WO2023040641A1 - 一种酰化液及利用酰化液连续合成酰基萘的工艺方法 - Google Patents

一种酰化液及利用酰化液连续合成酰基萘的工艺方法 Download PDF

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WO2023040641A1
WO2023040641A1 PCT/CN2022/115478 CN2022115478W WO2023040641A1 WO 2023040641 A1 WO2023040641 A1 WO 2023040641A1 CN 2022115478 W CN2022115478 W CN 2022115478W WO 2023040641 A1 WO2023040641 A1 WO 2023040641A1
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liquid
acylation
port
tower
temperature
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PCT/CN2022/115478
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English (en)
French (fr)
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毛学锋
李恒
胡发亭
钟金龙
李军芳
王通
张笑然
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煤炭科学技术研究院有限公司
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Priority claimed from CN202111075055.5A external-priority patent/CN113842857B/zh
Priority claimed from CN202111093577.8A external-priority patent/CN113773179B/zh
Priority claimed from CN202111638857.2A external-priority patent/CN114369018B/zh
Priority claimed from CN202210095938.0A external-priority patent/CN114516788B/zh
Priority claimed from CN202210135569.3A external-priority patent/CN114394880A/zh
Priority claimed from CN202210530956.7A external-priority patent/CN114890623B/zh
Application filed by 煤炭科学技术研究院有限公司 filed Critical 煤炭科学技术研究院有限公司
Priority to JP2022573236A priority Critical patent/JP2023546762A/ja
Priority to US18/057,394 priority patent/US20230095165A1/en
Publication of WO2023040641A1 publication Critical patent/WO2023040641A1/zh

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C45/00—Preparation of compounds having >C = O groups bound only to carbon or hydrogen atoms; Preparation of chelates of such compounds
    • C07C45/78—Separation; Purification; Stabilisation; Use of additives
    • C07C45/81—Separation; Purification; Stabilisation; Use of additives by change in the physical state, e.g. crystallisation
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J19/0086—Processes carried out with a view to control or to change the pH-value; Applications of buffer salts; Neutralisation reactions
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J19/0093—Microreactors, e.g. miniaturised or microfabricated reactors
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J19/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J19/18—Stationary reactors having moving elements inside
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C45/00—Preparation of compounds having >C = O groups bound only to carbon or hydrogen atoms; Preparation of chelates of such compounds
    • C07C45/45—Preparation of compounds having >C = O groups bound only to carbon or hydrogen atoms; Preparation of chelates of such compounds by condensation
    • C07C45/46—Friedel-Crafts reactions
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C45/00—Preparation of compounds having >C = O groups bound only to carbon or hydrogen atoms; Preparation of chelates of such compounds
    • C07C45/78—Separation; Purification; Stabilisation; Use of additives
    • C07C45/81—Separation; Purification; Stabilisation; Use of additives by change in the physical state, e.g. crystallisation
    • C07C45/82—Separation; Purification; Stabilisation; Use of additives by change in the physical state, e.g. crystallisation by distillation
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C45/00—Preparation of compounds having >C = O groups bound only to carbon or hydrogen atoms; Preparation of chelates of such compounds
    • C07C45/78—Separation; Purification; Stabilisation; Use of additives
    • C07C45/81—Separation; Purification; Stabilisation; Use of additives by change in the physical state, e.g. crystallisation
    • C07C45/82—Separation; Purification; Stabilisation; Use of additives by change in the physical state, e.g. crystallisation by distillation
    • C07C45/84—Separation; Purification; Stabilisation; Use of additives by change in the physical state, e.g. crystallisation by distillation by azeotropic distillation
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2219/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J2219/00049—Controlling or regulating processes
    • B01J2219/00051—Controlling the temperature
    • B01J2219/00074—Controlling the temperature by indirect heating or cooling employing heat exchange fluids
    • B01J2219/00087—Controlling the temperature by indirect heating or cooling employing heat exchange fluids with heat exchange elements outside the reactor
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2219/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J2219/00049—Controlling or regulating processes
    • B01J2219/00051—Controlling the temperature
    • B01J2219/00074—Controlling the temperature by indirect heating or cooling employing heat exchange fluids
    • B01J2219/00087—Controlling the temperature by indirect heating or cooling employing heat exchange fluids with heat exchange elements outside the reactor
    • B01J2219/00094—Jackets
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2219/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J2219/00781—Aspects relating to microreactors
    • B01J2219/00873—Heat exchange
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2219/00—Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J2219/00781—Aspects relating to microreactors
    • B01J2219/00891—Feeding or evacuation
    • B01J2219/00898—Macro-to-Micro (M2M)

Definitions

  • the application relates to the technical field of chemical synthesis, in particular to an acylation solution and a process for continuously synthesizing acylnaphthalene using the acylation solution.
  • 2,6-naphthalene dicarboxylic acid is an important monomer of special high-end polyester polyethylene naphthalate (PEN), and is a key monomer for the synthesis of various high-performance polynaphthyl esters, polyurethanes and liquid crystal polyester resins.
  • PEN polyethylene naphthalate
  • PET polyethylene terephthalate
  • packaging containers and electronic components and other fields have broad application prospects.
  • 2,6-naphthalene dicarboxylic acid (2,6-NDCA) can be divided into two categories: one is the traditional process method based on petroleum-based benzene series compounds, such as BP-Amoco, Japan Mitsubishi Chemical Corporation, Chevron and Finnish Optatech adopt this method, which has many steps and high production cost; the other is an improved process method based on coal-based naphthalene compounds as raw materials, using naphthalene and methylnaphthalene as raw materials to carry out 2,6-di
  • the synthesis of methylnaphthalene (2,6-DMN) is one-step preparation of DMN products, and then 2,6-DMN can be obtained through isomerization, separation and purification. Exxon-Mobil Company adopts this method. This method has simple steps and low cost. relatively low.
  • the industrialization route and semi-industrialization route are:
  • Naphthalene undergoes liquid-phase alkylation to produce 2,6-diisopropylnaphthalene, which is then oxidized to produce 2,6-naphthalene dicarboxylic acid.
  • Naphthalene reacts with propylene under the catalysis of acid-treated mordenite, and the reaction mixture is obtained by rectification to obtain crude dialkylnaphthalene (a mixture of isomers), and then further separated to obtain 2,6-dialkyl Naphthalene; 2,6-naphthalene dicarboxylic acid can be obtained by air liquid-phase oxidation.
  • This method has many reaction steps, and alkylation and purification are the key points of this technology. At present, there are two main industrial synthesis routes of 2,6-NDCA.
  • One is acetylation of 2-methylnaphthalene to obtain 2-methyl-6-acetylnaphthalene, and further oxidation to obtain 2,6-NDCA.
  • the second is the oxidation of 2,6-dialkylnaphthalene to 2,6-NDCA, including liquid-phase catalysis of 2,6-dimethylnaphthalene, 2,6-diethylnaphthalene, and 2,6-diisopropylnaphthalene oxidation.
  • the catalyst used in the preparation of 2-methyl-6-acetylnaphthalene is relatively expensive, and it is difficult to recycle, the cost is high, the reactants are highly toxic, and the pollution is large, so the production It is also smaller.
