WO2016101517A1 - 制备氯甲酰基取代苯的清洁工艺 - Google Patents

制备氯甲酰基取代苯的清洁工艺 Download PDF

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WO2016101517A1
WO2016101517A1 PCT/CN2015/079272 CN2015079272W WO2016101517A1 WO 2016101517 A1 WO2016101517 A1 WO 2016101517A1 CN 2015079272 W CN2015079272 W CN 2015079272W WO 2016101517 A1 WO2016101517 A1 WO 2016101517A1
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gas stream
hydrogen chloride
containing gas
reactor
oxygen
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French (fr)
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王农跃
瞿雄伟
李国华
邵建明
赵全忠
闻国强
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上海方纶新材料科技有限公司
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Priority to KR1020177020136A priority Critical patent/KR102360688B1/ko
Priority to RU2017126032A priority patent/RU2676310C1/ru
Priority to EP15871593.8A priority patent/EP3239130B1/en
Priority to JP2017531997A priority patent/JP6615205B2/ja
Priority to ES15871593T priority patent/ES2902248T3/es
Publication of WO2016101517A1 publication Critical patent/WO2016101517A1/zh
Priority to IL253047A priority patent/IL253047B/en
Priority to US15/629,617 priority patent/US10196340B2/en

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    • C07C63/34—Polycyclic acids with carboxyl groups bound to condensed ring systems containing two condensed rings

Definitions

  • the invention belongs to the technical field of chemical industry and relates to a cleaning process for preparing chloroformyl substituted benzene.
  • the process of the present invention passes the oxidative chlorination reaction and the tail gas hydrogen chloride of the acid chloride reaction, and the obtained oxidation product chlorine gas is recycled to the chlorination reaction.
  • the present invention is a cleaning process for preparing polymeric grade chloroformyl substituted benzenes.
  • the preparation method of chloroformyl substituted benzene mainly includes photochlorination method (see DE31 468 68, JP 47-130931), thionyl chloride method, phosphorus trichloride method, phosphorus pentachloride method and phosgene method. Wait.
  • the thionyl chloride process is most commonly used (see, for example, CN102516060A, CN102344362A), but requires 99.99% high purity phthalic acid as the starting material, which makes the process costly.
  • these methods all have problems with environmentally unfriendly by-products such as hydrogen chloride, sulfur dioxide, carbon dioxide, and phosphorous acid. These by-products cause inconvenience to the subsequent processing of the product and are highly likely to cause environmental pollution.
  • the photochlorination method can use a methyl aromatic compound as a raw material, but the amount of hydrogen chloride as a by-product is enormous. How to deal with a large amount of hydrogen chloride has become an urgent problem to be solved.
  • the main treatment measures actually adopted in the industry are the sale of hydrogen chloride after absorption of hydrogen chloride by water; due to the low price of hydrochloric acid and limited market demand, the production of hydrogen chloride into hydrochloric acid has actually become a burden rather than a waste. For treasure.
  • Some of the treatments used are direct discharges after neutralizing hydrogen chloride with alkali; however, as environmental laws and regulations are becoming more sophisticated, the environmental standards for various emission methods are already very strict.
  • the method of directly forming the by-product hydrogen chloride into chlorine gas can not only realize the closed loop of chlorine, but also realize the zero discharge of the reaction process, greatly improve the energy saving and emission reduction level of the industry, reduce the cost and eliminate the pollution to the environment. So far, the methods for preparing chlorine gas from hydrogen chloride can be mainly divided into three categories: electrolysis, direct oxidation and catalytic oxidation.
  • the energy consumption of the electrolysis process is too large, and the ion membrane needs to be replaced frequently, the cost is very high, the recovery cost per ton of chlorine gas is >4000 yuan; the yield of the direct oxidation method is low, and it cannot be industrialized; and the electrolysis method, the direct oxidation method
  • the catalytic oxidation process in particular the catalytic oxidation process via the Deacon reaction, has the greatest industrial potential.
  • the process of the invention solves the problems existing in the existing industry, realizes the closed circuit circulation of the chlorine resources, eliminates the pollution caused by the by-products from the source, and the product obtained by the process has low cost and high quality.
  • the present invention is achieved by the following technical scheme: firstly, a methyl arene of the formula (X) a C 6 H 6-ab (CH 3 ) b or a thiol side chain chloride of the compound and chlorine gas (for example, under illumination conditions)
  • the reaction is to prepare a trichloromethyl-substituted benzene; the produced trichloromethyl-substituted benzene is further reacted to prepare a chloroformyl-substituted benzene; the produced HCl gas is catalytically oxidized by a Deacon reaction to be a chlorine gas, and then used in the benzene gas.
  • the methyl aromatic hydrocarbon is chlorinated to prepare trichloromethyl substituted benzene.
  • the representative reaction of the entire process is as follows:
  • the chlorine gas obtained by oxidation is reintroduced as a raw material to the chlorination reaction, and the overall reaction equation for the preparation of the bis(chloroformyl)benzene of the present invention is:
  • the (X) a C 6 H 6-ab (CH 3 ) b is a methyl arene compound (the alkyl side chain chloride of the compound is also suitable for use in the present invention), (X) a C 6 H 6-ab (CCl 3 ) b is trichloromethyl substituted benzene, (X) a C 6 H 6-ab (COOH) b is the corresponding aromatic acid, (X) a C 6 H 6-ab (COCl) b is chlorine Acyl substituted benzene, in the above formula of the compound described herein, X is chlorine or a bromine or fluorine atom, a is an integer selected from 0, 1, 2, 3, 4 or 5, and b is selected from 1, 2, 3 Or an integer of 4, and a+b ⁇ 6.
  • the corresponding aromatic acid as used in the present application means that the substituent on the aromatic acid nucleus is at the same or corresponding substitution position as the substituent on the methyl aryl hydrocarbon core; the substituent on the aromatic acid nucleus is The substituents on the methylene aromatic hydrocarbon core may also be the same.
  • the alkyl side chain chloride of the methyl aromatic hydrocarbon compound as used herein refers to a compound in which the hydrogen atom on the sulfhydryl group is not all replaced by a chlorine atom; the target of the photochlorination reaction described herein
  • the product, trichloromethyl-substituted benzene refers to a product in which all of the hydrogen atoms on the alkyl group in the aromatic hydrocarbon compound are replaced by chlorine atoms.
  • a cleaning process for preparing chloroformyl substituted benzene comprising the steps of:
  • Step 1 chlorination reaction, reacting the methyl aromatic hydrocarbon of the formula (X) a C 6 H 6-ab (CH 3 ) b or its alkyl side chain chloride and chlorine gas (for example, under light) Preparing trichloromethyl substituted benzene and obtaining by-product hydrogen chloride;
  • Step 2 acyl chloride reaction
  • the trichloromethyl-substituted benzene prepared in the first step is further reacted with the corresponding aromatic acid or water of the formula (X) a C 6 H 6-ab (COOH) b to prepare a chlorine group.
  • X aromatic acid or water of the formula (X) a C 6 H 6-ab (COOH) b
  • a chlorine group Acyl substituted benzene and obtained by-product hydrogen chloride;
  • Step 3 catalyzing oxidation by-product hydrogen chloride), collecting the by-product hydrogen chloride in the above steps 1 and 2, and performing catalytic oxidation reaction (Deacon reaction) through the catalyst to prepare chlorine gas;
  • step four separating the gas stream from step three
  • a chlorine-containing, oxygen-containing, and/or hydrogen-containing hydrogen gas stream is separated from the product gas stream in step three above.
  • Step 5 (recycling the separated product), introducing the chlorine-containing gas stream separated in the above step 4 as a raw material into the chlorination reaction including the first step;
  • the hydrogen chloride-containing and/or oxygen-containing gas stream separated in the above step four is introduced as a raw material into the reaction of the catalytic oxidation by-product hydrogen chloride in the third step.
  • the purification step of trichloromethyl substituted benzene may be further present or absent after the chlorination reaction; further may be present after the acid chloride reaction or There is no purification step for the chloroformyl substituted benzene.
  • the process of the present invention without any abandonment of the normal loss during the purification process and the reaction process, is a completely green chemical process.
  • the process of the invention realizes the clean production of the chloroformyl-substituted benzene, especially the raw material polymerization grade bis(chloroformyl)benzene, and has important economic and social benefits for producing high-performance aramid fiber at low cost.
  • the by-product hydrogen chloride produced in the chlorination and acid chloride step is further catalytically oxidized to obtain chlorine gas, and obtained.
  • the chlorine gas is chlorinated to achieve a closed loop of chlorine, which reduces production costs and reduces environmental pollution.
  • the third catalytic oxidation step of the present invention is the key and core for realizing the recycling of chlorine resources.
  • the process of the Deacon reaction is realized by the technique of directly recycling the product gas stream obtained by catalytic oxidation of hydrogen chloride without separation.
  • the dispersion of heat extends the life of the catalyst, while the heat carried by the recycled product gas stream reduces the cost of preheating the hydrogen chloride-containing feed gas, further saving the cost of industrialization.
  • step four the present invention separates the product gas stream from step three using a separation process comprising a condensation, drying, adsorption step, optionally further comprising a liquefaction separation step.
  • the separation method of the present invention does not produce a large amount of dilute hydrochloric acid due to the absence of a water washing step.
  • no further liquefaction separation treatment step is required.
  • the chlorine gas containing a small amount of hydrogen chloride gas is directly recycled to the chlorination step such as the first step, the presence of a small amount of hydrogen chloride does not affect the reaction of the chlorine gas with the methyl aromatic hydrocarbon to produce trichloromethyl substituted benzene.
  • the separation method of the product gas stream in the fourth step of the invention has the simple process flow and environmental friend Good, low energy consumption, high separation efficiency, low cost, etc.
  • the purity of chlorine gas in the separated and recovered chlorine-containing gas stream is ⁇ 99.6% (vol%), and such a chlorine-containing gas stream can satisfy the photochlorination reaction to the chlorine gas raw material. Gas quality requirements.
  • step three in addition to the closed loop of the chlorine removal resource, other substances generated in the production process of the product can also be recycled, thereby achieving clean production.
  • step three the unreacted hydrogen chloride and/or oxygen in the catalytic oxidation step is subjected to the catalytic oxidation reaction again after separation.
  • the process of the present invention can obtain a polymerization grade chloroformyl substituted benzene such as bis(chloroformyl)benzene, and the production cost is reduced by more than 30% compared with the conventional process.
  • Figure 1 is a flow chart showing the cleaning process for preparing bis(chloroformyl)benzene.
  • the inventors of the present invention creatively integrated the synthesis of methylated aromatic chlorinated and chloroformyl substituted benzene with chlorine chloride to form a complete process (while improving the hydrogen chloride oxidation process and the separation process of the mixed gas).
  • the recycling of chlorine is achieved by catalytically oxidizing a large amount of hydrogen chloride generated in the chlorination and acid chlorination to chlorine gas and introducing the obtained chlorine gas into the chlorination process.
  • the entire process is a clean production process.
  • Step 1 chlorination reaction
  • the methyl aromatic hydrocarbon of the formula (X) a C 6 H 6-ab (CH 3 ) b or the alkyl side chain chloride of the compound and chlorine gas (for example, under light conditions)
  • the reaction produces trichloromethyl-substituted benzene and obtains by-product hydrogen chloride, wherein X is chlorine or a bromine or fluorine atom, a is an integer selected from 0, 1, 2, 3, 4 or 5, and b is selected from 1, An integer of 2, 3 or 4, and a+b ⁇ 6.
  • the alkyl side chain chloride of the methyl aromatic hydrocarbon compound as used herein means a compound in which the hydrogen atom on the alkyl group in the aromatic hydrocarbon compound is not completely substituted by a chlorine atom; the target of the photochlorination reaction described herein
  • the product, trichloromethyl-substituted benzene refers to a product in which all of the hydrogen atoms on the alkyl group in the aromatic hydrocarbon compound are replaced by chlorine atoms.
  • the obtained trichloromethyl-substituted benzene is optionally further purified or directly introduced into the acid chloride reaction, and the obtained by-product hydrogen chloride is recovered for use in the third step.
  • the chlorination reaction of the present invention relates to a photochemical process for the preparation of trichloromethyl substituted benzene, characterized in that the methyl arene of the formula (X) a C 6 H 6-ab (CH 3 ) b or the alkane of the compound
  • the pendant metal chloride and chlorine are reacted under light to prepare trichloromethyl substituted benzene, wherein the light source has a wavelength of about 350 nm to 700 nm and a light wave amplitude of about 200 nm, wherein the reaction temperature is about 0 ° C to 85 ° C.
  • Chlorine gas is introduced at an illuminance of about 2000 Lux to about 55000 Lux, and the first reaction stage is carried out at a reaction temperature of not more than about 120 ° C under the illuminance; then the remaining amount of chlorine gas is continuously passed at a higher reaction temperature until the reaction is completed.
  • the light source is preferably an LED lamp.
  • the present inventors have found that it is advantageous to increase the temperature and illuminance after the first reaction stage of chlorination, preferably by consuming a chlorine gas in an amount of at least about 1/6 of the total amount of chlorine required for the reaction.
  • the first reaction stage consumes from about 1/6 to about 1/2 of the total amount of chlorine required for the reaction; preferably, the first reaction stage consumes about 1/ of the total amount of chlorine required for the reaction. 4- about 1/3.
  • the reaction temperature is preferably from about 55 to 85 °C.
  • the illuminance is about 5000 Lux-about 55000 Lux, preferably from about 20,000 Lux to about 55,000 Lux, more preferably from about 35,000 Lux to about 45,000 Lux.
  • the inventors have found that the reaction after the first reaction stage of chlorination is to pass the balance chlorine gas at a reaction temperature of not more than about 350 ° C and an illuminance of not more than about 100,000 Lux.
  • the process after the first reaction stage of chlorination may be a single reaction stage, or may be divided into several reaction stages, for example, divided into two, three, four, five, six, seven, eight, nine, ten, etc. Reaction stage.
  • the illuminance is optionally also increased during the temperature rise of each stage during the first reaction stage.
  • the process after the first reaction stage of the photochlorination reaction can be further divided into a second reaction stage and a third reaction stage.
  • the second reaction stage controls the reaction temperature to be about 120 to about 160 ° C
  • the incident illuminance is about 10,000 to about 70,000 Lux
  • the amount of chlorine gas introduced is 1/4 to 2/5 of the total amount
  • the third reaction stage controls the temperature to be about 160.
  • incident illuminance of about 50,000 to about 100,000 Lux
  • the balance of chlorine gas In the second and third stages, the elevated temperature and the elevated illuminance can be in a sequential order.
