is uniformly distributed to the inner cylinder 120 by the second gas distributor 142, and is subjected to oxidation reaction again with the reaction liquid in the inner cylinder 120 to obtain a final oxidation reaction liquid, the final oxidation reaction liquid flows out from the bottom of the inner cylinder 120, and the reaction area of the air and the chloro-o-xylene can be increased by the first gas distributor 141 and the second gas distributor 141, thereby increasing the reaction rate. It should be understood that the oxidation reaction liquid of the outer tub 110 and the final oxidation reaction liquid each include raw material chloro-ortho-xylene, catalyst, acetic acid, water, chlorophthalic acid, and a part of intermediate products, and the gas distributor is at least one of a loop-shaped gas distributor and a disk-shaped gas distributor.

When the first filler 131 is installed at the middle upper part of the outer cylinder 110, and the second filler 132 is installed at the middle upper part of the inner cylinder 110, the feed liquid is the oxidation reaction liquid, the oxidation reaction liquid needs to be contacted with air, and the filler can disperse the oxidation reaction liquid, so that the contact area between the oxidation reaction liquid and the air is increased, and the chloro-o-xylene and intermediate products thereof are fully oxidized to generate the chloro-phthalic acid, thereby achieving higher conversion rate and yield, and greatly reducing the generation of partial oxidation products and byproduct impurities. According to the embodiment of the invention, acetic acid and water vapor which flow back to the oxidation reactor from the tail gas condensing unit, as the upper part of the outer cylinder of the oxidation reactor is provided with the filler, the cold reflux liquid can be in countercurrent contact with the acetic acid vapor and the water vapor which rise from the outer cylinder, so that the volatilization amount of acetic acid is reduced, the content of the discharged water vapor is improved, and the reaction is continuously more controllable.

In an example, fig. 14 shows a schematic structural diagram of an anhydride forming device according to an embodiment of the present invention, and as shown in fig. 14, in the method for producing chlorophthalic anhydride according to the embodiment of the present invention, an anhydride forming device 300 in a production system of chlorophthalic anhydride further includes an oxidation reaction liquid distributor 311 and an anhydride forming condenser 350, where the oxidation reaction liquid distributor 311 is disposed in the anhydride forming reactor 310. The oxidation reaction liquid outlet of the post-treatment unit is communicated with the inlet of the anhydride forming reactor 310, the exhaust port of the anhydride forming reactor 310 is communicated with the inlet of the anhydride forming condenser 350, the outlet of the anhydride forming condenser 350 is divided into two paths, one path is communicated with the inlet of the post-treatment unit, and the other path is communicated with the inlet of the anhydride forming reactor 310.

When the oxidation reaction liquid from the bottom of the stripping tower enters the anhydride forming reactor 310, the oxidation reaction liquid enters the oxidation reaction liquid distributor 311, and the oxidation reaction liquid distributor 311 has a heating function. At this time, the oxidation reaction liquid distributor 311 may uniformly distribute the oxidation reaction liquid into the anhydride-forming reactor 310, the operation pressure of the anhydride-forming reactor 310 is normal pressure, the reaction temperature is 205 ℃ to 250 ℃, and the oxidation product chlorophthalic acid is dehydrated into chlorophthalic anhydride in the anhydride-forming reactor 310. Meanwhile, at this temperature, the reaction liquid in the anhydride-forming reactor 310 is in a gaseous state, and enters the anhydride-forming condenser 350, and the temperature of the anhydride-forming condenser 350 is 90 to 130 ℃.

Therefore, the organic matters in the gaseous reaction liquid can be condensed and reflowed into the anhydride forming reactor, acetic acid vapor and water vapor in the gaseous reaction liquid are reflowed into the stripping tower through the reflow pipe, and acetic anhydride coming out of the anhydride forming reactor is pumped into the evaporator. In the embodiment of the invention, the acetic acid in the anhydride forming reactor is recycled to the stripping tower, and can be further recycled to the oxidation reactor through the acetic acid recycling unit, so that the production cost is reduced, meanwhile, the oxidation reaction liquid feed is firstly heated on the oxidation reaction liquid distributor, and the oxidation reaction liquid is uniformly sprayed on the crude anhydride liquid level in the reactor through the oxidation reaction liquid distributor, thereby being beneficial to the evaporation of acetic acid and water and having higher anhydride forming rate.