  • the latter after the alkylation of naphthalene, uses a Co-Mn-Br catalyst system, uses lower fatty acid as a solvent, and reacts in a titanium material reactor at a temperature of about 200°C and a pressure of about 3.0MPa. The reaction conditions are mild and easy. control.
  • 2-methylnaphthalene is used as a raw material to generate 2-methyl-6-propionylnaphthalene through acylation reaction, and then oxidized to obtain 2,6-naphthalene dicarboxylic acid.
  • This synthesis method has rich sources of raw materials, less side reactions, and easy product Refining is much easier than the synthetic route of preparing 2,6-naphthalene dicarboxylic acid by oxidation of 2,6-dialkylnaphthalene.
  • an acylation solution with high stability and high uniformity is very critical.
  • Lewis acid is used as a catalyst to form an electrophilic complex, which is easy to enter the para-position of the acylated product, so the selectivity of the reaction is high.
  • the catalyst and acylating agent due to the high activity of the catalyst and acylating agent, it is easy to react with water and deteriorate, and a large amount of white acid mist will be generated when it contacts the air with high humidity. Stablize.
  • the reaction needs to be carried out in a homogeneous phase. Undissolved catalysts in the solution and particulate matter generated by deterioration will affect the efficiency of the acylation liquid reaction and affect the yield and purity of the final product.
  • 2-methylnaphthalene is used as a raw material
  • propionyl chloride is an acylating agent
  • aluminum trichloride is a catalyst
  • nitrobenzene is a solvent
  • Friedel-Crafts acylation is carried out at normal temperature and pressure to prepare 2-methyl- 6-Propionylnaphthalene.
  • the reaction needs to be hydrolyzed and quenched, and the oil phase is hydrolyzed to a pH of 6-7, and then purification methods such as vacuum distillation, rectification, and recrystallization are used to obtain 2-methyl- 6-Propionylnaphthalene.
  • the reaction temperature of this reaction has certain characteristics.
  • the two materials need to be mixed at low temperature to react first, and then the temperature is raised to react for a period of time.
  • the reported acylation reaction adopts intermittent tank reactor or continuous microchannel reactor, but the heat exchange area of tank reactor is small, the reaction temperature is not easy to control, the residence time is inconsistent, and the product quality is not easy to stabilize; while the microchannel reaction
  • the device has the risk of clogging, and the cost of amplification is high, and the operation is more complicated.
  • acylation reaction adopts the shortcoming of intermittent tank reactor to be that heat exchange area is little, and reaction temperature is difficult for controlling, and reaction time is longer, and residence time is inconsistent, and product quality is not easy to stabilize, and product yield is lower;
  • Acylation reaction adopts continuous
  • the disadvantage of the microchannel reactor is that there is a risk of blockage, the cost of amplification is high, and the operation is more complicated; the reaction temperature of this reaction has certain characteristics, and the two materials of the acylation liquid and the raw material liquid need to be mixed at a low temperature to react first, and then the temperature Elevating the reaction for a period of time, the existing device does not fully meet the needs of the reaction characteristics.
  • the hydrolysis process of the acylation reaction solution containing 2-methyl-6-acylnaphthalene plays an important role in obtaining 2-methyl-6-acylnaphthalene with high purity and high yield.
  • the hydrolysis can stop the acylation reaction, hydrolyze the aluminum trichloride and the acylating agent, dissolve the HCl gas generated in the reaction, and keep the 2-methyl-6-acylnaphthalene in the oil phase.
  • the pH value of the oil phase is increased to 6-7 to eliminate the influence of acidity on the product.
  • the acyl group is oxidized in a weak acid and prone to condensation and asphaltification, generating tar and reducing the yield of the product and purity. This just requires that when hydrolyzing the acylation reaction liquid, accomplish synchronous and continuous hydrolysis to ensure that the 2-methyl-6-propionylnaphthalene with higher purity is obtained.
  • the first one is to first flow the acylation reaction liquid into the mixer for storage, and then enter the microchannel reactor of the hydrolysis section at the back end after mixing with water, which cannot be done
  • the acylation reaction and the hydrolysis reaction are carried out simultaneously, and a large amount of heat will be released when the acylation reaction liquid contacts with water, and the temperature cannot be cooled down in time;
  • the second is to pass the reacted product into the hydrolysis reactor, and the reaction
  • the mixed solution was poured into a large beaker equipped with ice cubes, and mechanically stirred continuously during the pouring process. After pouring, distilled water was added to continue stirring for half an hour to completely hydrolyze the acylated product.
  • This scheme is in the intermittent hydrolysis reaction, the time is long, the operation is more complicated, and the hydrolysis effect is general;
  • the solvent nitrobenzene was recovered by distillation to obtain the crude product 2-methyl-6-acetylnaphthalene.
  • This scheme is the hydrolysis of ethanol aqueous solution, does not highlight its advantage.
  • the prior art has the following disadvantages: intermittent or semi-continuous hydrolysis does not achieve synchronous continuous hydrolysis. After the acylation reaction solution is first obtained, there is a certain time interval before the hydrolysis reaction is carried out; the acylation reaction solution cannot be hydrolyzed in time.
  • the wastewater after the acylation reaction is mainly highly toxic, high-salt organic wastewater that is difficult to biodegrade (Table 1), which is extremely difficult to treat. It can be said that under the current situation of increasingly strict environmental regulation, the difficult biodegradable organic wastewater produced by the acylation reaction is the main sticking point restricting the industrialization of the Friedel-Crafts acylation reaction.
  • This application aims to solve one of the technical problems in the related art at least to a certain extent, and proposes a process for continuously synthesizing acylnaphthalene using an acylation liquid and acylating liquid. Poor homogeneity and the use of acylating liquid to prepare acyl naphthalene have low reaction efficiency, high loss, unstable hydrolysis process and easy blockage of pipelines and low yield and low purity of acyl naphthalene.
  • a new acylating liquid is proposed, and the use of acyl
  • the chemical solution improves the reaction rate and hydrolysis stability of acylnaphthalene synthesis, thereby increasing the yield of acylnaphthalene; in addition, the embodiment of the present application also treats the hydrolysis wastewater after acylnaphthalene synthesis by acylation solution, so as to realize green production.
  • an acylation liquid which is configured by a configuration system, and the configuration system includes:
  • a kettle body the kettle body defines a first chamber, and the kettle body has a first feeding port, a second feeding port, a stirring port and a liquid outlet communicated with the first chamber;
  • the feed port cover is detachably sealed and mounted on the first feed port
  • a feed pump the feed pump is sealingly connected to the second feed port, so as to add the acylating agent to the second feed port;
  • a stirrer a part of the stirrer is sealed and inserted in the stirring port, the stirrer includes a stirring shaft and a stirring blade, the stirring blade is arranged on the stirring shaft, at least a part of the stirring shaft and the each of the stirring blades is disposed within the first chamber;
  • Using the configuration system to prepare the acylation solution includes the following steps:
  • step b Add the solvent and the Lewis catalyst weighed in step a into the first chamber through the first feeding port and stir evenly with the agitator to obtain a mixed solution;
  • step c using the feed pump to add an acylating agent to the mixed solution obtained in step b to obtain an acylated liquid.