  • the LED in the chlorination preferably has a peak wavelength in the range of from 350 nm to 490 nm, or preferably a peak wavelength in the range of from 460 nm to 490 nm.
  • the source of light has a wavelength of light at most about 50 nm, preferably from about 10 to about 30 nm, more preferably from about 10 to about 25 nm.
  • the chlorination reaction does not contain an additional solvent and an initiator in the reaction system.
  • the total amount of chlorine gas chlorinated in the present invention is the amount of chlorine gas which can completely chlorinate the side chain hydrogen atom of the methyl aromatic hydrocarbon, and the total amount of chlorine gas is at least the theoretical molar amount of the chlorination of the starting methyl aromatic hydrocarbon compound.
  • the total amount of chlorine is at least six times the molar amount of the raw material of di(methyl)benzene.
  • the total amount of chlorine in the chlorination of the present invention is a molar amount more than six times the number of moles of di(methyl)benzene; the excess amount of chlorine gas can be conventionally determined.
  • the amount of chlorine gas introduced in each stage described herein can also be appropriately adjusted according to the reaction monitoring result.
  • the optical wave amplitude of the present invention refers to a wavelength range at a half maximum of the light emitted by the light source, and does not refer to a peak wavelength of a certain light.
  • a light wave amplitude of 50 nm means that the wavelength range of the half-height of the light emitted by the light source does not exceed 50 nm.
  • the peak wavelength of the LED light source of the present invention can be varied within the range of 350 nm to 700 nm.
  • the incident light source of the present invention can achieve a control wave amplitude within 50 nm, for example, 465 nm is a peak amplitude of 50 nm, and 360 nm is a peak amplitude.
  • the LED light source also has the advantage of low heat generation, so that the cost of the production equipment can be reduced, for example, no additional cooling device is needed, and the corresponding cooling device is required for the photochlorination reaction of the high-pressure mercury lamp source (see, for example, US5514254). .
  • the illuminance described in the present invention can be measured by a conventional apparatus in the art, such as an illuminance meter or the like.
  • the wavelengths described in the present invention can be measured by conventional instruments in the art, such as monochromators and the like.
  • the term "about” as used in the present invention means that the temperature is up and down by a value not exceeding 2.5 ° C (indicating a value of ⁇ 2.5 ° C), preferably ⁇ 2.5 ° C, ⁇ 2 ° C or ⁇ 1 ° C.
  • the upper and lower values of the number do not exceed 2500 Lux (representing the value of ⁇ 2500 Lux), preferably the values are ⁇ 2500 Lux, ⁇ 2000 Lux, ⁇ 1500 Lux, ⁇ 1000 Lux, ⁇ 500 Lux, ⁇ 200 Lux, ⁇ 100 Lux;
  • the value is not more than 5 nm up and down (indicated as a value of ⁇ 5 nm), and the value is ⁇ 4 nm, ⁇ 3 nm or ⁇ 1 nm.
  • the light amplitude it means that the number is centered.
  • the variation is not more than 3 nm (indicating a value of ⁇ 3 nm), and preferably the value is ⁇ 2 nm or ⁇ 1 nm.
  • the reaction system of the chlorination reaction of the present invention preferably does not contain an external solvent and an initiator, and more preferably does not contain other components other than di(methyl)benzene and chlorine.
  • the chlorination of the present invention can be monitored at various stages by conventional sampling and detection methods, such as gas chromatography, to properly adjust the above parameters to save reaction time.
  • the description of the three-stage time aspect herein is not limiting, and the staged reaction time can be freely adjusted based on the chlorination progress monitoring results.
  • Chlorine gas velocity as described herein Not limited to a specific feed rate. When the slow, gradual and other terms are used to describe the rate of chlorine gas introduction, the meaning is not unclear. Since the rate of introduction of chlorine gas can be appropriately adjusted by those skilled in the art based on the reaction monitoring results.
  • the method of the present application produces a product with a high purity.
  • the purity obtained directly after the reaction is between about 70% and about 75%, between about 75% and about 80%, between about 80% and about 85%, and between about 85% and about 90%.
  • the purity is directly between about 90.0% and about 90.5%, between about 90.0% and about 91.0%.
  • the trichloromethyl substituted benzene obtained by the chlorination of the present invention can be further purified according to a conventional purification method such as recrystallization, rectification, molecular distillation or the like.
  • the molecular distillation method is preferred in the present invention.
  • the method of this step of the present application can be carried out in a continuous or batch manner, preferably in a continuous reaction mode.
  • Step 2 acyl chloride reaction
  • the trichloromethyl-substituted benzene obtained in the first step is further reacted to prepare a chloroformyl-substituted benzene, and a by-product hydrogen chloride is obtained.
  • the resulting chloroformyl-substituted benzene is optionally further purified or collected directly as a final product; the resulting by-product hydrogen chloride is recovered for use in step three.
  • the acid chlorination reaction of the present invention comprises the following steps:
  • the molar ratio of the trichloromethyl-substituted benzene to the corresponding aromatic acid in the step i) is a measurement value in which the chemical reaction is completely carried out, for example, the molar ratio of the bis(trichloromethyl)benzene to the phthalic acid is preferably 1:1.01 to 1.03.
  • the catalyst reacted in step i) is a Lewis acid such as aluminum trichloride, zinc chloride, ferric chloride or the like, preferably ferric chloride; in particular, in step i), trichloromethyl substituted benzene is reacted with water.
  • a small amount of the corresponding aromatic acid of the formula (X) a C 6 H 6-ab (COOH) b is also present.
  • the amount of the catalyst added in the step i) is preferably 0.2% to 0.3% by mass of the trichloromethyl-substituted benzene.
  • the chloroformyl-substituted benzene obtained by acid chlorination may also be further purified according to an optional purification step such as rectification, distillation, molecular distillation or recrystallization, and is preferably rectified in the present invention.
  • the method of this step of the present application can be carried out in a continuous or batch manner, preferably in a continuous reaction mode.
  • Step 3 catalytic oxidation of by-product hydrogen chloride
  • the hydrogen chloride is subjected to catalytic oxidation reaction (Deacon reaction) through a catalyst to prepare chlorine gas.
  • the by-product hydrogen chloride gas is first subjected to deep purification or adsorption for pre-purification to remove organic impurities, followed by catalytic oxidation.
  • hydrogen chloride gas can be purified by adsorption, and suitable adsorbent materials include, for example, activated carbon, alumina, titania, silica, iron oxide, silica gel, zeolite, and molecular sieves.
  • Step 3 of the present invention relates, in one aspect, to a method for catalytically oxidizing hydrogen chloride to produce chlorine gas, comprising the steps of:
  • the reactor provides a hydrogen chloride containing gas stream and/or an oxygen containing gas stream for oxidizing the hydrogen chloride containing gas stream for catalytic oxidation of hydrogen chloride;
  • step four The remainder of the product gas stream from the last reactor is provided to step four for separation.
  • the method of this step of the present application can be carried out in a continuous or batch manner, preferably in a continuous reaction mode.
  • the third step of the present invention relates to a method for catalytically oxidizing hydrogen chloride to produce chlorine gas, the method comprising:
  • step four The remainder of the product gas stream from the last reactor is provided to step four for separation.
  • the first first reactor in one or more reactors is provided with a hydrogen chloride containing gas stream and an oxygen containing gas stream for oxidizing the hydrogen chloride containing gas stream to the one
  • the downstream reactor in the plurality of reactors provides an oxygen-containing gas stream for oxidizing the hydrogen chloride-containing gas stream; the oxygen-containing gas stream supplied to each reactor for oxidizing the hydrogen chloride-containing gas stream is required as needed
  • the portion of the oxygen-containing gas stream which oxidizes the hydrogen chloride-containing gas stream is distributed between the reactors in an arbitrary ratio, preferably the desired number of oxygen-containing gas streams for oxidizing the hydrogen chloride-containing gas stream is equally distributed according to the number of reactors. For the corresponding number of copies.
  • step four The remainder of the product gas stream from the last reactor is provided to step four for separation.
  • the oxygen-containing gas stream entering each reactor has an oxygen content greater than the theoretical oxygen amount required to oxidize the hydrogen chloride-containing gas stream entering each reactor.
  • This particularly preferred embodiment can be carried out, for example, by providing a first reactor in the one or more reactors with an oxygen-containing gas stream for oxidizing a hydrogen chloride-containing gas stream and a hydrogen chloride-containing gas stream, A downstream reactor in the one or more reactors provides a hydrogen chloride-containing gas stream; a portion of the hydrogen chloride-containing gas stream is distributed between each reactor in a ratio of hydrogen chloride gas to be oxidized as desired.
  • the hydrogen chloride-containing gas stream to be oxidized is preferably evenly distributed to the corresponding number of parts in accordance with the number of reactors.
  • step four The remainder of the product gas stream from the last reactor is provided to step four for separation.
  • a portion of the product gas stream from the last reactor is preferably Returning to each of the reactors provided without separation; more preferably, before returning to each reactor feed port, mixing with the hydrogen chloride-containing gas stream and/or the oxygen-containing gas stream for oxidizing the hydrogen chloride-containing gas stream, The reactor is then introduced to carry out the catalytic oxidation reaction.
  • the process of the present invention can dilute the concentration of the feed gas to each reactor to prevent a violent reaction at the inlet of the reactor and avoid causing too many hot spots; After the mixing, the process of the invention increases the feed temperature of the feed gas of the feedstock and substantially eliminates the need to preheat the feed gas.
  • the returned product gas stream may be in each reactor at any ratio
  • the inter-distribution can be rationally distributed according to the operating conditions of the individual reactors, preferably by returning the returned product gas stream equally to the corresponding fractions according to the number of reactors and returning to each reactor separately.
  • the reactor described in step 3 of the present application is preferably an adiabatic reactor.
  • a heat exchanger can be connected between the reactors to remove the heat of reaction, that is, a heat exchanger is optionally present after each reactor.
  • the heat exchanger installed after the last reactor is a gas heat exchanger
  • the heat exchangers installed after the rest of the reactor may be heat exchangers well known to those skilled in the art, such as tube bundle heat exchangers, plate exchange Heaters, or gas heat exchangers, etc.
  • the present application preferably optimizes the remainder of the product gas stream (high temperature) after the catalytic oxidation reaction in step three (or all parts after the end of the third reaction, those skilled in the art can understand that the last part of the product gas stream may not return)
  • Separation is carried out after heat exchange by a gas heat exchanger, preferably a gas stream containing hydrogen chloride gas entering the first reactor and/or an oxygen-containing gas stream for oxidizing the hydrogen chloride-containing gas stream as a cooling medium in the gas Performing heat exchange in the heat exchanger; preferably, the heat exchanged hydrogen chloride-containing gas stream and/or the oxygen-containing gas stream for oxidizing the hydrogen chloride-containing gas stream is supplied to the first reactor and returned from the third A portion of the product gas stream exiting the stage reactor is combined and then passed to the first reactor for catalytic oxidation of hydrogen chloride.
  • the product gas stream is reduced in temperature after heat exchange.
  • Hydrogen chloride-containing gas used as a cooling medium
  • the stream and/or the oxygen-containing gas stream for oxidizing the hydrogen chloride-containing gas stream is heated by heat exchange, and then the heat-exchanged hydrogen chloride-containing gas stream and/or the oxygen-containing gas used to oxidize the hydrogen chloride-containing gas stream
  • the stream is supplied to the first reactor for catalytic oxidation of hydrogen chloride; preferably the heat exchanged hydrogen chloride containing gas stream and/or the oxygen containing gas stream for oxidizing the hydrogen chloride containing gas stream is provided to the first reactor It is mixed with a portion of the product gas stream that is returned from the third stage reactor and then passed to the first reactor for catalytic oxidation of hydrogen chloride.
  • a chlorine-containing, oxygen-containing, and/or hydrogen-containing hydrogen gas stream by dehydrating and removing (partially residual) hydrogen chloride in part or all of the product gas stream in step three.
  • the gas stream and the oxygen-containing gas stream provide a chlorine-containing gas stream.
  • the present application can provide the (unreacted residual) hydrogen chloride and/or oxygen separated from the product gas stream in step four to the catalytic oxidation reaction of step three again.
  • the hydrogen chloride (or vaporized hydrochloric acid) and/or oxygen separated in step four may also be returned to one or more of the reactors in step three.
  • the portion of the product gas stream (returned product gas stream) that is returned to the reactor without separation and the remainder of the product gas stream (remaining product gas stream portion) is preferably selected in step three.
  • the volume ratio is from 0.25:0.75 to 0.75:0.25, preferably from 0.35:0.65 to 0.45:0.55.
  • the hydrogen chloride containing gas stream (according to pure chlorine)
  • the hydrogenation calculation) and the oxygen-containing stream for the oxidation of the hydrogen chloride gas stream (calculated as pure oxygen) have a feed volume ratio of from 1:2 to 5:1, preferably from 1:1.2 to 3.5:1, more preferably 1:1 to 3:1.
  • the feed volume ratio of the hydrogen chloride-containing gas stream (calculated as pure hydrogen chloride) to the oxygen-containing gas stream (calculated as pure oxygen) for the oxidation of the hydrogen chloride gas stream is 2 : 1 to 5:1.
  • the feed volume of the hydrogen chloride-containing gas stream (calculated as pure hydrogen chloride) and the oxygen-containing stream (calculated as pure oxygen) for the oxidation of the hydrogen chloride gas stream The ratio is from 1:2 to 2:1, preferably from 0.9:1.1 to 1.1:0.9.
  • the pressure in the reactor is from 0.1 to 1 MPa.
  • the feed gas temperature of the reactor is from 250 to 450 ° C, preferably from 300 to 380 ° C.
  • the catalyst described in the third step of the present application is a conventional catalyst capable of oxidizing hydrogen chloride gas and oxygen to form chlorine gas and water.
  • Suitable catalysts include copper compounds or/and ruthenium compounds, preferably copper compounds or/and ruthenium compounds supported on supported alumina, or titania or the like.
  • alumina supported with copper chloride or barium chloride is preferably a barium compound.
  • Suitable catalysts described herein may also contain other promoters, such as metals, metals such as gold, palladium, platinum, rhodium, iridium, nickel or chromium, alkali metals, alkaline earth metals and rare earth metals.
  • Suitable catalysts can have different shapes, such as rings, cylinders or spheres, and the like, preferably a suitable catalyst has similar outer dimensions.
  • the reactor in the third step of the present application is a conventional reaction device, such as a fixed bed or a fluidized bed reaction.
  • a fixed bed reactor is preferred in which the desired catalyst can be charged.
  • the reactor described in the present application may be a reactor of any material that meets the requirements of the reaction, preferably a reactor of pure nickel or nickel alloy or quartz. If a plurality of reactors are selected, they may be connected in series or in parallel, preferably in series, so that the oxidation reaction of hydrogen chloride can be carried out in multiple stages.