Illustratively, as shown in FIG. 14, the anhydride-forming reactor 310 further includes a first agitator 312 and a second agitator 313, the first agitator 312 and the second agitator 323 being configured to agitate the oxidation reaction solution in the anhydride-forming reactor 310. Wherein the first agitator 312 may be an axial flow agitator and the second agitator 313 may be a radial flow agitator. For example: after the oxidation reaction liquid distributor 311 uniformly distributes the oxidation reaction liquid into the anhydride-forming reactor 310, the axial flow stirrer and the radial flow stirrer can uniformly stir the oxidation reaction liquid, so that the reaction rate is improved while the oxidation reaction liquid is uniformly distributed into the anhydride-forming reactor 310.

Example 1

1) Oxidation reaction

67.5kg/h of chloro-ortho-xylene (mass composition: 54.8% of 4-chloro-ortho-xylene, 44.8% of 3-chloro-ortho-xylene, 0.3% of ortho-xylene, 0.1% of dichloro-ortho-xylene, 284.9kg/h of solvent acetic acid and 0.85kg/h of catalyst cobalt acetate, 0.42kg/h of manganese acetate and 0.50kg/h of tetrabromoethane are mixed and then enter a gas-liquid reaction section of an outer cylinder of a bubble column reactor from a raw material feed pipe, 260.4kg/h of compressed air (oxygen mole content 21%) enters the gas-liquid reaction section of the outer cylinder from an air inlet of the outer cylinder of the bubble column reactor through a loop-shaped gas distributor, and is subjected to oxidation reaction with the chloro-ortho-xylene to generate chloro-phthalic acid, the reacted feed liquid enters the inner cylinder from the upper part of the inner cylinder, flows downwards in a mode similar to a plug flow, 13.0kg/h of compressed air (oxygen mole content 21%) enters from the lower part of the inner cylinder, flows reversely with the feed liquid of the inner cylinder after passing through the loop-shaped gas distributor and undergoes secondary oxidation reaction, the oxidation reaction temperature is 227.6 ℃, the pressure is 2.2MPa (G), the retention time of the outer cylinder is 70min, the retention time of the oxidation reaction is 60% of the oxygen content is 5% of the water in the bubble column from the top of the gas-column reactor, the end product is discharged from the outlet of the end product is discharged from the end of the gas-stream of the oxidation reactor, the oxidation reactor is discharged from the end of the end product of the oxidation reactor, the end product is discharged from the end of the oxidation reactor, the end of the product is discharged from the oxidation reaction product is discharged from the end of the product is discharged from the product of the product is discharged from the product is discharged.

The bubble column reactor of this embodiment external diameter is 700mm, internal diameter is 500mm, urceolus liquid level height is 2350mm, and the inner tube height is 1700mm, and inner tube silk screen packing height is 700mm, and the theoretical plate number of urceolus upper portion packing is 4, and bubble column reactor (urceolus) total height is 8230mm, and liquid level and equipment height datum line are the lower tangent head.