  • the kettle body includes:
  • the liquid outlet is arranged on the first housing, and the upper end of the first housing is open;
  • the upper cover, the upper cover sealing cover is mounted on the first housing, the first housing and the upper cover define the first chamber, the first feeding port, the second Both the feeding port and the stirring port are arranged on the upper cover;
  • a jacket the jacket is sheathed on the first shell, and the jacket has a first heat exchange medium inlet for the heat exchange medium to enter and a first heat exchange medium outlet for the heat exchange medium to flow out of.
  • the configuration system further includes a heating device, the heating device has a second heat exchange medium inlet and a heat exchange medium second outlet, and the heat exchange medium first outlet is connected to the heat exchange medium second outlet.
  • the inlets are connected, and the first inlet of the heat exchange medium is connected with the second outlet of the heat exchange medium;
  • the kettle body has a temperature measuring port communicating with the first chamber, a part of the temperature sensor is sealed and inserted in the temperature measuring port, and the detection end of the temperature sensor is set at the first chamber inside the chamber;
  • a controller connected to each of the heating device and the temperature sensor, so that the controller controls the heating device according to the temperature detected by the temperature sensor.
  • the configuration system also includes a suction filtration device, and the suction filtration device includes:
  • a filter device the filter device includes a second housing and a filter membrane, the second housing defines a second chamber, the filter membrane is arranged in the second chamber, the filter membrane divides the first The two chambers are divided into a first part and a second part, the second housing is provided with a filter device inlet and a filter device outlet communicating with the second chamber, and the filter device inlet is set corresponding to the first part, The outlet of the filter device is set corresponding to the second part, and the inlet of the filter device is connected to the liquid outlet;
  • a liquid storage tank includes a third housing, the third housing defines a third chamber, the third housing is provided with a liquid storage tank inlet and a liquid storage tank outlet, the storage tank The inlet of the liquid tank communicates with the outlet of the filtering device;
  • a vacuum pump the vacuum pump has a first vacuum port, the third housing is provided with a second vacuum port communicated with the third chamber, the first vacuum port is connected to the second The vacuum port is connected.
  • the first housing, the second housing and the third housing are of an integrated structure, wherein the liquid outlet is arranged at the bottom of the first housing, and the The filter device inlet is set on the top of the second housing, the filter device is set at the bottom of the kettle body, the filter device outlet is set at the bottom of the second shell, and the liquid storage tank inlet is set On the top of the third housing, the liquid storage tank is arranged at the bottom of the filtering device.
  • the solvent is nitrobenzene
  • step b is:
  • the nitrobenzene is added into the first chamber through the first feeding port;
  • step a the Lewis catalyst weighed in step a is added in the nitrobenzene;
  • the nitrobenzene and the Lewis catalyst are heated to 50° C.-60° C. using a heat exchange medium, while stirring at 200 rpm-400 rpm using the stirrer.
  • using the configuration system to prepare the acylation liquid further includes step d, using the suction filtration device to perform suction filtration treatment on the acylation liquid obtained in step c under an inert gas atmosphere, so as to obtain a filtered acylation liquid.
  • step c the adding speed of the acylating agent is 3-10 drops per second, and the stirring speed of the agitator is 200rpm-400rpm.
  • the molar mass ratio of the acylating agent to the Lewis catalyst is (1.1-1.5):(1.3-1.7); and/or
  • the molar mass ratio of the Lewis catalyst to the solvent is (1.3-1.7):5.
  • a method for continuously synthesizing acylnaphthalene is proposed, and the synthesis using any of the acylating liquids described above includes the following steps:
  • acylation reaction solution enters the microchannel reactor and the tank reactor connected in series for acylation reaction and simultaneous hydrolysis and subsequent rectification and crystallization to obtain 2-methyl-6-propionylnaphthalene .
  • microchannel reactors are provided and connected in parallel.
  • step S1 the raw material solution and the acylation solution are injected into a three-way mixer through a syringe for mixing.
  • the three-way mixer is a T-type mixer or a Y-type mixer.
  • the three-way mixer and the microchannel reactor are placed in the first constant temperature tank, the temperature is controlled at -5-0°C, and the tank reactor is placed in the second constant temperature tank , the temperature is controlled at 30-50°C.
  • the total residence time of the acylation reaction liquid in the tank reactor is 50-80 min.
  • the microchannel reactor and the tank reactor are made of materials resistant to strong acid corrosion.
  • the inner diameter of the microchannel reactor is 0.5-3.175 mm.
  • the acylating agent in the acylating solution is any one of propionyl chloride, acetyl chloride, acetic anhydride, and propionic anhydride.
  • the hydrolysis method is that in the hydrolysis section, water is passed into the water phase pipeline of the hydrolysis section before the acylation reaction solution after the acylation reaction is passed into the hydrolysis section, and after the water flows out of the outlet of the hydrolysis section
  • the acylation reaction solution is passed into the oil phase pipeline of the hydrolysis section, and mixed rapidly with water in a low-temperature cooling bath, and then enters a tubular reactor for hydrolysis reaction, and the hydrolyzed mixed solution is obtained from
  • the outlet of the tubular reactor flows out, and the liquid separator is used to collect the mixed liquid to separate the water phase and the oil phase; at this stage, the water is always kept in a state of circulation until the mixed liquid is completely discharged, so that the hydrolysis reaction and the acylation reaction are synchronized conduct.
  • the flow rate of water in the water phase pipeline is controlled to be 3-15 mL/min by turning on the water injection pump to pass water to the water phase pipeline of the hydrolysis section.
  • the acylation reaction liquid directly enters the oil phase pipeline of the hydrolysis section through a one-way valve on the pipeline.
  • the one-way valve is a stainless steel ferrule one-way valve
  • the material of the internal flow channel is polytetrafluoroethylene
  • the temperature of the cryocooler is 0-20°C.
  • the reaction temperature of the tubular reactor is 30-40°C.
  • the liquid separator includes a container, a stirring device, a feed port, a discharge port and an exhaust port, and the stirring device extends into the container and is connected to an external reducer for driving,
  • the discharge port is connected to the bottom of the container, and the feed port and the exhaust port are connected to the upper end of the container;
  • the outlet of the reactor and the exhaust port are connected to a suction pump through a pipeline; control valves are provided at the feed inlet, the material outlet and the exhaust port.
  • the exhaust pump when the hydrolysis reaction is carried out in the tubular reactor in the hydrolysis section, the exhaust pump simultaneously extracts the waste gas generated.
  • the method for continuously synthesizing acyl naphthalene also includes step S3: acylation wastewater treatment; specifically includes the following steps
  • the method for adjusting the pH of the aqueous phase separated by the liquid separator to alkaline comprises: adjusting the pH of the aqueous phase to 8-10 with an alkali raw material, so The alkali raw material is one or more of sodium hydroxide, potassium hydroxide or liquid ammonia.