  • the present application preferably employs 2, 3, 4, 5, 6, 7, 8, 9, 10, more preferably 3 or 4 reactors.
  • 2, 3, 4, 5, 6, 7, 8, 9, 10, particularly preferably 3 or 4 in series adiabatic reactors are preferably provided.
  • the reactors connected in parallel and connected in series can also be combined with each other.
  • the process according to the invention particularly preferably has a reactor which is only connected in series. If it is preferred to use reactors connected in parallel, in particular up to five, preferably three, particularly preferably up to two production lines (optionally comprising reactors consisting of reactors connected in series) are connected in parallel.
  • the methods described herein can operate, for example, up to 60 reactors.
  • the method of this step of the present application can be carried out in a continuous or batch manner, preferably in a continuous reaction mode.
  • the hydrogen chloride-containing gas stream described herein includes a fresh hydrogen chloride-containing gas stream and a gas stream comprising hydrochloric acid recovered by the process of the present invention or hydrochloric acid recovered by gasification.
  • the fresh hydrogen chloride-containing gas stream may also be a hydrogen chloride-containing gas stream in the form of by-products from the related industries such as the production of isocyanates, the production of acid chlorides, the chlorination of aromatic compounds, and the like.
  • a hydrogen chloride-containing gas stream in the form of a by-product from steps 1 and 2 of the present invention is preferred.
  • the hydrogen chloride-containing gas stream in the form of a by-product may be a hydrogen chloride-containing gas stream in the form of a preliminary treated by-product or a hydrogen chloride-containing gas stream in the form of a by-product directly from the related industry without any treatment.
  • the hydrogen chloride-containing gas stream in the form of by-products may contain, depending on the source, little or no other impurity gases derived from the relevant industries that have no effect on the catalytic oxidation of hydrogen chloride. The amount of other impurity gases is determined by the nature of the production in the relevant industry. Those skilled in the art will appreciate that so-called exhaust hydrogen chloride produced in the relevant industries may be an appropriate raw material for the present application.
  • the unreacted hydrogen chloride-containing gas stream described herein refers to a hydrogen chloride-containing gas stream that has not undergone a catalytic oxidation reaction by the reactor described herein.
  • the oxygen-containing gas stream described herein includes a fresh oxygen-containing gas stream and an oxygen-containing gas stream recovered by the process of the present invention.
  • the fresh oxygen-containing gas stream can be pure oxygen or other oxygen-containing gas (e.g., air).
  • the product gas stream as referred to herein refers to a mixed gas comprising hydrogen chloride, oxygen, water vapor and chlorine obtained from a catalytic oxidation reaction from a reactor.
  • the product gas stream returned by the present invention is a mixed gas from the last reactor.
  • Step 4 separating the gas stream from step three), separating the product comprising the chlorine, oxygen and/or hydrogen chloride containing gas from the product gas stream in step three above.
  • the separation of the chlorine-containing, oxygen-containing, and/or hydrogen-containing hydrogen gas stream obtained in the fourth step of the present invention includes the following steps:
  • condensation condensation treatment of the product gas stream from step three; product from step three The water in the gas stream, together with a portion of the unreacted hydrogen chloride, is coagulated and precipitated as an aqueous solution of hydrochloric acid;
  • Deep dehydration deep dehydration of the gas stream condensed in step a, including deep dehydration by means of concentrated sulfuric acid, molecular sieve, or by techniques such as temperature swing adsorption and pressure swing adsorption to remove residual moisture and reduce Corrosiveness of the gas stream;
  • step b Adsorption: The gas stream after the deep dehydration treatment in step b is adsorbed by the adsorbent to separate chlorine gas and oxygen gas.
  • the adsorption may be selected from adsorbents capable of adsorbing a large amount of oxygen and only a small amount of chlorine gas, such as carbon molecular sieves, silica gel, etc., to remove and remove oxygen by adsorption; after the adsorption treatment of the adsorbent, the main component is chlorine.
  • a chlorine gas stream optionally containing a small amount of hydrogen chloride; the oxygen adsorbed to the adsorbent after the adsorbent is adsorbed and then desorbed to obtain a separated oxygen-containing gas stream; the desorbed adsorbent can continue in the step In c, it is used for adsorption separation to remove oxygen.
  • the adsorption can also select an adsorbent capable of adsorbing a large amount of chlorine gas and adsorbing only a small amount of oxygen, such as fine pore silica gel, activated carbon, etc., to remove and remove chlorine gas by adsorption, and the main component is obtained after adsorption treatment by the above adsorbent.
  • step c further comprising d, liquefying: liquefying the chlorine-containing gas stream obtained in step c, and separating the hydrogen chloride-containing gas stream and the liquefied chlorine-containing gas stream.
  • the condensation conditions in the step a are: a temperature of -5 to 5 ° C and a pressure of 0.05 to 10 MPa.
  • the drying in the step b is preferably carried out by a temperature swing adsorption drying or a pressure swing adsorption drying process, and in the temperature swing adsorption drying process, a composite adsorbent layer of two adsorbents is preferably used.
  • the adsorbent is an alumina dehydrating agent placed on the upper part of the adsorption tower, and the other is a dehydrated and dried adsorbent placed in the lower part of the adsorption tower, and the volume ratio of the upper alumina dehydrating agent and the lower deep dehydrated adsorbent is 20 ⁇ . 80%: 80% to 20%.
  • a composite adsorbent layer of two adsorbents is preferably used, one adsorbent is an alumina dehydrating agent placed on the upper part of the adsorption tower, and the other is a dehydrated and dried adsorption placed in the lower part of the adsorption tower.
  • the volume ratio of the upper alumina dehydrating agent to the lower dehydrated adsorbent is 20-80%:80%-20%.
  • the variable temperature adsorption drying process described in the step b is: passing the gas stream condensed through the step a from the bottom to the composite adsorbent layer, and the gas stream leaves the temperature swing adsorption drying device to reach the drying target; during the temperature-temperature adsorption drying process
  • the adsorption pressure is 0.30 to 0.80 MPa, and the adsorption temperature is 20 to 50 °C.
  • the temperature swing adsorption drying process comprises an alternating process of adsorption and regeneration operations wherein the alternating processes of adsorption and regeneration are achieved by conventional means including pressure reduction, displacement, temperature rise and cooling steps.
  • the regeneration operation includes a desorption and dehydration process.
  • the desorption pressure of the regeneration operation is 0.01 to 0.005 MPa, and the desorption temperature of the regeneration operation is 110 to 180 ° C; the dehydration process for the regeneration operation uses a carrier gas (raw material gas or nitrogen gas) at a temperature of 50 to 180 ° C, and the raw material gas is used as a carrier gas.
  • the raw material gas is dried by the pre-drying tower, heated by the steam heater, and then enters the adsorption drying tower which needs to be heated and regenerated and dehydrated. After the aqueous carrier gas is discharged from the adsorption tower, it is cooled, condensed, separated and returned to the raw material gas system for recovery. use.
  • the pressure swing adsorption drying process in step b comprises an alternating process of adsorption and desorption processes, wherein: The adsorption pressure is 0.40-0.80 MPa, the desorption pressure is 0.02--0.07 MPa, and the adsorption temperature is normal temperature; the alternating process of the adsorption and desorption processes is carried out according to a conventional setting (including pressure equalization, flushing replacement, vacuum suction, etc.);
  • the apparatus required for the pressure swing adsorption drying process is usually set as a four-column process. In this process, the dry process product gas stream is used for flushing replacement, and the flushing and vacuum suctioning tail gas are sent to the dehydrochlorination after cooling and dehydration.
  • the product gas logistics system is recycled.
  • the adsorbent for drying the molecular sieve in step b is zeolite molecular sieve or silica gel.
  • the adsorption in the step c is preferably a variable temperature pressure swing adsorption technology, comprising an adsorption and desorption process, wherein: the adsorption pressure is 0.20-0.7 MPa, and the temperature in the adsorption stage is gradually lowered from 40 to 70 ° C to 20 to 35 ° C;
  • the desorption pressure is -0.07 MPa, the desorption temperature is 40-70 ° C;
  • the gas stream used as a raw material is introduced at a temperature of less than 40 ° C during adsorption, and the adsorption and cooling are started; and the hot chlorine gas replacement system of more than 50 ° C is introduced before the desorption regeneration.
  • the gas, and the temperature rise promotes desorption.
  • the hot chlorine gas is stopped and the vacuum desorption is started. After the desorption regeneration is completed, the replacement before the adsorption is started by using oxygen; the exhaust gas of the hot chlorine gas replacement tail gas and the oxygen replacement is returned to the raw material gas system.
  • step d means that when the ratio of hydrogen chloride and oxygen participating in the catalytic oxidation reaction is appropriately controlled (for example, when the ratio of pure hydrogen chloride to pure oxygen is 0.5:1 to 1:0.5 ), the residual The unreacted hydrogen chloride is substantially absorbed by the water formed by the reaction during the condensation process.
  • the amount of hydrogen chloride contained in the chlorine gas obtained after the treatment in the step c is small, the chlorine gas is not involved in the chlorination reaction, and the chlorine gas is not required to be further processed.
  • the liquefaction conditions in the step d are: a temperature of -20 to 20 ° C and a pressure of 0.05 to 10 MPa.
  • Step 5 (recycling the separated product), introducing the chlorine-containing gas stream separated in the above step 4 as a raw material into the chlorination reaction including the first step; and using the hydrogen chloride-containing and/or oxygen-containing gas stream separated in the above step 4 as The raw material is introduced into the reaction including the catalytic oxidation by-product hydrogen chloride in the third step.
  • the chlorine-containing gas stream obtained by the separation of the fourth step can also be referred to other independent chlorination reactions.
  • the purity of the chlorine gas in the chlorine-containing gas stream obtained by the third catalytic oxidation method of the invention can reach 99.6% (vol%) or more, and can meet the quality requirement of the photochlorination reaction for the raw material chlorine gas.
  • the closed loop of the chlorine resource (or chlorine element or chlorine atom) of the present invention means that the process of the present invention allows the chlorine element to be recycled in the process of the present invention by rational treatment of the by-product hydrogen chloride.
  • the product 1,3-bis(trichloromethyl)benzene, 1,4-bis(trichloromethyl)benzene, 1,3-bis(chloroformyl)benzene, 1,4-di The purity of chloroformyl)benzene, p-chloro(trichloromethyl)benzene, and trichloromethylbenzene was quantitatively determined using a gas chromatograph.
  • 1,3-di(methyl)benzene is continuously added from the top of the first column at a rate of 95 kg/h, and the first tower temperature is controlled at 80.
  • °C ⁇ 120 °C the incident light center peak wavelength is 460nm
  • the average illuminance in the tower is 20,000 ⁇ 39,000 Lux
  • the chlorine gas is passed through the bottom of the tower at a flow rate of 135kg / h for continuous chlorination
  • the temperature is controlled at 135 to 145 ° C
  • the chlorine gas is introduced into the second tower at a flow rate of 128 kg/h
  • the reaction liquid of the tower overflows from the bottom of the tower to the third tower.
  • the peak wavelength of the incident light center of the third tower is 586 nm, the average illuminance is 60,000 to 86,000 Lux, the temperature is controlled at 170 to 180 ° C, and the flow rate is 148 kg/h to the third tower.
  • the chlorine gas is introduced into the reaction system, and the total amount of chlorine gas introduced into the reaction system consisting of three columns is 411 kg/h.
  • the reaction mixture obtained from the outlet of the third column is 1,3-bis(trichloromethyl)benzene, which is purified by one-time rectification to obtain purified 1,3-bis(trichloromethyl)benzene.
  • the by-product hydrogen chloride gas generated in the photochlorination reaction is collected.
  • the purified 1,3-bis(trichloromethyl)benzene obtained in the first step is added to the batching kettle with temperature measurement, condensing reflux and stirring device, and two or more batching kettles can be set to realize continuous feeding. .
  • Step three catalytic oxidation by-product hydrogen chloride
  • the hydrogen chloride as a by-product in the first step and the second step is collected, and the catalytic oxidation reaction (Deacon reaction) is carried out through the catalyst to prepare chlorine gas, which specifically comprises the following steps:
  • the by-product hydrogen chloride gas is purified by adsorption to remove organic impurities.
  • the hydrogen chloride-containing gas stream and the oxygen-containing gas stream entering the first-stage reactor are first mixed, preheated, and passed to the first-stage reactor.
  • Step 4 separating the gas stream from step three
  • step three Separating the chlorine, oxygen or hydrogen chloride containing gas stream from the product gas stream in step three above comprises the following steps:
  • step three The product gas stream from step three is subjected to condensation treatment at a condensation temperature of -5 to 5 ° C and a pressure of 0.05 to 10 MPa. , water together with a portion of unreacted hydrogen chloride, coagulated as an aqueous solution of hydrochloric acid;
  • step b Deep dehydration: The gas stream condensed through step a is deeply dehydrated by concentrated sulfuric acid.
  • step b Adsorption: the gas stream after the deep dehydration treatment in step b is passed through the adsorbent silica gel to remove and remove oxygen by adsorption, and the variable temperature pressure swing adsorption technology is adopted, wherein: the adsorption pressure is 0.5 MPa, and the temperature in the adsorption phase is gradually decreased from 60 ° C to 60 ° C. 25 ° C; decompression desorption pressure of -0.07 MPa, desorption temperature of 50 ° C, desorption to obtain a separate oxygen-containing gas stream.
  • the residual gas after adsorption is a chlorine-containing gas stream whose main component is chlorine.
  • liquefaction liquefying the chlorine-containing gas stream obtained in step c, the liquefaction temperature is -20 to 20 ° C, the pressure is 0.05 to 10 MPa, and the chlorine-containing gas stream after the hydrogen chloride gas stream and the liquefaction treatment are separated.
  • Step 5 recycling the separated substance
  • the chlorine-containing gas stream separated in the above step four is introduced as a raw material into the chlorination reaction including the first step.
  • the oxygen-containing and hydrogen-containing hydrogen gas stream separated in the above step 4 is again introduced as a raw material into the hydrogen chloride catalytic oxidation reaction including the third step.
  • the specific process conditions are shown in Table 2.
  • Embodiment 2 the specific operation process is as follows:
  • the chlorination reaction operation was the same as that in Example 1. The difference is that the raw material is 1,4-bis(methyl)benzene, the velocity of entering the first column is 100 kg/h, the peak wavelength of the incident light center of the first tower is 460 nm, the average illuminance in the tower is 20,000 to 39,000 Lux; the second tower is incident.