2) Tail gas treatment

The oxidation tail gas from the top of the bubble column reactor is sequentially subjected to step-by-step heat recovery through a first tail gas condenser, a second tail gas condenser and a third tail gas condenser, wherein the first condenser byproducts 0.4MPa (G) steam, the second condenser byproducts 0.1MPa (G) steam, the third condenser cools the tail gas by desalted water at 40 ℃, the tail gas temperature from the tube of the third condenser is 165 ℃, 140 ℃ and 102 ℃, then the tail gas enters a high-pressure tail gas absorption tower, acetic acid in the tail gas in the absorption tower is absorbed by desalted water, the high-pressure tail gas absorption tower adopts a packing tower, the operating pressure is 2.0MPa (G), the operating temperature is 83-94 ℃, the theoretical plate number is 3, the desalted water consumption is 24.5kg/h, the concentration of an aqueous acetic acid solution is 25.02wt%, and the aqueous acetic acid solution is sent into a subsequent acetic acid recovery tower, and the acetic acid content in the tail gas from the top of the absorption tower is 6.1 multiplied by 10- ⁵ And (5) treating the waste gas entering a tail gas treatment device of a factory and then evacuating the waste gas after reaching the standard. The condensate liquid from the three-stage condenser is aqueous solution containing acetic acid, most of the condensate liquid flows back into the oxidation tower, and the other part of the condensate liquid is sent into the subsequent acetic acid recovery tower, wherein the reflux ratio of the condensate liquid is 24.

3) Acetic acid recovery

The material from the bottom of the inner barrel of the oxidation tower enters a flash tank for flash evaporation, the operating pressure is normal pressure, the operating temperature is 120.6 ℃, the flash evaporated acetic acid gas enters an acetic acid recovery tower for rectification, and the flash evaporated liquid material enters a stripping tower; the stripper adopts a packed tower, the operating pressure is normal pressure, the tower bottom temperature is 182 ℃, the tower top temperature is 114.2 ℃, the reflux ratio is 0.098, the theoretical plate number is 6, the reboiler adopts an external forced circulation heating mode to heat and vaporize most of the residual acetic acid and water, and the acetic acid and water vapor are partially condensed and purified by an upper built-in condenser and then enter an acetic acid recovery tower for rectification. The feeding of the acetic acid recovery tower has a gas phase and a liquid phase, and a stream of acetic acid aqueous solution from the bottom of the high-pressure tail gas absorption tower enters from the upper part; a stream of condensate from a three-stage tail gas condenser enters from the lower part; a flash vapor from the flash tank entering from the lower portion; a stream of acetic acid and water vapor from the top of the stripping column enters from the lower part. Acetic acid recovery tower rectifying separation acetic acid and water, operating pressure is normal pressure, tower top temperature is 99.3 ℃, reflux ratio is 3.12, theoretical plate number is 34, concentration of acetic acid separated at tower bottom is 98.2wt%, the acetic acid is circulated into oxidation reactor through recycling pipe, acetic acid content in waste water separated at tower top is 0.1wt%, the waste water is sent to waste water treatment device, noncondensable gas is sent to tail gas treatment device of factory to treat and reach standard, and then is emptied.

4) Anhydride formation

The oxidation product from the bottom of the stripping tower enters an anhydride-forming reactor, and the oxidation product chlorophthalic acid is dehydrated into anhydride in the anhydride-forming reactor. The anhydride forming reactor adopts a double-layer stirrer, the upper layer is an axial flow impeller type stirrer, the lower layer is a radial flow blade type stirrer, a liquid distribution disc is arranged at the upper part of a stirring shaft, liquid feed flows to the upper part of the disc and is sprayed onto the liquid surface below the disc through the periphery of the disc, the operation pressure of the anhydride forming reactor is normal pressure, the operation temperature is 240 ℃, more organic matters are carried out by evaporated acetic acid and water, the organic matters are partially condensed by a condenser, the condensation temperature is 130-140 ℃, the organic matters are condensed and flow back into the anhydride forming reactor, and acetic acid and water gas are sent into an acetic acid recovery tower for rectification and recovery after entering a stripping tower for purification. The crude anhydride from the anhydride-forming reactor enters a scraper film evaporator, the crude anhydride is continuously vaporized, the catalyst and the multi-ring heavy component residual liquid are separated, the operating pressure is 5kPa (A), and the operating temperature is 231 ℃. The evaporated residual liquid is sent to a catalyst recovery device for recovering the catalyst, the crude anhydride steam is condensed to 160 ℃ by a crude anhydride condenser, the liquid enters a crude anhydride storage tank for buffering, and then a pump is used for refining by a rectifying device.