  • adjusting the pH of the filtrate to acidity includes: adjusting the pH of the filtrate to 2-3 with an acid, where the acid is one or more of hydrochloric acid, nitric acid or sulfuric acid.
  • the extractant is a non-polar organic solvent, and the volume ratio of the extractant to the filtrate is 0.5-5:1.
  • the extractant is at least one of n-heptane, n-octane, n-hexane, benzene, toluene, xylene, and carbon tetrachloride.
  • step (3) the biochemical treatment adopts activated sludge method; in step (5), the temperature of heating and dissolving is 40-50°C, and the weight ratio of filter cake to concentrated hydrochloric acid is 0.5-2.5:1 .
  • the additive in step (5), is added in an amount of 2-10 wt% of the weight of the filter cake, and the additive is calcium aluminate or/and magnesium aluminate.
  • rectification includes the steps of:
  • the oil phase is fed into the first rectification tower from the middle part of the first rectification tower for rectification, the light fraction is steamed from the top of the tower, recovered after condensation, and the tower bottom liquid flows out from the bottom of the tower;
  • the pressure of the first rectification column is 0.05KPa-210KPa; the reflux ratio is (1-2):1.
  • the condensation temperature of the first rectification column is 10-220°C.
  • the pressure of the second rectification column is 0.05KPa-210KPa; the reflux ratio is (5-210):1.
  • the condensation temperature of the second rectification column is 50-290°C.
  • the 2-methylnaphthalene is extracted from washing oil with a purity of 99.0-99.9%; wherein the extraction method is:
  • wash oil obtains methylnaphthalene enrichment fraction by rectification separation
  • the acylation solution prepared in the examples of the present application has good homogeneity and high stability and can be used in a microchannel reactor, and when the acylation solution is used for acylation to prepare 2-methyl-6-propionylnaphthalene,
  • the acylation reaction liquid is immediately connected to the hydrolysis section, and the continuous and stable synchronous hydrolysis improves the material balance rate and there is no risk of pipeline blockage.
  • the hydrolyzed liquid can directly enter the liquid separator for synchronous liquid separation, and the oil phase can be conveniently collected , the exhaust gas generated can be drawn out from the upper exhaust port of the liquid separator.
  • the pH value of the oil phase produced by synchronous hydrolysis is 6-7, which meets the requirement of purification.
  • FIG. 1 is a schematic structural diagram of a configuration system provided by an embodiment of the present application.
  • Fig. 2 is the internal structure schematic diagram of kettle body, filtering device and liquid storage tank place in Fig. 1;
  • Fig. 3 is the structural representation of the structural representation of the pH meter in Fig. 2;
  • Fig. 4 is a schematic structural view of a welding system provided by an embodiment of the present application.
  • Fig. 5 is a schematic structural diagram of a punching system provided by an embodiment of the present application.
  • Figure 6 is a rectification process flow diagram provided according to an embodiment of the present application.
  • Fig. 7 is a flow chart of the treatment process of the water phase separated by the liquid separator that can recover resources according to the embodiment of the present application;
  • Fig. 8 is a conventional activated sludge treatment system diagram
  • Fig. 9 is a change diagram of COD in the aqueous phase separated by the liquid separator when utilizing the treatment process of Fig. 7 in Example 8 of the present application;
  • Fig. 10 is a change diagram of COD in the aqueous phase separated by the liquid separator when utilizing the treatment process of Fig. 7 in Example 9 of the present application;
  • Fig. 11 is a change diagram of COD in the aqueous phase separated by the liquid separator when utilizing the treatment process of Fig. 7 in Example 10 of the present application;
  • Fig. 12 is the process flow chart of extracting high-purity 2-methylnaphthalene from wash oil in the embodiment of the present application;
  • Fig. 13 is a structural schematic diagram of the rectification column of the azeotropic rectification tower in Fig. 12;
  • Fig. 14 is the structural representation of the intermittent melting crystallizer among Fig. 12;
  • Figure 15 is a schematic structural view of the crystallization plate in Figure 12 of the present application.
  • Kettle body 101, first feeding port; 102, second feeding port; 103, stirring port; 104, liquid outlet; 1041, first control valve; 105, temperature measuring port; 106, pH meter port; 107 , the first shell; 108, the upper cover; 109, the jacket; 1091, the first inlet of the heat exchange medium; 1092, the first outlet of the heat exchange medium; 110, the first chamber;
  • Feeding pump 600. The first hose;
  • Filtration device 701, second housing; 7011, filter device inlet; 7012, filter device outlet; 7013, first part; 7014, second part; 702, filter membrane;
  • Liquid storage tank 801, the third shell; 8011, the inlet of the liquid storage tank; 8012, the outlet of the liquid storage tank; 8013, the second vacuum port;
  • Vacuum pump 900, the second hose; 901, the first vacuum port;
  • Liquid separator 1601, exhaust port; 1602, feed port; 1603, control valve; 1604, container; 1605, discharge port;
  • 1801 washing oil; 1802, naphthalene distillate light oil; 1803, methylnaphthalene enriched distillate; 1804, heavy distillate; 1805, azeotropic distillate; 1806, residual oil; 1807, entrainer mixture; 1808 1809, water; 1810, entrainer; 1811, heating and cooling medium; 1812, residual mother liquor; 1813, methylnaphthalene product; 1814, temperature control medium outlet; 1815, temperature control medium inlet; 1816, Material inlet; 1817, material outlet; 1818, crystal plate temperature control medium inlet; 1819, crystal plate temperature control medium outlet; 1820, protrusion; 1821, groove; 1822, spiral groove;
  • V1 washing oil storage tank; V2, atmospheric rectification tower; V3, azeotropic rectification tower; V4, ultrasonic static mixer; V5, separator; V6, rectification tower; V7, digital temperature-controllable oil bath ; V8, intermittent melting crystallizer.
  • the configuration system 100 for preparing an acylation liquid includes a kettle body 1 , a feeding port cover 2 , a feeding pump 6 and an agitator 3 .
  • the kettle body 1 defines a first chamber 110 , and the kettle body 1 has a first feeding port 101 , a second feeding port 102 , a stirring port 103 and a liquid outlet 104 communicating with the first chamber 110 .
  • each of the first feeding port 101 , the second feeding port 102 and the stirring port 103 is disposed above the liquid outlet 104 .
  • the solvent and Lewis catalyst used to prepare the acylating solution can be fed into the first chamber 110 through the first feed port 101 , and the acylating agent can be fed into the first chamber 110 through the second feed port 102 .
  • the feed port cover 2 is detachably sealed and mounted on the first feed port 101 .
  • the feed port cover 2 When it is necessary to feed materials (solvent and Lewis catalyst) into the kettle body 1, remove the feed port cover 2 from the first feed port 101, and after the feeding is completed, install the seal cover of the feed port cover 2 on the first feed port in time 101, thereby reducing the contact time of the solvent and Lewis catalyst with the external environment, and at the same time reducing the time that the first chamber 110 is exposed to the external environment through the first feeding port 101.
  • the feeding pump 6 is sealed and connected with the second feeding port 102 so as to add the acylating agent into the second feeding port 102 .