  • the peak wavelength of the optical center is 505 nm, the average illuminance is 40,000 to 61,000 Lux, the temperature of the tower is controlled at 135 to 145 ° C; the peak wavelength of the incident light center of the third tower is 586 nm, the average illuminance is 60,000 to 86,000 Lux, and the temperature is controlled at 170 to 180 °C.
  • the reaction mixture obtained from the outlet of the third column is 1,4-bis(trichloromethyl)benzene, which is purified by one-time rectification to obtain purified 1,4-bis(trichloromethyl)benzene.
  • the by-product hydrogen chloride gas generated in the photochlorination reaction is collected.
  • the acid chloride reaction operation is the same as that in the second step of the first embodiment, and 1,4-bis(chloroformyl)benzene is obtained from the outlet of the second stage reactor, and after one rectification, the purified 1,4-di(chloroform) is obtained.
  • Acyl) benzene is obtained from the outlet of the second stage reactor, and after one rectification, the purified 1,4-di(chloroform) is obtained.
  • Step three catalytic oxidation by-product hydrogen chloride
  • the by-product hydrogen chloride gas is purified by adsorption to remove organic impurities.
  • the oxygen-containing gas stream and the hydrogen chloride-containing gas stream entering the first-stage reactor are first mixed, preheated, and passed to the first-stage reactor.
  • Step 4 separating the gas stream from step three
  • step three Separating the chlorine, oxygen and hydrogen chloride containing gas streams from the product gas stream in step three above comprises the following steps:
  • condensation the product gas stream from step three is condensed, the condensation temperature is -5 ⁇ 5 ° C, the pressure is 0.05 ⁇ 10MPa, water together with some unreacted hydrogen chloride, coagulation and precipitation in the form of aqueous hydrochloric acid;
  • Deep dehydration deep dehydration of the gas stream condensed in step a, drying by pressure swing adsorption technology, and using a composite adsorbent layer of two adsorbents, one adsorbent is placed in the upper part of the adsorption tower for oxidation
  • the aluminum dehydrating agent the other is a zeolite molecular sieve adsorbent which is deeply dehydrated and dried in the lower part of the adsorption tower, and the volume ratio of the upper alumina dehydrating agent and the lower deep dehydrated adsorbent is 40%:60%.
  • the adsorption pressure is 0.40 MPa
  • the desorption pressure is 0.02 MPa
  • the adsorption temperature is normal temperature.
  • step b Adsorption: the gas stream after the deep dehydration treatment in step b is passed through the adsorbent carbon molecular sieve to remove and remove oxygen by adsorption, and adopts a variable temperature pressure swing adsorption technology, including adsorption and desorption processes, wherein: the adsorption pressure is 0.20 MPa, and the adsorption phase is The temperature was gradually lowered from 40 ° C to 20 ° C; the vacuum desorption pressure was -0.07 MPa, the desorption temperature was 40 ° C, and the separated oxygen-containing gas stream was desorbed.
  • the residual gas after adsorption is a chlorine-containing gas stream whose main component is chlorine.
  • Step 5 recycling the separated substance
  • the chlorine-containing gas stream separated in the above step four is introduced as a raw material into the chlorination reaction including the first step.
  • the oxygen-containing gas stream separated in the above step 4 is again introduced as a raw material into the hydrogen chloride catalytic oxidation reaction including the third step.
  • the specific process conditions are shown in Table 2.
  • Example 3 was operated according to the specific procedure of Example 2 above, wherein the methyl aromatic hydrocarbon material was p-chlorotoluene, and reacted with chlorine gas to obtain p-chloro(trichloromethyl)benzene, and the corresponding aromatic acid was p-chlorobenzoic acid.
  • Example 4 was operated according to the specific procedure of Example 2 above, wherein the methyl aromatic hydrocarbon material was methylbenzene, which was reacted with chlorine gas to obtain trichloromethylbenzene, and the corresponding aromatic acid was benzoic acid.
  • Example 5 was carried out according to the specific procedure of Example 1 above, wherein the methyl aromatic hydrocarbon material was mesitylene, and reacted with chlorine gas to obtain tris(trichloromethyl)benzene, and the corresponding aromatic acid was trimesic acid.
  • the illuminance was maintained to 60,000 Lux, and after the system temperature was raised to 160 ° C, 155.2 kg of chlorine gas was introduced, and the third reaction stage took a total of 16 hours and 35 minutes.
  • the total amount of chlorine consumed in the reaction was 427 kg.
  • the reaction mixture after completion of the reaction was crude 1,3-bis(trichloromethyl)benzene, which was purified by one-time distillation to obtain 255 kg of 1,3-bis(trichloromethyl)benzene having a purity of 99.42%.
  • the gas generated in the chlorination reaction was collected to obtain 122 m 3 of by-product hydrogen chloride.
  • the obtained product was subjected to rectification to obtain purified 1,3-bis(chloroformyl)benzene having a purity of 99.97%.
  • the gas generated in the acid chloride reaction was collected to obtain 34 m 3 of by-product hydrogen chloride.
  • Step three catalytic oxidation by-product hydrogen chloride
  • the two by-product hydrogen chloride gases produced in this example were continuously compressed into a catalytic oxidation system in the same manner as in the third step of Example 2.
  • Step 4 separating the gas stream from step three
  • Separating the chlorine-containing, oxygen-containing gas stream from the product gas stream in the above step 3 includes the following steps:
  • step three The product gas stream from step three is subjected to condensation treatment at a condensation temperature of -5 to 5 ° C and a pressure of 0.05 to 10 MPa. , water together with a portion of unreacted hydrogen chloride, coagulated as an aqueous solution of hydrochloric acid;
  • Deep dehydration deep dehydration of the gas stream condensed in step a, drying by variable temperature adsorption technology, and using a composite adsorbent layer of two adsorbents, one adsorbent is alumina placed on the upper part of the adsorption tower
  • the dehydrating agent the other is a zeolite molecular sieve adsorbent which is deeply dehydrated and dried in the lower part of the adsorption tower, and the volume ratio of the upper alumina dehydrating agent and the lower deep dehydrated adsorbent is 30%:70%.
  • the adsorption pressure during the temperature-dependent adsorption drying process is 0.70 MPa, and the adsorption temperature is 30 ° C; the regeneration operation includes a desorption and dehydration process.