5) Rectifying

The crude anhydride is continuously subjected to product separation in three vacuum rectification columns. Firstly, removing light components such as phthalic anhydride and the like in a phthalic anhydride separating tower, then separating 4-chlorophthalic anhydride in a second tower, namely a 4-chlorophthalic anhydride separating tower, and finally separating 3-chlorophthalic anhydride in a 3-chlorophthalic anhydride separating tower. The operation pressure at the top of the phthalic anhydride separation column is 5kPa (A), the temperature is 184.6 ℃, the reflux ratio is 230, the theoretical plate number is 39, the mixture mainly containing phthalic anhydride is obtained at the top of the column, the phthalic anhydride content is 70.9wt%, the chlorophthalic anhydride is 27.8wt%, and the chlorobenzoic acid is 1.3wt%. The operating pressure at the top of the 4-chlorophthalic anhydride separation tower is 2.1kPa (A), the temperature is 174.6 ℃, the reflux ratio is 4.2, the theoretical plate number is 34,4, the yield of the chlorophthalic anhydride is 44.2kg/h, and the purity is 99.6wt%; the operating pressure at the top of the 3-chlorophthalic anhydride separation column is 2.1kPa (A), the temperature is 208 ℃, the reflux ratio is 3.8, the theoretical plate number is 28, the 3-chlorophthalic anhydride yield is 35.0kg/h, and the purity is 99.8wt%. The reboiler types of the three vacuum rectifying towers all adopt falling film type reboilers.

Example two

1) Oxidation reaction

122.2kg/h of feed chloro-ortho-xylene (mass composition: 54.93 percent of 4-chloro-ortho-xylene, 45.02 percent of 3-chloro-ortho-xylene, 0.05 percent of ortho-xylene, 516kg/h of solvent acetic acid (the mass composition: acetic acid 98.0 percent, water 2 percent) and 1.44kg/h of catalyst cobalt acetate, 0.48kg/h of manganese acetate and 1.2kg/h of tetrabromoethane are mixed and then enter a gas-liquid reaction section of an outer cylinder of a bubble column reactor from a raw material feeding pipe, simultaneously 476kg/h of compressed air (the molar content of oxygen) enters the gas-liquid reaction section of the outer cylinder after passing through a loop-shaped gas distributor from an air inlet of the outer cylinder of the bubble column reactor, and is subjected to oxidation reaction with the chloro-ortho-xylene to generate chloro-phthalic acid, the reacted feed liquid enters the inner cylinder from the upper part of the inner cylinder in a mode of approximately plug flow, simultaneously 19kg/h of compressed air (the molar content of oxygen 21 percent) enters from the lower part of the inner cylinder, flows reversely with the feed liquid of the inner cylinder after passing through the loop-shaped gas distributor and is subjected to secondary oxidation reaction, the oxidation reaction temperature is 1.2MPa (G), the reaction time is 90min, the retention time is 60 wt% of the oxygen content of the oxygen is 5.49 percent of the end product is discharged from the top of the gas-phase oxidation reaction device, the end product is discharged from the top of the oxidation reaction device, the oxidation reaction device is discharged from the outlet of the end of the oxidation reaction device, the end product is discharged from the outlet of the oxidation reaction device is discharged from the end of the oxidation reaction device, and the end product is discharged from the oxidation reaction device is discharged from the end of the oxidation reaction product is discharged from the oxidation reaction is discharged.

The bubble column reactor of this embodiment external diameter is 1000mm, internal diameter is 700mm, and urceolus liquid level height is 2450mm, and the inner tube height is 1590mm, and inner tube silk screen packing height 800mm, and the theoretical plate number of urceolus upper portion packing is 3, and bubble column reactor (urceolus) total height is 9200mm, and liquid level and equipment height datum line are the lower tangent head.