  • the acylating agent can be effectively prevented from contacting the external environment, and at the same time, the first chamber 110 is prevented from being exposed to the outside world through the second feeding port 102 Environment.
  • the stirrer 3 includes a stirring shaft 301 and a stirring blade 302, the stirring blade 302 is arranged on the stirring shaft 301, at least a part of the stirring shaft 301 and each of the stirring blades 302 or disposed in the first chamber 110.
  • the materials in the first chamber 110 can be fully stirred and mixed by using the agitator 3 to fully dissolve the Lewis catalyst and improve the preparation efficiency of the acylation liquid; and a part of the agitator 3 is sealed and inserted in the stirring port 103 , so that during the working process of the stirrer 3 , the stirring port 103 is always kept sealed, so as to prevent the first chamber 110 from being exposed to the external environment through the stirring port 103 .
  • the method for preparing an acylation solution implemented by the configuration system 100 for preparing an acylation solution according to the embodiment of the present application includes the following steps:
  • step b Add the solvent and the Lewis catalyst weighed in step a into the first chamber 110 through the first feeding port 101, and stir with the stirrer 3, so that the Lewis catalyst weighed in step a dissolves to obtain a mixed solution;
  • step c Add an acylating agent to the mixed solution obtained in step b by using the feeding pump 6, and stir by using the stirrer 3, so as to obtain an acylating liquid.
  • the method for preparing the acylation liquid in the embodiment of the present application when weighing the Lewis catalyst, weighs under the protection of an inert gas, which effectively ensures the activity of the Lewis catalyst, is not affected by the weather environment, and avoids high humidity in summer During the formation of acid mist; in the method for preparing the acylation liquid in the embodiment of the present application, after weighing the Lewis catalyst in an inert environment, add it to the solvent, and cover the feeding port cover 2, the liquid seal reduces the The contact time between the Lewis catalyst and the air can effectively reduce or even contact the Lewis catalyst and the acylating agent with moisture during the preparation of the acylation solution, thereby making the prepared acylation solution better in stability and uniformity; In the method for preparing the acylation solution in the embodiment of the present application, due to the agitation of the agitator 3, the Lewis catalyst can be fully and quickly dissolved in the solvent, which is beneficial to shorten the preparation time of the acylation solution and improve the acylation solution. Preparation efficiency: the acylation liquid
  • the Lewis catalyst is selected from at least one of AlCl 3 , BF 3 , ZnCl 2 or FeCl 3 .
  • the Lewis catalyst is not particularly limited, and the Lewis catalyst that can be used to synthesize 2-methyl-6-propionylnaphthalene can be prepared by the method for preparing the acylating liquid in the embodiment of the present application. preparation.
  • the acylating agent may be at least one selected from the group consisting of acetylating agents, propionylating agents and butyrylating agents.
  • the second feeding port 102 is a grinding port, and the second feeding port 102 is sealed and connected with the pagoda joint, and the discharge pipe of the peristaltic pump is sealed and connected with the pagoda joint.
  • the feeding pump 6 is a peristaltic pump.
  • the flow rate of the acylating agent added to the first chamber 110 of the still body 1 can be controlled by adjusting the flow rate of the peristaltic pump, so as to facilitate the control of the amount of the acylating agent added to the first chamber 110. It is beneficial to further improve the stability and uniformity of the prepared acylation solution.
  • the feeding pump 6 is connected to the second feeding port 102 through a first hose 600 .
  • the first hose 600 may be a polytetrafluoroethylene tube.
  • the stirring shaft 301 and the stirring blade 302 are made of metal
  • the stirrer 3 includes a shaft anti-corrosion layer and a blade anti-corrosion layer
  • the shaft anti-corrosion layer is coated on a part of the stirring shaft 301
  • the material of the stirring shaft 301 and the stirring blade 302 is stainless steel, and the material of the anti-corrosion layer of the shaft and the anti-corrosion layer of the blade is polytetrafluoroethylene.
  • the stirring shaft 301 and the stirring blade 302 are made of metal, which can effectively ensure that the stirring shaft 301 and the stirring blade 302 have sufficient structural strength. Coating the shaft anti-corrosion layer on the stirring shaft 301 and coating the blade anti-corrosion layer on the stirring blade 302 can prevent the metal material part of the stirring shaft 301 and the stirring blade 302 from contacting with solvent, Lewis catalyst and acylating agent, not only Corrosion of the stirring shaft 301 and the stirring blade 302 can be avoided, and corrosion products can be prevented from entering the acylation liquid and affecting the quality of the prepared acylation liquid.
  • the adding speed of the acylating agent is 3-10 drops per second, and the stirring speed of the agitator 3 is 200rpm-400rpm.
  • the kettle body 1 includes a first housing 107 , an upper cover 108 and a jacket 109 , the liquid outlet 104 is arranged on the first housing 107 , and the upper end of the first housing 107 wide open.
  • the upper cover 108 sealing cover is installed on the first housing 107, the first housing 107 and the upper cover 108 define the first chamber 110, each of the first feeding port 101, the second feeding port 102 and the stirring port 103 Or set on the upper cover 108.
  • the jacket 109 is sleeved on the first shell 107, and the jacket 109 has a first heat exchange medium inlet 1091 for the heat exchange medium to enter and a heat exchange medium first outlet 1092 for the heat exchange medium to flow out.
  • the first shell 107 and the upper cover 108 can be processed separately, which facilitates the processing of the kettle body 1 .
  • the heat exchange medium can flow into the jacket 109 through the first heat exchange medium inlet 1091, and exchange heat with the material in the first chamber 110 through the first shell 107, and then the heat exchange medium The heat exchange medium flows out of the jacket 109 through the first outlet 1092 .
  • the heating of the material in the first chamber 110 is realized, so that the temperature in the first chamber 110 is maintained at a preset temperature suitable for the Lewis catalyst solvent, and the Lewis catalyst is quickly dissolved in the solvent, thereby greatly shortening the temperature of the acylation liquid.
  • the preparation time is improved, and the preparation efficiency of the acylation solution is improved.
  • the heat exchange medium can be water, oil and other liquids.
  • the solvent is nitrobenzene
  • the above step b is:
  • nitrobenzene is added into the first chamber 110 through the first charging port 101;
  • step a the Lewis catalyst weighed by step a is added in the nitrobenzene;
  • the temperature in the first chamber 110 is kept at the preset temperature suitable for dissolving the Lewis catalyst, and the stirring speed of the stirrer 3 is also maintained at the stirring speed suitable for dissolving the Lewis catalyst, so that the Lewis catalyst is rapidly dissolved in the nitrobenzene In this way, the preparation time of the acylation solution is greatly shortened, and the preparation efficiency of the acylation solution is improved.
  • the acylation liquid preparation system further includes a heating device, a temperature sensor 5 and a controller.
  • the heating device has a second heat exchange medium inlet and a second heat exchange medium outlet, the first heat exchange medium outlet 1092 is connected to the second heat exchange medium inlet, and the first heat exchange medium inlet 1091 is connected to the second heat exchange medium outlet.