  • the desorption pressure of the regeneration operation was 0.009 MPa, the desorption temperature of the regeneration operation was 160 ° C, and the dehydration process of the regeneration operation employed a carrier gas having a temperature of 180 ° C.
  • adsorption the gas stream after the deep dehydration treatment in step b is passed through the adsorbent carbon molecular sieve to Adsorption separation and removal of oxygen, using variable temperature pressure swing adsorption technology, including adsorption and desorption process, wherein: adsorption pressure is 0.5MPa, the adsorption phase temperature is gradually reduced from 60 ° C to 25 ° C; decompression desorption pressure is -0.07MPa, desorption temperature At 50 ° C, a separate oxygen-containing gas stream was obtained by desorption.
  • the residual gas after adsorption is a chlorine-containing gas stream whose main component is chlorine.
  • the chlorine-containing gas stream obtained in the step c is subjected to liquefaction treatment, the liquefaction temperature is -20 to 20 ° C, and the pressure is 0.05 to 10 MPa, and the chlorine-containing gas stream and the liquefied chlorine-containing gas stream are separated.
  • the amount of chlorine gas obtained was 195 Kg, the purity of chlorine gas was 99.97 (v.%), and the amount of recovered hydrogen chloride obtained after separation was 18 m 3 ; the amount of recovered oxygen obtained after separation was 30 m 3 .
  • Step 5 recycling the separated substance
  • the chlorine-containing gas stream separated in the above step 4 is introduced as a raw material into the chlorination reaction including the first step, and the obtained crude product is subjected to one-step rectification purification to obtain 1,3-bis(trichloromethane) having a purity of 99.4%. Further, the obtained purified 1,3-bis(trichloromethyl)benzene is further introduced into the acid chloride reaction comprising the second step to obtain a 1,3-bis(chloroformyl) group having a purity of 99.96%. benzene.
  • the reaction results are shown in Table 1.
  • Table 1 shows the entry and output quantities of the main materials of Examples 1-5;
  • Table 2 shows the operating conditions of the step dioxide units of the respective examples.

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Abstract

本发明属于化工技术领域，涉及一种制备氯甲酰基取代苯的清洁工艺。本发明的工艺通过氧化氯化反应和酰氯化反应的尾气氯化氢，并将所获得的氧化产物氯气回用到氯化反应中。本发明是一种制备聚合级氯甲酰基取代苯的清洁工艺。

Description

制备氯甲酰基取代苯的清洁工艺 技术领域
本发明属于化工技术领域，涉及一种制备氯甲酰基取代苯的清洁工艺。本发明的工艺通过氧化氯化反应和酰氯化反应的尾气氯化氢，并将所获得的氧化产物氯气回用到氯化反应中。本发明是一种制备聚合级氯甲酰基取代苯的清洁工艺。
背景技术
现有技术中，氯甲酰基取代苯的制备方法主要有光氯化法(参见DE31 468 68、JP47-130931)、氯化亚砜法、三氯化磷法、五氯化磷法和光气法等。氯化亚砜法应用最为普遍(例如，参见，CN102516060A、CN102344362A)，但需要采用99.99％高纯度苯二甲酸作为原料，这使得该方法成本很高。而且，这些方法都存在产生氯化氢、二氧化硫、二氧化碳、亚磷酸等对环境不友好副产物问题。这些副产物给产品的后续处理带来不便，且极易造成环境污染。
光氯化法可以采用甲基芳烃类化合物作为原料，但是副产物氯化氢的量是巨大的。如何处理大量的氯化氢，已经成为一个亟待解决的问题。目前工业上实际采用的主要处理措施是用水吸收氯化氢后制成质低价廉的盐酸出售；由于盐酸价格低廉且市场需求容量有限，将氯化氢制成盐酸实际上已经成为一种负担而不是变废为宝。还有一些采用的处理措施是用碱中和氯化氢后直接排放；然而随着环保法律法规日趋完善，各种排放方式的环保标准已经非常严格。
将副产物氯化氢直接制成氯气的方法，不仅能实现氯元素的闭路循环，还能实现反应过程的零排放，大大提升行业的节能减排水平、降低成本并且消除对环境的污染。到目前为止，氯化氢制备氯气的方法主要可以分为三类：电解法，直接氧化法和催化氧化法。但是，电解法工艺的能源消耗太大，且离子膜也需要经常更换，成本非常高，每吨氯气回收成本＞4000元；直接氧化法的收率低，不可工业化；与电解法、直接氧化法相比，催化氧化法尤其是经由Deacon反应的催化氧化法最具工业化潜力。
在氯化产业许多生产采用工业氯气的要求为≥99.6％(体积％)。因此从Deacon反应获得的氯气，还需解决反应所得混合气体的分离问题，才能获得可以循环使用的高纯度氯气。特别地，现有技术中，为了从氯化氢氧化混合气体中回收氯气，对氯化氢一般采用水吸收的分离方法，例如CN102502498A，US2008/0159948A1，这样又会进一步造成大量稀盐酸的生成，需要进一步处理。
综上可以看出，现有技术中需要一种能实现氯资源闭路循环的成本低、质量优、无污染的生产氯甲酰基取代苯的清洁工艺。所述清洁工艺的关键在于氯化氢氧化过程以及产物气体物流的分离过程。解决了所述关键过程后就可以得到高纯度氯气，实现氯气的回用。氯甲酰基取代苯的清洁生产工艺是实现相关化工产业例如芳纶产业工业化的关键。
发明概述
本发明目的是提供一种制备氯甲酰基取代苯的清洁工艺。本发明工艺解决了现有产业存在的问题，实现了氯资源的闭路循环，从源头上消除了由副产物引起的污染，同时该工艺所得产品成本低、品质高。
本发明是通过以下技术方案实现的：首先将结构式为(X)aC6H6-a-b(CH3)b的甲基芳烃或者该化合物的炕基侧链氯化物和氯气(例如在光照条件下)反应制备三氯甲基取代苯；再将所产生的三氯甲基取代苯进一步反应制备氯甲酰基取代苯；将所产生的HCl气体经Deacon反应催化氧化制成氯气后再用于所述甲基芳烃氯化制备三氯甲基取代苯。整个工艺过程的代表性反应式如下：
(X)aC6H6-a-b(CH3)b+3bCl2＝(X)aC6H6-a-b(CCl3)b+3bHCl
(X)aC6H6-a-b(CCl3)b+(X)aC6H6-a-b(COOH)b＝2(X)aC6H6-a-b(COCl)b+bHCl
或(X)aC6H6-a-b(CCl3)b+bH2O＝(X)aC6H6-a-b(COCl)b+2bHCl
4bHCl+bO2＝2bCl2+2bH2O。
出于简便的需要，上述反应式中仅示出了所述甲基芳烃。本领域技术人员能理解所述甲基芳烃的烷基侧链氯化物将依照类似的反应进行。
将氧化获得的氯气作为原料再引入到氯化反应，综合得到本发明二(氯甲酰基)苯制备的总反应方程式为：
(X)aC6H6-a-b(CH3)b+bCl2+(X)aC6H6-a-b(COOH)b+bO2＝2(X)aC6H6-a-b(COCl)b+2bH2O。
本领域技术人员能够理解使用水与三氯甲基取代苯制备酰氯的反应也具有类似的总反应方程式。
所述(X)aC6H6-a-b(CH3)b为甲基芳烃类化合物(该化合物的烷基侧链氯化物同样适用于本发明)，(X)aC6H6-a-b(CCl3)b为三氯甲基取代苯，(X)aC6H6-a-b(COOH)b为相应的芳香酸，(X)aC6H6-a-b(COCl)b为氯甲酰基取代苯，本申请所述的上述化合物分子式中，X为氯或溴或氟原子，a为选自0、1、2、3、4或5的整数，b为选自1、2、3或4的整数，且a+b≤6。本申请所述相应芳香酸是指所述芳香酸母核上的取代基与所述甲基芳烃母核上的取代基 处于相同或相应的取代位置；所述芳香酸母核上的取代基与所述甲基芳烃母核上的取代基也可以是相同的。
本申请所述甲基芳烃类化合物的烷基侧链氯化物，是指所述芳烃类化合物中炕基上的氢原子未全部被氯原子取代的化合物；本申请所述光氯化反应的目标产物即三氯甲基取代苯是指所述芳烃类化合物中烷基上的氢原子全部被氯原子取代的产物。
一种制备氯甲酰基取代苯的清洁工艺，包含下述步骤：
步骤一(氯化反应)、将所述结构式为(X)aC6H6-a-b(CH3)b的甲基芳烃或其烷基侧链氯化物和氯气(例如在光照条件下)反应制备三氯甲基取代苯，并得到副产物氯化氢；
步骤二(酰氯化反应)、将步骤一制备得到的三氯甲基取代苯与所述结构式为(X)aC6H6-a-b(COOH)b的相应芳香酸或水进一步反应制备氯甲酰基取代苯，并得到副产物氯化氢；
步骤三(催化氧化副产物氯化氢)、收集上述步骤一和步骤二中的副产物氯化氢，经催化剂进行催化氧化反应(Deacon反应)，制备氯气；
步骤四(分离来自步骤三的气体物流)、从上述步骤三中的产物气体物流中分离获得含氯气、含氧气和/或含氯化氢气体物流。
步骤五(循环利用分离物)、将上述步骤四分离所得的含氯气体物流作为原料引入到包括步骤一的氯化反应中；
任选地，将上述步骤四分离所得的含氯化氢和/或含氧气体物流作为原料引入到步骤三中的催化氧化副产物氯化氢的反应中。
本发明所述制备氯甲酰基取代苯的清洁工艺中，在氯化反应后可进一步存在或不存在三氯甲基取代苯的纯化步骤；在酰氯化反应后可进一步存在或 不存在氯甲酰基取代苯的纯化步骤。
本发明工艺，除纯化过程和反应过程中的正常损耗外无任何遗弃，是一个完全绿色的化学工艺过程。本发明工艺实现了氯甲酰基取代苯，特别是芳纶制备原料聚合级二(氯甲酰基)苯的清洁生产，对于以低成本生产高性能芳纶纤维，具有重要的经济价值和社会效益。
本发明的有益效果：
1)本发明将氯甲酰基取代苯，例如二(氯甲酰基)苯生产过程中，氯化与酰氯化步骤中所产生的副产物氯化氢，进行进一步催化氧化，制取氯气，并将所获得的氯气循环进行氯化反应，实现了氯元素的闭路循环，降低生产成本并减少了环境污染。
2)本发明步骤三催化氧化步骤是实现氯资源循环利用的关键与核心，在该步骤中，通过将氯化氢催化氧化反应所得产物气体物流部分不经分离直接循环的技术，实现对Deacon反应过程中热量的分散作用，延长催化剂寿命，同时利用循环产物气体物流其自身携带的热量减少了对含氯化氢原料气预热的燃料费用，进一步节省了工业化的成本。
3)在步骤四中，本发明采用包含冷凝、干燥、吸附步骤的分离方法来分离来自步骤三的产物气体物流，所述方法中任选进一步包括液化分离步骤。本发明分离方法因不存在水洗步骤而不会产生大量稀盐酸。特别地，当经冷凝、干燥、吸附分离去除氧气后的氯气中氯化氢浓度较低时，无需进行进一步的液化分离处理步骤。因为含有少量氯化氢气体的氯气直接循环回用到例如步骤一的氯化步骤，存在少量氯化氢并不影响所述氯气与甲基芳烃反应生产三氯甲基取代苯。
本发明步骤四中的产物气体物流的分离方法具有工艺流程简单、环境友 好、能耗低、分离效率高、成本低等优点，分离回收的含氯气体物流中氯气的纯度达≥99.6％(体积％)，这样的含氯气体物流能满足光氯化反应对氯气原料气的质量要求。
4)本发明工艺中，除氯资源实现闭路循环以外，产品生产过程中产生的其它物质也可进行循环利用，从而实现清洁生产。例如，步骤三催化氧化步骤中未反应完全的氯化氢和/或氧气，经分离后会再次进行所述催化氧化反应。
5)本发明工艺可以获得聚合级氯甲酰基取代苯，例如二(氯甲酰基)苯，且生产成本比传统的工艺降低30％以上。
附图说明
图1为制备二(氯甲酰基)苯的清洁工艺流程图。
发明详述
现有技术中氯甲酰基取代苯制备方法均存在副产物多、处理困难、产品收率低、环境污染严重等问题。要得到高品质氯甲酰基取代苯，需要付出巨大的经济成本和环境成本。
本专利发明人经过大量实验研究，创造性地将甲基芳烃的氯化、氯甲酰基取代苯的合成与氯化氢氧化制备氯气整合成一个完整的工艺(同时通过改进氯化氢氧化工艺以及混合气体的分离过程)，其中将氯化和酰氯化过程中产生的大量氯化氢通过催化氧化制成氯气并将所得氯气再次引入到氯化工艺中，实现了氯元素的循环利用。整个工艺是一个清洁生产过程。
整个清洁生产过程按照下述步骤进行：
步骤一(氯化反应)、将所述结构式为(X)aC6H6-a-b(CH3)b的甲基芳烃或 者该化合物的烷基侧链氯化物和氯气(例如在光照条件下)反应制备三氯甲基取代苯，并得到副产物氯化氢，其中X为氯或溴或氟原子，a为选自0、1、2、3、4或5的整数，b为选自1、2、3或4的整数，且a+b≤6。
本申请所述甲基芳烃类化合物的烷基侧链氯化物，是指所述芳烃类化合物中烷基上的氢原子未全部被氯原子取代的化合物；本申请所述光氯化反应的目标产物即三氯甲基取代苯是指所述芳烃类化合物中烷基上的氢原子全部被氯原子取代的产物。
本发明氯化反应过程中，所得三氯甲基取代苯任选地进行进一步纯化或者直接进入酰氯化反应中，所得的副产物氯化氢经回收用于步骤三。
本发明所述氯化反应涉及制备三氯甲基取代苯的光化学方法，其特征在于结构式为(X)aC6H6-a-b(CH3)b的所述甲基芳烃或者该化合物的烷基侧链氯化物和氯气在光照条件下反应制备三氯甲基取代苯，其中所述光照的光源波长为约350nm-700nm、光波幅为最大约200nm，其中在反应温度约0℃-85℃、光照度约2000Lux-约55000Lux下开始通入氯气，经历在所述光照度下反应温度不超过约120℃的第一反应阶段；然后在更高的反应温度下继续通入剩余量氯气直到反应完成。在该方法的一个优选的方面，光源优选为LED灯。