2) Tail gas treatment

The oxidation tail gas from the top of the bubble column reactor is subjected to gradual heat recovery sequentially through a first tail gas condenser, a second tail gas condenser and a third tail gas condenser, wherein the first condenser is used for producing steam with the pressure of 0.4MPa (G), the second condenser is used for producing steam with the pressure of 0.1MPa (G), the third condenser is used for cooling the tail gas by desalted water with the temperature of 40 ℃, the tail gas from the tube of the third condenser is 160 ℃, 135 ℃ and 60 ℃ sequentially, then the tail gas enters a high-pressure tail gas absorption tower, acetic acid in the tail gas in the absorption tower is absorbed by desalted water, the high-pressure tail gas absorption tower adopts a packing tower, the operating pressure is 1.0MPa (G), the operating temperature is 53-56 ℃, the theoretical plate number is 6, the consumption of desalted water is 15kg/h, the concentration of an acetic acid aqueous solution is 23.3wt%, the acetic acid aqueous solution is fed into a later acetic acid recovery tower, the acetic acid content in the tail gas from the top of the absorption tower is less than 1PPM, and the tail gas from the absorption tower enters a tail gas treatment device of a factory for treatment to reach the standard and is emptied; the condensate liquid from the three-stage condenser is aqueous solution containing acetic acid, most of the condensate liquid flows back into the oxidation tower, and the other part of the condensate liquid is sent into the subsequent acetic acid recovery tower, wherein the condensate liquid reflux ratio is 32.

3) Acetic acid recovery

The material from the bottom of the inner barrel of the oxidation tower enters a flash tank for flash evaporation, the operating pressure is normal pressure, the operating temperature is 118.4 ℃, the flash evaporated acetic acid gas enters an acetic acid recovery tower for rectification, and the flash evaporated liquid material enters a stripping tower; the stripping tower adopts a composite tower with upper filler and lower sieve plates combined, the operating pressure is normal pressure, the tower kettle temperature is 190 ℃, the tower top temperature is 114 ℃, the reflux ratio is 0.25, the theoretical plate number is 6, the reboiler adopts an external forced circulation heating mode to heat and vaporize most of the residual acetic acid and water, and the acetic acid and water vapor are partially condensed and purified by an upper built-in condenser and then enter an acetic acid recovery tower for rectification. The operation pressure of the acetic acid recovery tower is normal pressure, the tower top temperature is 99.9 ℃, the reflux ratio is 4.56, the theoretical plate number is 30, the concentration of acetic acid separated at the tower bottom is 99.0wt%, the acetic acid enters the oxidation reactor in a circulating way through a recycling pipe, the acetic acid content in the wastewater separated at the tower top is 0.3wt%, the wastewater is sent to a wastewater treatment device, and the non-condensable gas is sent to a tail gas treatment device of a factory to be treated and then is emptied after reaching the standard.

4) Anhydride formation

The oxidation product from the bottom of the stripping tower enters an anhydride-forming reactor, and the oxidation product chlorophthalic acid is dehydrated into anhydride in the anhydride-forming reactor. The anhydride forming reactor adopts a double-layer stirrer, the upper layer is an axial flow impeller type stirrer, the lower layer is a radial flow blade type stirrer, a liquid distribution disc is arranged on the upper part of a stirring shaft, liquid feed flows to the upper part of the disc and is sprayed onto the liquid surface below the disc through the periphery of the disc, the operation pressure of the anhydride forming reactor is normal pressure, the operation temperature is 230 ℃, more organic matters are carried out by evaporated acetic acid and water, the acetic acid and the water are partially condensed by a condenser, the condensation temperature is 130 ℃, the condensed organic matters flow back into the anhydride forming reactor, and acetic acid and water gas are sent into an acetic acid recovery tower for rectification and recovery after entering a stripping tower for purification. Feeding the crude anhydride from the anhydride-forming reactor into a scraper film evaporator, continuously vaporizing the crude anhydride, separating out catalyst and multi-ring heavy component residual liquid, wherein the operating pressure is 2.1kPa (A), and the operating temperature is 210 ℃; the evaporated residual liquid is sent to a catalyst recovery device for recovering the catalyst, the crude anhydride steam is condensed to 90 ℃ by a crude anhydride condenser, the liquid enters a crude anhydride storage tank for buffering, and then a pump is used for refining by a rectifying device.