  • the kettle body 1 has a temperature measuring port 105 communicating with the first chamber 110 , a part of the temperature sensor 5 is sealed and inserted in the temperature measuring port 105 , and the detection end of the temperature sensor 5 is set in the first chamber 110 .
  • the controller is connected to each of the heating device and the temperature sensor 5 so that the controller controls the heating device according to the temperature detected by the temperature sensor 5 .
  • the heat exchange medium heated by the heating device flows to the first heat exchange medium inlet 1091 through the second heat exchange medium outlet, and the heat exchange medium flows into the jacket 109 through the first heat exchange medium inlet 1091, and the jacket
  • the heat exchange medium in 109 heats the material in the first chamber 110; then the heat exchange medium in the jacket 109 flows out of the jacket 109 through the first heat exchange medium outlet 1092, and flows back through the second heat exchange medium inlet
  • the heating device is heated by the heating device to realize the circulation of the heat exchange medium between the heating device and the kettle body 1 .
  • the temperature sensor 5 can be used to detect the temperature of the material in the first chamber 110 in real time, and the temperature of the material in the first chamber 110 detected by the temperature sensor 5 is transmitted to the controller, so that the controller can control the heating device.
  • the controller controls the heating device to stop heating, so as to prevent the material in the first chamber 110 from being higher than the preset temperature;
  • the controller controls the heating device to start heating, so as to prevent the material in the first chamber 110 from falling below the preset temperature.
  • the temperature in the first chamber 110 is kept at a preset temperature suitable for the dissolution of the Lewis catalyst, which is beneficial to increase the dissolution rate of the Lewis catalyst and improve the preparation efficiency of the acylation solution.
  • the model of the controller is DSC350.
  • the temperature of the mixed solution obtained in the above step b is not higher than 60°C.
  • the Lewis catalyst is dissolved in the solvent in step b, it is not necessary to lower the temperature of the mixed solution, and directly add the acylating agent to prepare the acylating liquid, which not only reduces the energy Consumption and shorten the preparation time of the acylation solution, improve the preparation efficiency.
  • the acylating agent can also be fed into the jacket 109 through the first heat exchange medium inlet 1091 while using the feed pump 6 to add the acylating agent to the first chamber 110.
  • the heat exchange medium for example, cooling water or coolant
  • the configuration system 100 for preparing the acylating liquid further includes a pH meter 4, the kettle body 1 has a pH meter port 106 communicating with the first chamber 110, and a part of the pH meter 4 is sealed and inserted In the pH meter port 106 , the detection end of the pH meter 4 is arranged in the first chamber 110 .
  • the pH meter 4 can be used to detect the pH of the material in the first chamber 110 in real time, which is beneficial to improving the preparation efficiency of the acylation solution and improving the quality of the prepared acylation solution.
  • pH meter port 106 is a ground port. Thus, the sealing of the pH meter port 106 is facilitated.
  • a glass cover 10 is provided outside the pH meter 4 , and the glass cover 10 has a tapered mating surface 1001 , and the mating surface 1001 is sealed with the pH meter port 106 .
  • the model of the pH meter 4 is SIN-PH6.3-5022-AL/Y.
  • a suction filtration device is further included, and the suction filtration device includes a filtration device 7 , a liquid storage tank 8 and a vacuum pump 9 .
  • the filter device 7 includes a second housing 701 and a filter membrane 702, the second housing 701 defines a second chamber, the filter membrane 702 is arranged in the second chamber, and the filter membrane 702 divides the second chamber into a first part 7013 and a second chamber.
  • the second part 7014, the second housing 701 is provided with a filter device inlet 7011 and a filter device outlet 7012 communicating with the second chamber, the first part 7013 is arranged adjacent to the filter device inlet 7011, and the second part 7014 is arranged adjacent to the filter device outlet 7012 , the filter device inlet 7011 is connected with the liquid outlet 104 .
  • the liquid storage tank 8 includes a third housing 801, the third housing 801 defines a third chamber, the third housing 801 is provided with a liquid storage tank inlet 8011 and a liquid storage tank outlet 8012, the liquid storage tank inlet 8011 is connected with the filter
  • the device outlet 7012 communicates.
  • the vacuum pump 9 has a first vacuum port 901 , the third housing 801 is provided with a second vacuum port 8013 communicating with the third chamber, and the first vacuum port 901 communicates with the second vacuum port 8013 .
  • the method for preparing the acylation liquid further includes step d, performing suction filtration on the acylation liquid obtained in the above step c under an inert gas atmosphere to remove solid particles.
  • the prepared acylation solution is treated by suction filtration under an inert atmosphere to remove all undissolved solid particles in the solution, thereby further improving the homogeneity of the solution.
  • the vacuum pump 9 makes the filter device 7 form a negative pressure, so that the acylation liquid prepared by the kettle body 1 is quickly passed through the filter membrane 702 of the filter device 7 under the action of the negative pressure, and flows into the liquid storage tank 8 for storage. , so that the overall preparation efficiency of the acylation solution can be further improved.
  • the first vacuum port 901 and the second vacuum port 8013 are connected through a second hose 900 .
  • the first feeding port 101 can be connected to the inert gas source, and the first feeding port 101 can be used to fill the acylating liquid preparation system with an inert gas, so that the prepared The acylation solution is subjected to suction filtration under an inert atmosphere.
  • an inert gas is introduced into the first housing 107, the second housing 701 and the third housing 801 through the first feed port 101, so that the acylating liquid The whole preparation process is carried out under an inert atmosphere.
  • the first housing 107 , the second housing 701 and the third housing 801 are integrated.
  • the liquid outlet 104 is arranged at the bottom of the first casing 107
  • the inlet 7011 of the filter device is arranged at the top of the second casing 701
  • the filter device 7 is arranged at the bottom of the kettle body 1 .
  • the outlet 7012 of the filter device is arranged at the bottom of the second casing 701
  • the inlet 8011 of the liquid storage tank is arranged at the top of the third casing 801
  • the liquid storage tank 8 is arranged at the bottom of the filter device 7 .
  • the acylation liquid prepared in the kettle body 1 can directly flow out of the kettle body 1 through the liquid outlet 104 by its own gravity, and enter the filter device 7 through the filter device inlet 7011, without the need for a liquid outlet 104 and a filter device.
  • a liquid pump for pumping acylation liquid is additionally provided between the inlet 7011 and between the outlet 7012 of the filter device and the inlet 8011 of the liquid storage tank. It is beneficial to simplify the overall structure of the dispensing system 100 for preparing the acylation liquid, and reduce the manufacturing and operating costs of the dispensing system 100 for preparing the acylation liquid.
  • the connection between the liquid outlet 104 and the inlet 7011 of the filter device, the outlet 7012 of the filter device and the inlet 8011 of the liquid storage tank can be omitted, which is convenient for preparing the acylation liquid. Assembly of the liquid configuration system 100.
  • the material of the first shell 107 , the jacket 109 , the second shell 701 and the third shell 801 is high borosilicate glass.
  • the material of the first shell 107 , the jacket 109 , the second shell 701 and the third shell 801 is high borosilicate glass.