本发明人发现，在氯化第一反应阶段中，优选消耗占反应所需氯气总量比例为至少约1/6的氯气量后升高温度和光照度是有利的。在本发明的一些优选方面，第一反应阶段消耗反应所需氯气总量的约1/6-约1/2；优选地，所述第一反应阶段消耗反应所需氯气总量的约1/4-约1/3。
本发明人发现，在氯化第一反应阶段中，优选所述反应温度在约55-85℃。
本发明人发现，在氯化第一反应阶段中，所述光照度为约5000Lux-约 55000Lux，优选约20000Lux-约55000Lux，更优选约35000Lux-约45000Lux。
本发明人发现，氯化第一反应阶段之后反应是在反应温度不超过约350℃、光照度不超过约100000Lux的条件下通入余量氯气。
本发明方法，氯化第一反应阶段之后的过程可以是一个单一的反应阶段，或者分为若干个反应阶段例如分为二、三、四、五、六、七、八、九、十个等反应阶段。第一反应阶段之后过程中，每个阶段的温度升高时任选地也升高光照度。更为优选的，光氯化反应第一反应阶段之后的过程还可以进一步分为第二反应阶段和第三反应阶段。第二反应阶段控制反应温度为约120-约160℃、入射光照度为约10000-约70000Lux、通入的氯气量为总量的1/4～2/5；第三反应阶段控制温度为约160-约300℃、入射光照度为约50000-约100000Lux、通入余量氯气。在第二、三阶段中，升高温度和升高光照度可互为先后顺序。
本发明方法，氯化中LED优选峰值波长范围350nm-490nm之间，或优选峰值波长460nm～490nm。
本发明方法，氯化中优选所述光源的光波幅为最大约50nm，优选为约10-约30nm，更优选为约10-约25nm。
本发明方法，氯化反应中优选在反应体系中不含外加溶剂和引发剂。
本发明氯化所述氯气的总量是能将所述甲基芳烃侧链氢原子全部氯代的氯气量，氯气总量至少是原料甲基芳烃类化合物氯化的理论摩尔量。以二(甲基)苯为例，氯气总量至少是原料二(甲基)苯摩尔数六倍的摩尔量。优选的，本发明氯化中氯气总量是相对于二(甲基)苯摩尔数六倍以上的摩尔量；氯气的过量数量可以按常规确定。优选的，出于节约反应时间的目的，本文所述各阶段中各自通入的氯气量也可根据反应监测结果适当调整。
本发明所述光波幅是指光源发出光的半峰高处的波长范围，而不是指某种光的峰值波长。例如，光波幅50nm是指光源发出光的半峰高处的波长范围不超过50nm。本发明LED光源的峰值波长可在350nm～700nm范围内变动，对于任意给定的波长，本发明入射光光源都能实现控制光波幅在50nm之内，例如465nm为峰值波幅50nm、360nm为峰值波幅50nm或586nm为峰值波幅50nm。本发明发现，LED光源还具有发热小的优点，因此可以降低生产设备的成本，例如不需要额外的降温装置，而高压汞灯光源光氯化反应时则需要相应的降温装置(例如参见US5514254)。
本发明中所述光照度可通过本领域常规仪器测定，例如照度表等。本发明中所述波长可通过本领域常规仪器测定，例如单色仪等。
本发明所述的“约”，对于温度而言是指以本数为中心值上下变动不超过2.5℃(表示为本数值±2.5℃)，优选本数值±2.5℃、±2℃或±1℃；对于光照度而言是指以本数为中心值上下变动不超过2500Lux(表示为本数值±2500Lux)，优选本数值±2500Lux、±2000Lux、±1500Lux、±1000Lux、±500Lux、±200Lux、±100Lux；对于波长而言是指以本数为中心值上下变动不超过5nm(表示为本数值±5nm)，优选本数值±4nm、±3nm或±1nm；对于光波幅而言是指以本数为中心值上下变动不超过3nm(表示为本数值±3nm)，优选本数值±2nm或±1nm。
本发明氯化反应的反应体系中优选不含外加溶剂和引发剂，更优选除了二(甲基)苯和氯以外不加入其它组分。本发明氯化可在不同阶段通过常规取样和检测方法监测反应进度，例如气相色谱，从而适当调节上述参数而节约反应时间。本文中对于三阶段时间方面的描述不是限制性的，分阶段反应时间可依氯化进度监测结果而自由调整。本文所述通入氯气速度 并不局限于特定的进料速率。当使用缓慢、逐渐等用语描述通入氯气速度时，其含义并非不清楚的。因为通入氯气的速度可由本领域技术人员根据反应监测结果适当调整。
本申请方法制得产品纯度很高。在一些实施方式中，反应后直接获得纯度在约70％～约75％之间，约75％～约80％之间，约80％～约85％之间，约85％～约90％之间，约90％～约95％之间，约95％～约99.9％之间，优选反应后直接获得纯度为约90.0％-约90.5％之间、约90.0％-约91.0％之间、约90.0％-约91.5％之间、约90.0％-约92.0％之间、约90.0％-约92.5％之间、约90.0％-约93.0％之间、约90.0％-约93.5％之间、约90.0％-约94.0％之间、约90.0％-约94.5％之间、约90.0％-约95.0％之间、约90.0％-约95.5％之间、约90.0％-约96.0％之间、约90.0％-约96.5％之间、约90.0％-约97.0％之间、约90.0％-约97.5％之间、约90.0％-约98.0％之间、约90.0％-约98.5％之间、约90.0％-约99.0％之间、约90.0％-约99.1％之间、约90.0％-约99.2％之间、约90.0％-约99.3％之间、约90.0％-约99.4％之间、约90.0％-约99.5％之间、约90.0％-约99.6％之间、约90.0％-约99.7％之间、约90.0％-约99.8％之间、约90.0％-约99.9％之间的三氯甲基取代苯。
如有需要，本发明氯化所得三氯甲基取代苯也可根据常规纯化方法进一步纯化，例如重结晶、精馏、分子蒸馏等方法。本发明优选分子蒸馏法。
本申请该步骤方法可以采用连续或者间歇方式进行，优选采用连续反应方式。
步骤二(酰氯化反应)、将步骤一制备得到的三氯甲基取代苯进一步反应制备氯甲酰基取代苯，并得到副产物氯化氢。
所得氯甲酰基取代苯任选地，进一步纯化或者直接作为成品收集；所得副产物氯化氢经回收用于步骤三。
本发明所述酰氯化反应，包含以下步骤：
i)升高温度使三氯甲基取代苯完全熔化，再加入水或所述结构式为(X)aC6H6-a-b(COOH)b的相应芳香酸及催化剂，搅拌均匀；其中X为氯或溴或氟原子，a为选自0、1、2、3、4或5的整数，b为选自1、2、3或4的整数，且a+b≤6；所述芳香酸母核上的取代基与步骤一所述甲基芳烃母核上的取代基也可以是相同的。
ii)加热反应体系维持反应进行，例如将反应体系升温至90～125℃，从而得到酰氯化反应混合物。
优选地，步骤i)中的三氯甲基取代苯与相应的芳香酸的投料摩尔比为化学反应完全进行的计量值，例如二(三氯甲基)苯与苯二甲酸的投料摩尔比优选1∶1.01～1.03。步骤i)中反应的催化剂为路易斯酸，例如三氯化铝、氯化锌、三氯化铁等，优选三氯化铁；特别地，步骤i)中为三氯甲基取代苯与水反应时优选还存在少量的相应结构式为(X)aC6H6-a-b(COOH)b的相应芳香酸。步骤i)中的催化剂加入量优选为三氯甲基取代苯质量的0.2％～0.3％。
如有需要，酰氯化所得的氯甲酰基取代苯，也可根据任选的纯化步骤，例如精馏、蒸馏、分子蒸馏或重结晶等进一步纯化，本发明优选精馏。
本申请该步骤方法可以采用连续或者间歇方式进行，优选采用连续反应方式。
步骤三(催化氧化副产物氯化氢)、收集上述步骤一和步骤二中的副产 物氯化氢，经催化剂进行催化氧化反应(Deacon反应)，制备氯气。
任选地，先将副产物氯化氢气体进行深冷或吸附进行预纯化，除去有机杂质后，再进行催化氧化。例如，氯化氢气体可以通过吸附进行纯化，适合的吸附剂材料包括例如：活性炭、氧化铝、氧化钛、二氧化硅、氧化铁、硅胶、沸石和分子筛等。本发明步骤三在一个方面涉及一种氯化氢催化氧化制备氯气的方法，包含以下步骤：
1)提供一个或多个串联或并联的装填有催化剂的反应器；
2)向所述一个或多个反应器中的第一反应器提供含氯化氢气体物流以及用于氧化所述含氯化氢气体物流的含氧气体物流，向所述一个或多个反应器中的下游反应器提供含氯化氢气体物流和/或用于氧化所述含氯化氢气体物流的含氧气体物流，以进行催化氧化氯化氢的反应；
3)将来自最后一个反应器的经过催化氧化反应的产物气体物流的一部分不经分离直接返回至任意一个或任意多个反应器；
4)将来自最后一个反应器的产物气体物流的剩余部分提供至步骤四用于分离。
本申请该步骤方法可以采用连续或者间歇方式进行，优选采用连续反应方式。
在另一个方面，本发明步骤三涉及一种氯化氢催化氧化制备氯气的方法，该方法包括：
1)提供一个或多个串联或并联的装填有催化剂的反应器；
2)向所述一个或多个反应器中的第一反应器提供含氯化氢气体物流和 用于氧化所述含氯化氢气体物流的含氧气体物流，向所述一个或多个反应器中的下游反应器提供含氯化氢气体物流和/或用于氧化所述含氯化氢气体物流的含氧气体物流，以进行催化氧化氯化氢的反应；
3)将来自最后一个反应器的产物气体物流的一部分不经分离直接返回至任意一个或任意多个反应器，优选返回至任意一个或任意多个反应器进料口之前，与要进入所述的任意一个或多个反应器的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流混合，然后进入反应器进行该催化氧化反应；
4)将来自最后一个反应器的产物气体物流的剩余部分提供至步骤四用于分离。
在本发明步骤三的一个优选实施方案中，向一个或多个反应器中的第一第一反应器提供含氯化氢气体物流以及用于氧化含氯化氢气体物流的含氧气体物流，向所述一个或多个反应器中的下游反应器提供用于氧化含氯化氢气体物流的含氧气体物流；向各反应器提供的用于氧化含氯化氢气体物流的含氧气体物流是根据需要将所需的用于氧化含氯化氢气体物流的含氧气体物流按照任意比例在各反应器之间分配的部分，优选地按照反应器的个数将所需的用于氧化含氯化氢气体物流的含氧气体物流平均分配为相应的份数。
该优选实施方案进一步优选包括以下步骤的方法：
1)提供一个或多个串联或并联的装填有催化剂的反应器；
2a)向所述一个或多个反应器中的第一反应器提供含氯化氢气体物流以 及用于氧化所述含氯化氢气体物流的含氧气体物流，以进行催化氧化氯化氢的反应；
2b)将来自所述第一反应器的产物气体物流通过换热器后提供进入下游反应器，向所述下游反应器提供用于氧化所述含氯化氢气体物流的含氧气体物流，依次向各剩余下游反应器提供来自前一反应器的产物气体物流以及用于氧化含氯化氢气体物流的含氧气体物流；
3)将来自最后一个反应器的产物气体物流的一部分不经分离返回至任意一个或任意多个反应器，优选返回至任意一个或任意多个反应器进料口之前，与要进入所述的任意一个或多个反应器的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流混合，然后进入反应器进行该催化氧化反应；
4)将来自最后一个反应器的产物气体物流的剩余部分提供至步骤四用于分离。
在本发明步骤三的另一特别优选的实施方案中，进入每个反应器的含氧气体物流的含氧量大于氧化进入每个反应器的含氯化氢气体物流所需的理论用氧量。该特别优选的实施方案，可以通过例如下述方法实施：向所述一个或多个反应器中的第一反应器提供用于氧化含氯化氢气体物流的含氧气体物流以及含氯化氢气体物流，向所述一个或多个反应器中的下游反应器提供含氯化氢气体物流；所述含氯化氢气体物流的一部分是根据需要将待氧化的含氯化氢气体物流按照任意比例在每个反应器之间分配的部分，优选地按照反应器的个数将待氧化的含氯化氢气体物流平均分配为相应的份数。
该特别优选的实施方案进一步优选包括以下步骤的方法：
1)提供一个或多个串联或并联的装填有催化剂的反应器；
2a)向所述一个或多个反应器中的第一反应器提供用于氧化氯化氢的含氧气体物流以及含氯化氢气体物流，以进行催化氧化氯化氢的反应；
2b)将来自所述第一反应器的产物气体物流通过换热器后提供进入下游反应器，向所述下游反应器提供含氯化氢气体物流，依次向各剩余下游反应器提供来自前一反应器的产物气体物流以及含氯化氢气体物流；
3)将来自最后一个反应器的产物气体物流的一部分不经分离返回至任意一个或任意多个反应器，优选返回至任意一个或任意多个反应器进料口之前，与要进入所述的任意一个或任意多个反应器的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流混合，然后进入反应器进行该催化氧化反应；
4)将来自最后一个反应器的产物气体物流的剩余部分提供至步骤四用于分离。
进一步地，进行所述的将来自最后一个反应器的产物气体物流的一部分不经分离返回至任意一个或任意多个反应器的步骤时，是优选将来自最后一个反应器的产物气体物流的一部分不经分离返回至所提供的每一个反应器；更优选返回至每一个反应器进料口之前，与所述含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流混合，然后进入反应器进行该催化氧化反应。一方面，本发明方法可以稀释进入每一个反应器的原料反应气的浓度，防止在反应器入口发生剧烈的反应，避免造成太多的热点；另一方面， 经所述混合后，本发明方法提高了原料反应气的进料温度，基本上不需要对原料反应气再进行预热。
进一步地，进行所述的将来自最后一个反应器的产物气体物流的一部分不经分离返回至所提供的每一个反应器的步骤时，返回的产物气体物流可以按照任意比例在每个反应器之间分配，例如，可以根据各个反应器的运行状况进行合理分配，优选地按照反应器的个数将返回的产物气体物流平均分配为相应的份数后分别返回至每一个反应器。本申请步骤三所述反应器优选为绝热反应器。本发明可在反应器之间连接有换热器而去除反应热，即每一反应器之后任选存在换热器。优选地，最后一个反应器后安装的换热器为气体换热器，其余反应器后面安装的换热器可以是本领域技术人员所熟知的换热器，例如管束式换热器，板式换热器、或气体换热器等。
本申请优选将所述步骤三经过催化氧化反应后的产物气体物流(高温)的剩余部分(或步骤三反应结束后的全部部分，本领域技术人员能理解最后一部分产物气体物流可以不返回)先通过气体换热器换热后再进行分离，所述换热优选是以需要进入第一反应器的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流作为冷却介质在气体换热器内进行换热；优选所述经换热后的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流被提供至第一反应器之前与被返回的从第三级反应器流出的产物气体物流的一部分混合，然后再进入第一反应器以进行催化氧化氯化氢的反应。所述产物气体物流经换热后温度降低。用作冷却介质的含氯化氢气体 物流和/或用于氧化含氯化氢气体物流的含氧气体物流经换热后温度升高，然后再将经换热后的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流提供至第一反应器以进行催化氧化氯化氢的反应；优选所述经换热后的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流被提供至第一反应器之前与被返回的从第三级反应器流出的产物气体物流的一部分混合，然后再进入第一反应器以进行催化氧化氯化氢的反应。
本发明所述提供至步骤四分离获得含氯气、含氧气和/或含氯化氢气体物流是将步骤三中产物气体物流的一部分或全部在步骤四中通过脱水、脱除(部分残余的)含氯化氢气体物流和含氧气体物流、从而得到含氯气体物流。
本申请优选可以将步骤四中从产物气体物流中分离得到的(未反应残余的)氯化氢和/或氧气再次提供到步骤三的催化氧化反应中。步骤四分离的氯化氢(或经气化后的盐酸)和/或氧气也可以被返回至步骤三中的一个或多个反应器。
在本发明的全部实施方式中，优选步骤三中所述不经分离返回至反应器的产物气体物流的部分(返回的产物气体物流)与该产物气体物流的剩余部分(剩余的产物气体物流部分)的体积比为0.25∶0.75～0.75∶0.25，优选0.35∶0.65～0.45∶0.55。