5) Rectifying

The crude anhydride is continuously subjected to product separation in three vacuum rectification columns. Firstly, removing light components such as phthalic anhydride and the like in a phthalic anhydride separating tower, then separating 4-chlorophthalic anhydride in a second tower, namely a 4-chlorophthalic anhydride separating tower, and finally separating 3-chlorophthalic anhydride in a 3-chlorophthalic anhydride separating tower. The operation pressure at the top of the phthalic anhydride separation tower is 5kPa (A), the temperature is 183.5 ℃, the reflux ratio is 270, the theoretical plate number is 40, the mixture mainly containing phthalic anhydride is obtained at the top of the tower, the phthalic anhydride content is 86.8wt%, the chlorophthalic anhydride is 12.6wt%, the chlorobenzoic acid is 0.6wt%; the operating pressure at the top of the 4-chlorophthalic anhydride separation tower is 2.1kPa (A), the temperature is 174.4 ℃, the reflux ratio is 5, the theoretical plate number is 38,4, the yield of the chlorophthalic anhydride is 80.9kg/h, and the purity is 99.8wt%; the operating pressure at the top of the 3-chlorophthalic anhydride separation column is 2.1kPa (A), the temperature is 208.3 ℃, the reflux ratio is 4.5, the theoretical plate number is 35, the 3-chlorophthalic anhydride yield is 62.1kg/h, and the purity is 99.8wt%. The reboiler types of the three vacuum rectifying towers all adopt falling film type reboilers.

Comparative example one

The comparative example one has the same operating parameters as example one, but using a conventional batch process, only one oxidation reactor was used, without the outer drum, inner drum, and continuous apparatus of the present application, and without the post-treatment unit and multistage separation unit of the present application.

Comparative example two

The operating parameters of comparative example two were the same as in example two, but multiple oxidation reaction vessels in series were used, without the outer and inner drums of the present application, and without the post-treatment unit and multistage separation unit of the present application.

The effects of examples of a production system of chlorophthalic anhydride of the present invention versus comparative examples are shown in the following table:

as is clear from the above table, the product yields and the quality stability and purity of the first and second examples of the present invention are higher than those of the first and second comparative examples. Meanwhile, the 4-chlorophthalic anhydride and the 3-chlorophthalic anhydride can be separated and higher purity can be obtained in the first and second embodiments. And, the amounts of o-xylene chloride and acetic acid consumed in examples one and two were significantly less than in the comparative example. Therefore, the system for producing the chlorophthalic anhydride can obtain higher yield and purity under the condition of consuming a small amount of raw materials and acetic acid. Meanwhile, the batch method cannot keep the operation of each batch the same, and because the human factors are different, the repeatability is poor, and therefore the purity of each batch of the product is different, and the stability is poor. The production system of the chlorophthalic anhydride is full continuous flow, the logistics are continuous, and the raw materials, the equipment and the operation conditions are the same, so that the change of the quality and the purity of the product is small, and the stability is strong.

The foregoing is merely illustrative of the present invention, and the present invention is not limited thereto, and any person skilled in the art will readily recognize that variations or substitutions are within the scope of the present invention. Therefore, the protection scope of the present invention shall be subject to the protection scope of the claims.

Claims

1. A process for the production of chlorophthalic anhydride, said process comprising:

2. The method for producing chlorophthalic anhydride according to claim 1, characterized in that said separating the gas phase product from acetic acid to obtain acetic acid and purifying the reaction liquid comprises:

Condensing the gas-phase product to obtain an acetic acid solution, separating acetic acid from the acetic acid solution, and sending the acetic acid solution into the oxidation reactor to continue oxidation reaction;

and (3) acetic acid separation is carried out on the oxidation reaction liquid to obtain a purified reaction liquid and a gas phase component containing acetic acid, and then the acetic acid is separated from the gas phase component.