  • the configuration system 100 for preparing the acylation liquid further includes a first control valve 1041 , and the first control valve 1041 is arranged on the liquid outlet 104 so as to control the opening and closing of the liquid outlet 104 .
  • the prepared acylation liquid flows out of the kettle body 1 through the liquid outlet 104, and when the first control valve 1041 is closed, the kettle body 1 is sealed, and the prepared acylation liquid
  • the chemical solution is stored in the still body 1.
  • the configuration system 100 for preparing the acylating liquid further includes a second control valve, the second control valve is arranged on the outlet 8012 of the liquid storage tank, so as to control the opening and closing of the outlet 8012 of the liquid storage tank.
  • the filtered acylation liquid flows out of the liquid storage tank 8 through the liquid storage tank outlet 8012, and when the second control valve is closed, the liquid storage tank 8 is sealed, and the filtered acylation liquid is stored. In the liquid storage tank 8.
  • the molar mass ratio of the acylating agent, the Lewis catalyst to the solvent is (1.1-1.5):(1.3-1.7):5; and/or the molar mass ratio of the Lewis catalyst to the solvent is (1.3-1.7) :5.
  • the molar mass of acylating agent, Lewis catalyst and solvent satisfies: the molar mass ratio of acylating agent, Lewis catalyst and solvent is (1.1-1.5):(1.3-1.7):5, and the molar mass of Lewis catalyst and solvent Ratio is (1.3-1.7): 5; Or, the molar mass ratio of acylating agent, Lewis catalyst and solvent is (1.1-1.5): (1.3-1.7): 5; Or, the molar mass ratio of Lewis catalyst and solvent is (1.3-1.7): 5.
  • the proportion of each substance is optimized, which can make full use of the raw materials and reduce the production cost.
  • the configured acylation solution is continuously synthesized acyl naphthalene, which includes,
  • Preparation of raw material solution 14 Take 5 L of acylation liquid as an example, add 700 g of nitrobenzene and 284 g of 2-methylnaphthalene to the 5 L acylation liquid in a reactor equipped with a stirring device at room temperature to prepare the raw material Liquid 14.
  • 2-methylnaphthalene is extracted from washing oil, and the purity is 99.0-99.9%; wherein the method for extracting 2-methylnaphthalene from washing oil is:
  • wash oil obtains methylnaphthalene enrichment fraction by rectification separation
  • the methylnaphthalene enriched fraction is passed into the azeotropic rectification column to carry out azeotropic rectification to obtain the azeotropic distillate;
  • step I washing the washing oil is rectified and separated through an atmospheric rectification tower to obtain naphthalene fraction light oil, methylnaphthalene-enriched fraction and heavy distillate oil, wherein a part of the heavy distillate oil Circulate back to the atmospheric distillation column, and another part is recycled.
  • the light oil of naphthalene fraction is obtained from the top of the atmospheric distillation tower, the enriched fraction of methylnaphthalene is obtained from the side line of the atmospheric distillation tower, and the heavy distillate is obtained from the bottom of the atmospheric distillation tower.
  • the temperature is 210-225°C, the temperature of the side line is 235-260°C, and the temperature at the bottom of the tower is 290-310°C.
  • methylnaphthalene-enriched fraction and entrainer are mixed in proportion, and the mixture is heated to a certain temperature and then enters the azeotropic rectification column for azeotropic rectification.
  • the entrainer is a single compound entrainer or a mixture type entrainer, wherein the single compound entrainer is any one of ethylene glycol, diethylene glycol, ethanolamine, diethylene glycol, and N-methylformamide , the mixture type entrainer is a mixture of heptane and ethanolamine or a mixture of heptane and ethylene glycol;
  • the top temperature of the azeotropic distillation column is 150-175 ° C, the top pressure is 1-4KPa, and the reflux ratio is 5 -15;
  • the temperature at the bottom of the tower is 220-245°C, and the pressure at the bottom of the tower is 10-15KPa.
  • the mass ratio of the methylnaphthalene-rich fraction mixed with the entrainer is 1-3:1; in response to the use of a mixture type entrainer, the heptane and ethanolamine or heptane in the mixture type entrainer
  • the mass ratio of alkane to ethylene glycol is 0.2-1:1
  • the mass ratio of the methylnaphthalene enriched fraction to the entrainer is 1:0.8-2.
  • the azeotropic distillate and water are passed into an ultrasonic static mixer for ultrasonic mixing, and then passed into a separator for standing separation of oil and water.
  • the operating temperature of the separator is 50-80 ° C, and the standing time is 0.2 -1 hour, separate the water phase and the oil phase from the separator, wherein the water phase is the entrainer and water, the water phase enters the rectification tower for the distillation separation of the entrainer and water, and the separated entrainer returns to the step II, mixed with the enriched fraction of methylnaphthalene, and recycled; the separated water returns and mixed with the azeotropic distillate, recycled; the oil phase is 2-methylnaphthalene crude product, and the oil phase enters the batch melting crystallizer for 2 - Crystalline purification of methylnaphthalene.
  • step IV there are two batch melting crystallizers, the initial temperature during the crystal growth process is 35-40°C, the cooling rate is 3-8°C/h, the final temperature is 8-12°C, and the constant temperature time is 0.5-1h;
  • the heating rate of the sweating process is 2-6°C/h, and the final temperature is 30-32°C; the heating rate of the melting process is 2-8°C/h, the final temperature is 45-50°C, and the final temperature constant temperature time is 0.5-1h.
  • the method for extracting 2-methylnaphthalene from the washing oil is as follows: the washing oil 1801 is added to the washing oil storage tank V1, and after stirring and mixing, the washing oil 1801 is sent to the atmospheric distillation tower with a pump V2, the washing oil 1801 is separated by rectification, and the naphthalene fraction light oil 1802 is obtained at the top of the tower, which can be used as a raw material for extracting naphthalene, or can be recycled back to the atmospheric rectification tower V2;
  • the produced product is the heavy oil fraction 1804, a part of the heavy oil fraction 1804 is recycled to the atmospheric distillation column V2, and the other part can be sent to the coal tar processing plant to further separate acenaphthene, oxyfluorene, industrial fluorene and other products; atmospheric pressure
  • the temperature at the top of the distillation column V2 is 210-225°C, the temperature at the side line is 235-260°C, and the temperature at the bottom of the tower is 290-
  • the methylnaphthalene-enriched fraction 1803 obtained above is mixed with the entrainer 1810 according to a certain ratio, and the mixture is heated to a certain temperature and then enters the azeotropic distillation tower V3 for azeotropic distillation.
  • the azeotropic distillate 1805 obtained at the top of the azeotropic distillation column V3 enters the separator to recycle the entrainer 1810, and the residual oil 1806 is obtained at the bottom of the tower, and the residual oil 1806 is thrown out as a means of extracting other fine chemicals. raw material.
  • the application adopts the method of water extraction to reclaim the entrainer 1810.
  • the azeotropic distillate 1805 obtained in step (2) is mixed with the water 1809, the mixture is first mixed in the ultrasonic static mixer V4 Ultrasonic mixing in the middle, and then into the separator V5 for static separation of oil and water.