在本发明的全部实施方式中，优选所述含氯化氢气体物流(按照纯氯 化氢计算)与所述用于氧化氯化氢气体物流的含氧气物流(按照纯氧计算)的进料体积比为1∶2～5∶1，优选为1∶1.2～3.5∶1，更优选为1∶1～3∶1。
在本发明步骤三的一个优选实施方案中，所述含氯化氢气体物流(按照纯氯化氢计算)与所述用于氧化氯化氢气体物流的含氧气物流(按照纯氧计算)的进料体积比为2∶1～5∶1。
在本发明步骤三的另一特别优选的实施方案中，所述含氯化氢气体物流(按照纯氯化氢计算)与所述用于氧化氯化氢气体物流的含氧气物流(按照纯氧计算)的进料体积比为1∶2～2∶1，优选0.9∶1.1～1.1∶0.9。
在本发明的全部实施方式中，优选地，反应器内压力为：0.1-1MPa。
在本发明的全部实施方式中，优选地，反应器的进料气体温度为250～450℃，优选为300～380℃。
本申请步骤三所述的催化剂是能将氯化氢气体和氧气经氧化反应生成氯气和水的常规催化剂。合适的催化剂包括铜化合物或/和钌化合物，优选负载在载体氧化铝、或二氧化钛等上的铜化合物或/和钌化合物。例如负载有氯化铜或氯化钌的氧化铝，优选钌化合物。本申请所述合适的催化剂还可以含有其他助催化剂，例如金、钯、铂、锇、铱、镍或铬等金属的化合物，碱金属，碱土金属和稀土金属等。合适的催化剂可以具有不同的形状，例如环状物、圆柱体或球状物等，优选合适的催化剂具有相似的外部尺寸。
本申请步骤三所述反应器是常规反应装置，例如固定床或流化床反应 器，优选固定床反应器，其中可以装填有所需催化剂。
本申请所述的反应器可以选用符合反应要求的任何材质的反应器，优选纯镍或镍合金或石英的反应器。如选用多个反应器，它们之间可以采用串联或并联的方式，优选采用串联的方式从而使得氯化氢的氧化反应可以在多个阶段进行。本申请优选采用2、3、4、5、6、7、8、9、10个、更优选3或4个反应器。有利地，本领域技术人员能理解作为Deacon反应的一些原料气体将依次通过各个反应器，然后逐一向下游反应器供应额外的含氯化氢气体物流和/或用于氧化氯化氢的含氧气体物流。在本发明的全部实施方式中，优选地提供2、3、4、5、6、7、8、9、10个、特别优选3或4个串联的绝热反应器。
特别地，并联连接和串联连接的反应器还可以彼此结合。然而本发明方法特别优选具有仅仅串联连接的反应器。如果优选使用并联连接的反应器，那么特别是至多五条、优选三条、特别优选至多两条生产线(任选地包含由串联连接的反应器组成的反应器组)是并联连接的。因此，本申请所述方法可以以例如高达60个反应器操作。
本申请该步骤方法可以采用连续或者间歇方式进行，优选采用连续反应方式。
本申请所述的含氯化氢气体物流包括新鲜的含氯化氢气体物流和包含经本发明所述方法回收氯化氢或者经气化回收的盐酸的气体物流。所述新鲜的含氯化氢气体物流也可以是来自相关行业生产例如异氰酸酯的生产、酰氯的生产、芳族化合物氯化等的副产物形式的含氯化氢气体物流， 优选来自本发明步骤一和步骤二的副产物形式的含氯化氢气体物流。所述副产物形式的含氯化氢气体物流可以是经过初步处理的副产物形式的含氯化氢气体物流或者是未经任何处理的直接来自相关行业的副产物形式的含氯化氢气体物流。所述副产物形式的含氯化氢气体物流依照来源不同，可以含有少量或不含有对催化氧化氯化氢反应没有影响的也是来源于相关行业生产的其他杂质气体。其他杂质气体的量是由相关行业生产的性质决定。本领域技术人员能理解，相关行业中产生的所谓废气氯化氢对本申请而言可以是恰当的原料。
本申请所述未经反应的含氯化氢气体物流是指未通过本申请所述反应器进行催化氧化反应的含氯化氢气体物流。
本申请所述的含氧气体物流包括新鲜的含氧气体物流和含有经本发明所述方法回收的含氧气体物流。所述新鲜的含氧气体物流可以是纯氧气或者是其他含氧气体(例如空气)。
本申请所述的产物气体物流是指来自反应器的、经催化氧化反应后得到的包含氯化氢、氧气、水蒸汽和氯气的混合气体。优选地，本发明返回的产物气体物流是来自最后一个反应器的混合气体。
步骤四(分离来自步骤三的气体物流)、从上述步骤三中的产物气体物流中分离获得含氯、含氧和/或含氯化氢气体物流。
本发明步骤四所述分离获得含氯气、含氧气和/或含氯化氢气体物流包括如下步骤：
a、冷凝：对来自步骤三的产物气体物流进行冷凝处理；来自步骤三的产物 气体物流中的水连同部分未反应的氯化氢，以盐酸水溶液形式凝结析出；
b、深度脱水：将经过步骤a冷凝后的气体物流进行深度脱水，所述的深度脱水包括例如通过浓硫酸、分子筛，或者通过变温吸附、变压吸附等技术进行深度脱水，去除残余水分，减少气体物流的腐蚀性；
c、吸附：将经过步骤b深度脱水处理后的气体物流通过吸附剂进行吸附，分离氯气和氧气。
一方面，所述的吸附可以选用能够大量吸附氧气而仅少量吸附氯气的吸附剂，例如碳分子筛、硅胶等，以吸附分离去除氧气；经过上述吸附剂吸附处理后得到主要组分为氯气的含氯气体物流，其中任选含有少量氯化氢；将上述吸附剂吸附处理后被吸附到吸附剂的氧气再经过解吸处理，可以得到分离的含氧气体物流；经解吸处理后的吸附剂可继续在步骤c中用于吸附分离去除氧气。
另一方面，所述的吸附也可以选择能够大量吸附氯气而仅少量吸附氧气的吸附剂，例如细孔硅胶、活性炭等，以吸附分离去除氯气，经过上述吸附剂吸附处理后得到主要组分为氧气的含氧气体物流；将上述吸附剂吸附处理后被吸附到吸附剂的氯气再经过解吸处理，可以得到分离的含氯气体物流，其中任选含有少量氯化氢；经解吸处理后的吸附剂可继续在步骤c中用于吸附分离去除氯气。
任选地，进一步包括d、液化：将步骤c中所得到的含氯气体物流进行液化处理，分离得到含氯化氢气体物流和液化处理后的含氯气体物流。
所述步骤a中的冷凝条件为：温度为-5～5℃，压力为0.05～10MPa。
所述步骤b所述的变温吸附干燥和变压吸附干燥去除残余水分的具体操作过程，以及步骤c所述的变温变压吸附技术分离氯气和氧气的具体操作过程可参见公开号为CN103752270A的专利申请，简述如下：所述步骤b中的干燥，优选采用变温吸附干燥或变压吸附干燥过程进行干燥，所述的变温吸附干燥过程中优选采用两种吸附剂組合的复合吸附剂层，一种吸附剂是放置在吸附塔上部的氧化铝脫水剂，另一种是放置在吸附塔下部脱水干燥的吸附剂，上部氧化铝脫水剂和下部深度脱水干燥的吸附剂的体积配比为20～80％∶80％～20％。所述变压吸附干燥过程中优选采用两种吸附剂组合的复合吸附剂层，一种吸附剂是放置在吸附塔上部的氧化铝脫水剂，另一种是放置在吸附塔下部脱水干燥的吸附剂，上部氧化铝脫水剂和下部脱水干燥的吸附剂的体积配比为20～80％∶80％～20％。
步骤b中所述的变温吸附干燥过程为：将经过步骤a冷凝后的气体物流从下而上经过复合吸附剂层，所述气体物流离开变温吸附干燥装置即达到干燥目标；变温吸附干燥过程中：吸附压力为0.30～0.80MPa、吸附温度为20～50℃。所述变温吸附干燥过程包含吸附和再生操作的交替过程，其中吸附和再生的交替过程是通过常规设置(包含降压、置换、升温和冷却步骤)而实现。所述再生操作包含解吸和脱水工艺。再生操作的解吸压力为0.01～0.005MPa、再生操作的解吸温度为110～180℃；再生操作的脱水工艺采用温度为50～180℃的载气(原料气或氮气)，用原料气作载气再生时，原料气经预干燥塔干燥、再进蒸汽加热器加热后进入需升温再生脫水的吸附干燥塔，含水的载气出吸附塔后经冷却、冷凝、分离水后送回原料气***回收利用。
步骤b中的变压吸附干燥过程包含吸附和解吸工艺的交替过程，其中： 吸附压力为0.40～0.80MPa、解吸压力为0.02～-0.07MPa、吸附温度为常温；吸附和解吸工艺的交替过程是按常规设置(包含均压、冲洗置换、真空抽吸等步骤)来实现；所述变压吸附干燥过程所需装置通常按常规设置为四塔流程，此过程中冲洗置换使用干燥后的产物气体物流，冲洗置换和真空抽吸的尾气均在冷却脱水后送入脱除氯化氢的产物气体物流***回收利用。
步骤b中所述分子筛干燥的吸附剂为沸石分子筛或硅胶。
所述步骤c中的吸附优选采用变温变压吸附技术，包含吸附和解吸工艺，其中：吸附压力为0.20～0.7MPa、吸附阶段的温度由40～70℃逐渐降到20～35℃；减压解吸压力为-0.07MPa、解吸温度为40～70℃；吸附时通入小于40℃的用作原料的所述气流，开始吸附并降温；解吸再生前通入大于50℃的热氯气置换***内气体，并且升温促进解吸，达40～70℃时停送热氯气并开始真空解吸；完成解吸再生后采用氧气开始吸附前的置换；热氯气置换尾气和氧气置换的尾气均送回原料气***。
所述任选地包含步骤d是指：当参与催化氧化反应的氯化氢和氧气的比例控制适当的时候(例如，以纯氯化氢和纯氧计的比例为0.5∶1～1∶0.5时)，残余的未反应的氯化氢基本被反应生成的水在冷凝过程中吸收，当经过步骤c处理后得到的氯气中所含的氯化氢量很小，不影响氯气再循环参与氯化反应时，无需对氯气进行进一步地液化处理以分离氯化氢；而当催化氧化反应的氯化氢和氧气的比例处于其他比例时，经步骤a-c处理后仍然残留部分氯化氢，此时如有必要可在步骤d中通过液化处理含氯气和氯化氢的气体物流以分离出含氯化氢气体物流。
所述步骤d中的液化条件为：温度为-20～20℃，压力为0.05～10MPa。
步骤五(循环利用分离物)、将上述步骤四分离所得的含氯气体物流作为原料引入到包括步骤一的氯化反应中；将上述步骤四分离所得的含氯化氢和/或含氧气体物流作为原料引入到包括步骤三中的催化氧化副产物氯化氢的反应中。
其中所述步骤四分离所得的含氯气体物流也可引用到其他独立的氯化反应中。
本发明步骤三催化氧化方法所得含氯气体物流中氯气的纯度可达99.6％(体积％)以上，能够满足光氯化反应对于原料气氯气的质量要求。
本发明所述氯资源(或氯元素或氯原子)的闭路循环是指本发明所述方法通过对副产物氯化氢的合理处理，使得氯元素能够在本发明方法中循环利用。
具体实施方式
下述实施例中产物1，3-二(三氯甲基)苯、1，4-二(三氯甲基)苯、1，3-二(氯甲酰基)苯、1，4-二(氯甲酰基)苯、对氯(三氯甲基)苯、三氯甲基苯的纯度是使用气相色谱仪进行定量测定。
下述实施例中产物均三(三氯甲基)苯纯度是使用液相色谱仪测定。
实施例1
如图1所示。
步骤一、氯化反应
在3个串联的反应塔组成的连续光氯化反应装置中，从第一塔顶部以95kg/h的速度连续地加入1，3-二(甲基)苯，第一塔塔温控制在80℃～120℃，入射光中心峰值波长为460nm，塔内平均光照度为20000～39000Lux，同时从塔底以135kg/h的流量通入氯气进行连续氯化反应；第一塔的反应液从塔底溢流到第二塔，第二塔入射光中心峰值波长为505nm，平均光照度为40000～61000Lux，温度控制在135～145℃，以128kg/h的流量向第二塔中通入氯气；第二塔的反应液从塔底溢流到第三塔，第三塔入射光中心峰值波长为586nm，平均光照度为60000～86000Lux，温度控制在170～180℃，以148kg/h的流量向第三塔中通入氯气，由三个塔组成的反应体系中合计通入氯气量为411kg/h。从第三塔出口所得反应混合物即为1，3-二(三氯甲基)苯，经一次精馏纯化后得到纯化的1，3-二(三氯甲基)苯。收集光氯化反应中产生的副产物氯化氢气体。
步骤二、酰氯化反应
将步骤一得到的纯化后的1，3-二(三氯甲基)苯加入到带测温、冷凝回流和搅拌装置的配料釜中，可以设置两个或多个配料釜，实现连续给料。升温使1，3-二(三氯甲基)苯完全熔化，再按摩尔比为1，3-二(三氯甲基)苯∶1，3-二(甲酸基)苯＝1∶1.01的比例加入纯度为99.50％的1，3-二(甲酸基)苯，再按1，3-二(三氯甲基)苯∶三氯化铁＝1∶0.003的重量比加入催化剂三氯化铁；按重量比为1，3-二(三氯甲基)苯∶1，3-二(氯甲酰基)苯＝1∶1的比例加入纯度为99.0％的1，3-二(氯甲酰基)苯作反应溶剂，这些酰氯化原料在配料釜中经过加热、混合成酰氯化原料液，连续加入二级串联的酰氯化反应器，控制第一级反应器内温为100℃，第二级反应器内温为110℃，从第二级反应器出口得到1，3-二(氯甲酰基)苯，进行一次精馏后，得到纯化的1，3-二(氯甲酰基)苯。
步骤三、催化氧化副产物氯化氢
收集上述步骤一和步骤二中的副产物氯化氢，经催化剂进行催化氧化反应(Deacon反应)，制备氯气，具体包括如下步骤：
(1)首先，将副产物氯化氢气体通过吸附进行纯化，除去有机杂质。将含氯化氢气体物流和进入第一级反应器的含氧气体物流先混合，预热后通入第一级反应器内。
(2)从第一级反应器流出的气体物流通过换热器后，与要进入第二级反应器的其它气体物流混合(其它气体物流是指进入所述反应器的返回的产物气体物流与含氧气体物流和/或含氯化氢气体物流，下同)，然后进入第二级反应器；从第二级反应器流出的气体物流通过换热器后，与要进入第三级反应器的其它气体物流混合，然后进入第三级反应器。
(3)将从第三级反应器流出的产物气体物流分成两部分：一部分返回的产物气体物流被平均分配并分别返回到第一级、第二级和第三级反应器的入口前，与要进入各级反应器的含氧气体物流和/或含氯化氢气体物流混合，然后进入各级反应器；另一部分剩余的产物气体物流部分先通过气体换热器换热后再进行分离，冷却介质为需要进入第一反应器的含氯化氢气体物流和用于氧化含氯化氢气体物流的含氧气体物流。
步骤四、分离来自步骤三的气体物流
从上述步骤三中的产物气体物流中分离获得含氯、含氧或含氯化氢气体物流，包括如下步骤：
a、冷凝：对来自步骤三的产物气体物流进行冷凝处理，冷凝温度为-5～5℃，压力为0.05～10MPa。，水连同部分未反应的氯化氢，以盐酸水溶液形式凝结析出；
b、深度脱水：将经过步骤a冷凝后的气体物流通过浓硫酸进行深度脱水。
c、吸附：将经过步骤b深度脱水处理后的气体物流通过吸附剂硅胶以吸附分离去除氧气，采用变温变压吸附技术，其中：吸附压力为0.5MPa、吸附阶段的温度由60℃逐渐降到25℃；减压解吸压力为-0.07MPa、解吸温度为50℃，解吸得到分离的含氧气体物流。吸附后的剩余气体是主要组分为氯气的含氯气体物流。
d、液化：将步骤c中所得到的含氯气体物流进行液化处理，液化温度为 -20～20℃，压力为0.05～10MPa，分离得到含氯化氢气体物流和液化处理后的含氯气体物流。
步骤五、循环利用分离物
将上述步骤四分离所得的含氯气体物流作为原料引入到包括步骤一的氯化反应中。将上述步骤四分离所得的含氧气和含氯化氢气体物流，作为原料再次引入到包括步骤三的氯化氢催化氧化反应中，具体的工艺条件见表二。
反应结果见表一。
实施例2，具体操作流程如下：
步骤一、氯化反应：
氯化反应操作过程同实施例1中的步骤一。其区别在于原料采用1，4-二(甲基)苯，进入第一塔速度为100kg/h，第一塔入射光中心峰值波长为460nm，塔内平均光照度为20000～39000Lux；第二塔入射光中心峰值波长为505nm，平均光照度为40000～61000Lux，塔温度控制在135～145℃；第三塔入射光中心峰值波长为586nm，平均光照度为60000～86000Lux，温度控制在170～180℃。从第三塔出口所得反应混合物即为1，4-二(三氯甲基)苯，经一次精馏纯化后得到纯化的1，4-二(三氯甲基)苯。收集光氯化反应中产生的副产物氯化氢气体。
步骤二、酰氯化反应
酰氯化反应操作同实施例1中的步骤二，从第二级反应器出口得到1，4-二(氯甲酰基)苯，进行一次精馏后，得到纯化的1，4-二(氯甲酰基)苯。
步骤三、催化氧化副产物氯化氢
收集上述步骤一和步骤二中的副产物氯化氢，经催化剂进行催化氧化反应，制备氯气，具体的氧化过程包括如下步骤：
(1)首先，将副产物氯化氢气体通过吸附进行纯化，除去有机杂质。将含氧气体物流和进入第一级反应器的含氯化氢气体物流先混合，预热后通入第一级反应器内。
(2)从第一级反应器流出的气体物流通过换热器后，与要进入第二级反应器的其它气体物流混合，然后进入第二级反应器；从第二级反应器流出的气体物流通过换热器后，与要进入第三级反应器的其它气体物流混合，然后进入第三级反应器。
(3)将从第三级反应器流出的产物气体物流分成两部分：一部分返回的产物气体物流被平均分配并分别返回到第一级、第二级和第三级反应器的入口前，与要进入各级反应器的含氧气体物流和/或含氯化氢气体物流混合，然后进入各级反应器；另一部分剩余的产物气体物流部分先通过气体换热器换热后再进行分离，冷却介质为需要进入第一反应器的含氯化氢气体物流和用于氧化含氯化氢气体物流的含氧气体物流。