3. The method for producing chlorophthalic anhydride according to claim 2, characterized in that said condensing said vapor phase product to obtain an acetic acid solution comprises:

condensing the gas-phase product to obtain a first acetic acid solution and tail gas;

absorbing the tail gas and separating the residual second acetic acid solution in the tail gas by using a tail gas absorption unit;

separating acetic acid from the first acetic acid solution and the second acetic acid solution.

4. The method for producing chlorophthalic anhydride according to claim 2, characterized in that after said condensing treatment of said gas phase product to obtain an acetic acid solution, said method further comprises, before separating acetic acid in said acetic acid solution:

dividing the acetic acid solution into a first path of acetic acid solution and a second path of acetic acid solution, refluxing the first path of acetic acid solution to the oxidation reactor for continuous reaction, and delivering the second path of acetic acid solution to the acetic acid recovery unit for recovery.

5. The method for producing chlorophthalic anhydride according to claim 2, characterized in that said separating acetic acid from said oxidation reaction liquid to obtain a purified reaction liquid and a gas phase component containing acetic acid comprises:

acetic acid separation is carried out on the oxidation reaction liquid to obtain a first purified reaction liquid and a first gas phase component containing acetic acid;

acetic acid separation is carried out on the first purification reaction liquid, so that a second purification reaction liquid and a second gas phase component containing acetic acid are obtained;

and sending the second purified reaction liquid into an anhydride forming reactor to perform an anhydride forming reaction.

6. The method for producing chlorophthalic anhydride according to claim 5, wherein when said purified reaction solution is dehydrated to obtain crude anhydride, said method further comprises:

and dehydrating the second purification reaction liquid to obtain an anhydride-forming crude product containing at least acetic acid and chlorophthalic anhydride.

7. The method for producing chlorophthalic anhydride according to claim 1, characterized in that the crude anhydride-forming product is a gaseous crude anhydride-forming product, and the method further comprises, after dehydrating the purified reaction solution to obtain the crude anhydride-forming product:

separating acetic acid contained in the anhydride-forming crude product and condensing the anhydride-forming crude product to obtain a condensed anhydride-forming crude product;

And sending the acetic acid separated from the anhydride-forming crude product into a post-treatment unit for post-treatment.

8. The method for producing chlorophthalic anhydride according to claim 1, characterized in that after dehydration of the purified reaction solution to obtain crude anhydride, the method further comprises, before separating at least two chlorophthalic anhydride monomers from the crude anhydride, respectively:

and separating the catalyst contained in the crude anhydride-forming product, and condensing the crude anhydride-forming product after separating the catalyst.

9. The method for producing chlorophthalic anhydride according to claim 1, characterized in that at least two chlorophthalic anhydride monomers are separated from the crude anhydride-forming product, respectively, comprising:

separating light components from the crude anhydride-forming product by using a first separation unit;

separating 4-chlorophthalic anhydride from the crude anhydride-forming product by a second separation unit;

and separating 3-chlorophthalic anhydride from the crude anhydride-forming product by a third separation unit.

10. The method for producing chlorophthalic anhydride according to claim 1, characterized in that the steps of oxidizing in an oxidation reactor with chloro-o-xylene and acetic acid as raw materials and air as an oxidizing agent under the action of a catalyst to obtain a gas phase product and an oxidation reaction liquid, comprise:

Air is distributed in the area between the outer cylinder and the inner cylinder, chloro-o-xylene and acetic acid are used as raw materials under the action of a catalyst, and air is used as an oxidant, so that a first oxidation reaction is carried out in the outer cylinder;

air is distributed to the inner cylinder, chloro-o-xylene and acetic acid are used as raw materials under the action of a catalyst, air is used as an oxidant, and a second oxidation reaction is carried out in the inner cylinder to obtain a gas-phase product and an oxidation reaction liquid.