步骤四、分离来自步骤三的气体物流
从上述步骤三中的产物气体物流中分离获得含氯、含氧和含氯化氢气体物流，包括如下步骤：
a、冷凝：对来自步骤三的产物气体物流进行冷凝处理，冷凝温度为-5～5℃，压力为0.05～10MPa，水连同部分未反应的氯化氢，以盐酸水溶液形式凝结析出；
b、深度脱水：将经过步骤a冷凝后的气体物流进行深度脱水，采用变压吸附技术干燥，并采用两种吸附剂組合的复合吸附剂层，一种吸附剂是放置在吸附塔上部的氧化铝脫水剂，另一种是放置在吸附塔下部深度脱水干燥的沸石分子筛吸附剂，上部氧化铝脫水剂和下部深度脱水干燥的吸附剂的体积配比为40％∶60％。变压吸附干燥过程中吸附压力为0.40MPa、解吸压力为0.02MPa、吸附温度为常温。
c、吸附：将经过步骤b深度脱水处理后的气体物流通过吸附剂碳分子筛以吸附分离去除氧气，采用变温变压吸附技术，包含吸附和解吸工艺，其中：吸附压力为0.20MPa、吸附阶段的温度由40℃逐渐降到20℃；真空解吸压力为-0.07MPa、解吸温度为40℃，解吸得到分离的含氧气体物流。吸附后的剩余气体是主要组分为氯气的含氯气体物流。
步骤五、循环利用分离物
将上述步骤四分离所得的含氯气体物流作为原料引入到包括步骤一的氯化反应中。将上述步骤四分离所得的含氧气体物流，作为原料再次引入包括步骤三的氯化氢催化氧化反应中，具体的工艺条件见表二。
反应结果见表一。
实施例3按照上述实施例2的具体步骤进行操作，其中甲基芳烃原料为对氯甲苯，与氯气反应后得到对氯(三氯甲基)苯，相应的芳香酸为对氯苯甲酸。
实施例4按照上述实施例2的具体步骤进行操作，其中甲基芳烃原料为甲基苯，与氯气反应后得到三氯甲基苯，相应的芳香酸为苯甲酸。
实施例5按照上述实施例1的具体步骤进行操作，其中甲基芳烃原料为均三甲苯，与氯气反应后得到均三(三氯甲基)苯，相应的芳香酸为均苯三甲酸。
实施例6
步骤一、氯化反应：
在带有测温、回流冷凝器、LED灯照射装置和加热冷却装置的反应塔中加入1，3-二(甲基)苯100千克，升温至80℃后停止升温，开启LED灯照射，入射光中心峰值波长为360nm，光照度为49000Lux，通入氯气反应，控制氯气通入速率使得体系温度不超过120℃，消耗氯气量为135.8千克，第一反应阶段耗时4小时30分钟。调节光照度至60000Lux，升高体系温度至140℃后，继续通入氯气，消耗氯气136千克，第二反应阶段耗时3小时55分钟。维持光照度至60000Lux，升高体系温度至160℃后，通入氯气155.2千克，第三反应阶段共耗时16小时35分钟。反应总计消耗氯气量为427千克。反应结束后的反应混合物即为1，3-二(三氯甲基)苯粗品，经过一次精馏 纯化后得到纯度为99.42％的1，3-二(三氯甲基)苯255千克。对氯化反应中产生的气体进行收集，得到122m3的副产物氯化氢。
步骤二、酰氯化反应
在带测温、冷凝回流和搅拌装置的反应釜中，加入步骤一制备得到的纯度为99.42％的1，3-二(三氯甲基)苯255千克，升温使1，3-二(三氯甲基)苯完全熔化，按照1，3-二(三氯甲基)苯摩尔数的1.01倍加入纯度为99.50％的1，3-二(甲酸基)苯137.3千克，再按照1，3-二(三氯甲基)苯重量的0.30％加入三氯化铁催化剂0.77千克，继续升温至110℃后维持60分钟反应结束。将所得产物进行精馏，得到纯化的1，3-二(氯甲酰基)苯，纯度为99.97％。对酰氯化反应中产生的气体进行收集，得到34m3的副产物氯化氢。
步骤三、催化氧化副产物氯化氢
将本实施例中产生的二种副产氯化氢气体连续地压缩进入催化氧化***，操作过程同实施例2中的步骤三。
步骤四、分离来自步骤三的气体物流
从上述步骤三中的产物气体物流中分离获得含氯、含氧气体物流，包括如下步骤：
a、冷凝：对来自步骤三的产物气体物流进行冷凝处理，冷凝温度为-5～5℃，压力为0.05～10MPa。，水连同部分未反应的氯化氢，以盐酸水溶液形式凝结析出；
b、深度脱水：将经过步骤a冷凝后的气体物流进行深度脱水，采用变温吸附技术干燥，并采用两种吸附剂組合的复合吸附剂层，一种吸附剂是放置在吸附塔上部的氧化铝脫水剂，另一种是放置在吸附塔下部深度脱水干燥的沸石分子筛吸附剂，上部氧化铝脫水剂和下部深度脱水的吸附剂的体积配比为30％∶70％。变温吸附干燥过程中吸附压力为0.70MPa、吸附温度为30℃；再生操作包含解吸和脱水工艺。再生操作的解吸压力为0.009MPa、再生操作的解吸温度为160℃；再生操作的脱水工艺采用温度为180℃的载气。
c、吸附：将经过步骤b深度脱水处理后的气体物流通过吸附剂碳分子筛以 吸附分离去除氧气，采用变温变压吸附技术，包含吸附和解吸工艺，其中：吸附压力为0.5MPa、吸附阶段的温度由60℃逐渐降到25℃；减压解吸压力为-0.07MPa、解吸温度为50℃，解吸得到分离的含氧气体物流。吸附后的剩余气体是主要组分为氯气的含氯气体物流。
d、液化：将步骤c中所得到的含氯气体物流进行液化处理，液化温度为-20～20℃，压力为0.05～10MPa，分离得到含氯化氢气体物流和液化处理后的含氯气体物流。
经计量步骤四分离后得到的氯气量为195Kg，氯气纯度为99.97(v.％)，分离后得到的回收氯化氢量为18m3；分离后得到的回收氧气量为30m3。
步骤五、循环利用分离物
将上述步骤四分离所得的含氯气体物流作为原料引入到包括步骤一的氯化反应中，并将所得粗品进行一次精馏纯化后，得到纯度为99.4％的1，3-二(三氯甲基)苯，进一步地将所得纯化后的1，3-二(三氯甲基)苯引入到包括步骤二的酰氯化反应中，得到纯度为99.96％的1，3-二(氯甲酰基)苯。反应结果见表一。
下表一为实施例1-5主要物料的进入和产出量；表二为各实施例步骤二氧化单元操作条件。
Figure PCTCN2015079272-appb-000001
Figure PCTCN2015079272-appb-000002

Claims (22)

  1. 一种制备氯甲酰基取代苯的清洁工艺，包含以下步骤：
    步骤一、将结构式为(X)aC6H6-a-b(CH3)b的甲基芳烃或其烷基侧链氯化物和氯气在光照条件下反应制备三氯甲基取代苯、并得到副产氯化氢，其中X为氯或溴或氟原子，a为选自0、1、2、3、4或5的整数，b为选自1、2、3或4的整数，且a+b≤6，并且所述烷基侧链氯化物是指所述甲基芳烃化合物中侧链烷基上的氢原子未全部被氯原子取代的化合物；
    步骤二、将步骤一制备得到的三氯甲基取代苯与水或结构式为(X)aC6H6-a-b(COOH)b的相应芳香酸进一步反应制备氯甲酰基取代苯，并得到副产氯化氢，所述相应芳香酸是指所述芳香酸母核上的取代基与上述甲基芳烃或其烷基侧链氯化物的母核上的取代基处于相同或相应的取代位置，其中X为氯或溴或氟原子，a为选自0、1、2、3、4或5的整数，b为选自1、2、3或4的整数，且a+b≤6；
    步骤三、将包含来自上述步骤一和步骤二的氯化氢气体物流进行催化氧化反应；
    步骤四、从上述步骤三中的产物气体物流中分离获得含氯气、含氧气和/或含氯化氢气体物流。
    步骤五将上述步骤四分离所得的含氯气体物流作为原料引入到步骤一的氯化反应中。
  2. 根据权利要求1所述的清洁工艺，其特征在于：在步骤五中包含将经步骤四分离所得的含氯化氢和/或含氧气体物流，作为原料再引入步骤三的氯 化氢催化氧化反应中。
  3. 根据权利要求1-2任一项所述的清洁工艺，其特征在于：所述步骤三中副产物氯化氢的氧化，包含以下步骤：
    1)提供一个或多个串联或并联的装填有催化剂的反应器(优选绝热反应器)；
    2)向所述一个或多个反应器中的第一反应器提供含氯化氢气体物流以及用于氧化所述含氯化氢气体物流的含氧气体物流，向所述一个或多个反应器中的下游反应器提供含氯化氢气体物流和/或用于氧化所述含氯化氢气体物流的含氧气体物流，以进行催化氧化氯化氢的反应；
    3)将来自最后一个反应器的经过催化氧化反应的产物气体物流的一部分不经分离直接返回至任意一个或任意多个反应器；
    4)将来自最后一个反应器的产物气体物流的剩余部分提供至步骤四用于分离。
  4. 根据权利要求3所述的清洁工艺，其特征在于：步骤3)中将来自最后一个反应器的产物气体物流的一部分不经分离返回至任意一个或任意多个反应器进料口之前，与要进入所述的任意一个或多个反应器的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流混合，然后再进入反应器进行该催化氧化反应。
  5. 根据权利要求3或4任一项所述的清洁工艺，其特征在于：步骤3)中将来自最后一个反应器的产物气体物流的一部分不经分离返回至所提供的每一个反应器。
  6. 根据权利要求5所述的清洁工艺，其特征在于：进行所述的将来自最后一个反应器的产物气体物流的一部分不经分离返回至所提供的每一个反应器的步骤时，返回的产物气体物流可以按照任意比例在每个反应器之间分配；优选地按照反应器的个数将返回的产物气体物流平均分配为相应的份数后分别返回至每一个反应器。
  7. 根据权利要求3-6任一项的清洁工艺，其特征在于：向所述一个或多个反应器中的第一反应器提供含氯化氢气体物流以及用于氧化所述含氯化氢气体物流的含氧气体物流，向所述一个或多个反应器中的下游反应器提供用于氧化含氯化氢气体物流的含氧气体物流。
  8. 根据权利要求7的清洁工艺，其特征在于：向各反应器提供的用于氧化含氯化氢气体物流的含氧气体物流是根据需要将所需的用于氧化含氯化氢气体物流的含氧气体物流按照任意比例在各反应器之间分配的部分，优选地按照反应器的个数将所需的用于氧化含氯化氢气体物流的含氧气体物流平均分配为相应的份数。
  9. 根据权利要求3-6任一项的清洁工艺，其特征在于：向所述一个或多个反应器中的第一反应器提供用于氧化含氯化氢气体物流的含氧气体物流以及含氯化氢气体物流，向所述一个或多个反应器中的下游反应器提供含氯化氢气体物流；优选进入每个反应器的含氧气体物流中的含氧量大于氧化进入每个反应器的含氯化氢气体物流所需的理论用氧量。
  10. 根据权利要求9的清洁工艺，其特征在于：向各反应器提供的含氯化氢气体物流是根据需要将待氧化的含氯化氢气体物流按照任意比例在每个反 应器之间分配的部分，优选地按照反应器的个数将待氧化的含氯化氢气体物流平均分配为相应的份数。
  11. 根据权利要求1-10任一项的清洁工艺，其特征在于，所述步骤三中副产物氯化氢的氧化，包含以下步骤：
    1)提供一个或多个串联或并联的装填有催化剂的反应器；
    2a)向所述一个或多个反应器中的第一反应器提供含氯化氢气体物流以及用于氧化所述含氯化氢气体物流的含氧气体物流，以进行催化氧化氯化氢的反应；
    2b)将来自所述第一反应器的产物气体物流通过换热器后提供进入下游反应器，向所述下游反应器提供用于氧化所述含氯化氢气体物流的含氧气体物流，依次向各剩余下游反应器提供来自前一反应器的产物气体物流以及用于氧化含氯化氢气体物流的含氧气体物流；
    3)将来自最后一个反应器的产物气体物流的一部分不经分离返回至任意一个或任意多个反应器，优选返回至所述任意一个或多个反应器进料口之前，与要进入所述的任意一个或多个反应器的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流混合，然后进入反应器以进行该催化氧化反应；
    4)将来自最后一个反应器的产物气体物流的剩余部分提供至步骤四用于分离。
  12. 根据权利要求1-10任一项的清洁工艺，其特征在于所述步骤三中副产物氯化氢的氧化，包含以下步骤：
    1)提供一个或多个串联或并联的装填有催化剂的反应器；
    2a)向所述一个或多个反应器中的第一反应器提供用于氧化氯化氢的含氧气体物流以及含氯化氢气体物流，以进行催化氧化氯化氢的反应；
    2b)将来自所述第一反应器的产物气体物流通过换热器后提供进入下游反应器，向所述下游反应器提供含氯化氢气体物流，依次向各剩余下游反应器提供来自前一反应器的产物气体物流以及含氯化氢气体物流；
    3)将来自最后一个反应器的产物气体物流的一部分不经分离返回至任意一个或多个反应器，优选返回至任意一个或多个应器进料口之前，与要进入所述的任意一个或任意多个反应器的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流混合，然后进入反应器进行该催化氧化反应；
    4)将来自最后一个反应器的产物气体物流的剩余部分提供至步骤四用于分离。
  13. 根据权利要求3-12任一项的清洁工艺，其特征在于：每一反应器之后任选配置换热器用以去除反应热，位于反应器之后的换热器可以是本领域技术人员所熟知的换热器，例如管束式换热器，板式换热器或体换热器等；优选在最后一个反应器之后配置气体换热器。
  14. 根据权利要求13的清洁工艺，其特征在于：来自最后一个反应器的经过催化氧化反应的产物气体物流的剩余部分先通过气体换热器换热后再进行分离，所述换热优选是以要进入第一反应器的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流作为冷却介质在气体换热器内 进行换热；优选所述经换热后的含氯化氢气体物流和/或用于氧化含氯化氢气体物流的含氧气体物流被提供至第一反应器之前与被返回的从第三级反应器流出的产物气体物流的一部分混合，然后再进入第一反应器以进行催化氧化氯化氢的反应。
  15. 根据权利要求3-14任一项的清洁工艺，其特征在于：所述来自最后一个反应器的不经分离直接返回至反应器的产物气体物流与来自最后一个反应器的剩余部分产物气体物流的体积比为0.25∶0.75～0.75∶0.25，优选0.35∶0.65～0.45∶0.55。
  16. 根据权利要求3-15任一项的清洁工艺，其特征在于：所述含氯化氢气体物流(按照纯氯化氢计算)与所述含氧气体物流(按照纯氧计算)的进料体积比为1∶2～5∶1，优选为1∶1.2～3.5∶1，更优选为1∶1～3∶1。
  17. 根据权利要求3-16任一项的方法，其特征在于：所述含氯化氢气体物流(按照纯氯化氢计算)与所述含氧气体物流(按照纯氧计算)的进料体积比为2∶1～5∶1。
  18. 根据权利要求3-16任一项的清洁工艺，其特征在于：所述含氯化氢气体物流(按照纯氯化氢计算)与所述含氧气体物流(按照纯氧计算)的进料体积比为1∶2～2∶1，优选1.1∶0.9～0.9∶1.1。
  19. 根据权利要1～18任一项所述的清洁工艺，其特征在于，所述步骤四中所述分离来自步骤三的产物气体物流的过程包括如下步骤：
    a、冷凝：对来自步骤三的产物气体物流进行冷凝处理，步骤三反应产生的水连同部分步骤三中未反应的氯化氢，以盐酸水溶液形式凝结出来；
    b、深度脱水：将经过步骤a冷凝后的产物气体物流进行深度脱水，所述的深度脱水包括例如通过浓硫酸、分子筛，或者通过变温吸附、变压吸附等技术进行深度脱水，去除残余水分；
    c、吸附：将经过步骤b深度脱水处理后的气体物流通过吸附剂进行吸附，分离氯气和氧气；
    任选地，进一步包括d、液化：将步骤c中所得到的含氯气体物流进行液化处理，分离得到含氯化氢气体物流和液化处理后的含氯气体物流。
  20. 根据权利要求1-19任一项所述的清洁工艺，其特征在于所述步骤一的氯化中甲基芳烃和氯气在光照条件下反应制备三氯甲基取代苯，其中所述光照的光源波长为约350nm-700nm、光波幅为最大约200nm，其中在反应温度约0℃-85℃、光照度约2000Lux-约55000Lux下开始通入氯气，经历在所述光照度下反应温度不超过约120℃的第一反应阶段；然后在更高的反应温度下继续通入剩余量氯气直到反应完成。
  21. 根据权利要求20所述的清洁工艺，其特征在于，所述光照的光源为LED灯。
  22. 根据权利要求1-21所述的清洁工艺，其特征在于，所述步骤二酰氯化优选包含以下步骤：
    i)升高温度使三氯甲基取代苯完全熔化，再加入水或相应的芳香酸及催化剂，搅拌均匀；
    ii)加热反应体系维持反应进行。
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