  • a water phase and an oil phase are separated from the separator, wherein the water phase is an entrainer mixture 1807, including an entrainer 1810 and water 1809, and the oil phase is crude 2-methylnaphthalene.
  • the water phase that comes out from separator V5 enters rectification column V6 and carries out the distillation separation of entrainer 1810 and water 1809, and the entrainer 1810 that obtains from the bottom of the tower after separation returns step (2) and goes and methyl naphthalene enrichment fraction 1803 Mixing and recycling; after separation, the water obtained from the top of the tower is returned and mixed with the azeotropic distillate 1805 for recycling.
  • the 2-methylnaphthalene crude product that comes out from the separator V5 is purified by melting and crystallization.
  • two batch melting crystallizers V8 are connected in parallel to alternately carry out the crystallization and purification of 2-methylnaphthalene to achieve the goal of continuous crystallization and purification.
  • the crude methylnaphthalene 1808 undergoes crystal growth, sweating, melting and other processes in the batch melting crystallizer V8, and finally achieves the purpose of crystallization and purification.
  • the crude methylnaphthalene product 1808 is purified by melt crystallization to obtain a high-purity 2-methylnaphthalene product 1813.
  • the liquid discharged during the crystal growth and sweating process is the residual mother liquor 1812, which can be used as a raw material for extracting 1-methylnaphthalene.
  • the purity of the 2-methylnaphthalene product 1813 obtained after crystallization and purification is 99.0-99.9%, and the product yield is 70%-90%.
  • the atmospheric rectification tower V2 is a packed tower, and the packing height is equivalent to a stainless steel packed tower with a theoretical plate number of 30-50 layers, and the heavy weight at the bottom of the tower Part of the oil fraction 1804 is recycled and mixed with the washing oil and then enters the atmospheric rectification tower V2, and part of it is thrown externally.
  • the ratio of external rejection and circulation is 2-5:1, and the 2-methylnaphthalene in the obtained methylnaphthalene enriched fraction 1803 is The content is 50-70%.
  • the azeotropic distillation column V3 adopts a precision fractionation column with concentric tubes, as shown in Figure 13.
  • the fractionation column is composed of two concentric tubes that have been extremelyly designed and calibrated. Efficient mass and heat transfer between the rising steam and the liquid film in the concentric annular gap, the material is glass or stainless steel, and the theoretical plate number of the fractionating column is 80-120.
  • the melting crystallizer is a full cube static crystallizer as shown in Figure 14. Its shape is a cube. It is provided with a temperature control medium outlet 1814, a temperature control medium inlet 1815, a material inlet 1816 and a material outlet 1817.
  • the melting crystallizer is equipped with Multiple groups of crystallization plates parallel to each other, as shown in Figure 15, each group of crystallization plates is arranged with spiral protrusions 1820, the crystal layer grows on the protrusions 1820 during crystallization, and is not easy to fall off; during the sweating process, the crystal layer grows from the crystallization
  • the growth surface of the plate falls off, enters the groove 1821 between the protrusions 1820, and keeps in contact with the heat exchange surface to prevent the crystal layer from easily falling off during the sweating process;
  • the temperature control medium flows in through the temperature control medium inlet 1818 of the crystallization plate, flows out through the temperature control medium outlet 1819 of the crystallization plate, circulates to the next group of crystallization plates, and finally flows out through the temperature control medium outlet 1814 of the melting crystallizer.
  • the crystal plate is made of stainless steel, plexiglass or other alloys.
  • Each intermittent melting crystallizer V8 is connected to a digital temperature-controllable oil bath V7, and a heating and cooling medium 1811 is transmitted between the digital temperature-controllable oil bath V7 and the intermittent melting crystallizer V8.
  • a change in the traditional process of utilizing benzene series compounds as raw materials in the related art and the improved process method of using coal-based naphthalene series compounds as raw materials is carried out with naphthalene and methylnaphthalene as raw materials for 2,6-
  • the method of synthesizing dimethylnaphthalene uses coal tar washing oil to extract high-purity 2-methylnaphthalene, which improves the production efficiency, product purity and yield in the separation and refining process of 2-methylnaphthalene; reduces the probability of pollution and equipment corrosion , so as to realize the synthesis of 2-methyl-6-propionylnaphthalene from the raw material 2-methylnaphthalene as the raw material liquid and the acylation liquid for acylation and hydrolysis, and achieve the realization of 2-methyl-6-propionylnaphthalene using a non-petroleum-based route. Purpose of industrial production of 6-propionylnaphthalene.
  • acylation liquid and raw material liquid 14 that will configure are carried out acylation reaction as shown in Figure 4:
  • the raw material liquid 14 and the acylation liquid 15 are injected into the three-way mixer 11 through a syringe for mixing, and after mixing, enter the microchannel reactor 12 for reaction.
  • the acylation reaction liquid flows out from the outlet of the microchannel reactor 12.
  • the tank reactor 13 of multi-still series stir and react in the tank reactor 13, pass through 2-4 tank reactors 13, the acylation reaction liquid in the tank reactor 13 of multi-still series
  • the total residence time is 50-80min, and the acylation reaction solution flowing out from the tank reactor 13 is hydrolyzed synchronously and rectified to obtain 2-methyl-6-propionylnaphthalene.
  • the tank reactor 13 adopts multi-tank series reactors to realize continuous reaction, and the flow rate is controlled by regulating valves, thereby controlling the residence time of materials in the tank reactor 13 .
  • Several microchannel reactors 12 are provided and connected in parallel with each other, and the flow rate of acylation liquid 15 and raw material liquid 14 is directly proportional to the number of parallel connections.
  • the three-way mixer 11 and the microchannel reactor 12 are placed in the first constant temperature tank, and the temperature is controlled at -5-0° C. In some embodiments, the three-way mixer 11 and the microchannel reactor 12 are controlled The temperature of the reactor is -5°C; the tank reactor 13 is placed in the second constant temperature tank, and the temperature is controlled at 30-50°C, and the temperature of the tank reactor 13 is controlled at 35°C in some embodiments.
  • the raw material liquid 14 of all parallel branch microchannel reactors 12 is transported by the same raw material liquid pump, and the acylation liquid 15 is transported by the same acylation liquid pump.
  • the parallel microchannel reactors 12 are directly connected to the tank reactors 13 with multiple tanks connected in series.
  • the three-way mixer 11 is a T-type mixer or a Y-type mixer with an inner diameter of 0.5mm-3mm.
  • the materials of the microchannel reactor 12 and the tank reactor 13 are both resistant to strong acid corrosion, and are completely sealed and isolated from air.
  • the inner diameter of the microchannel reactor 12 is 0.5-3.175 mm.
  • the acylating agent in the acylation liquid 15 is any one of propionyl chloride, acetyl chloride, acetic anhydride, and propionic anhydride.
  • the method of directly passing the acylation reaction liquid into the water stage for hydrolysis is: wherein in the hydrolysis section, the water injection pump is opened in advance, and the water (deionized water) is injected before the acylation reaction liquid is passed into the hydrolysis section.
  • the acylation reaction liquid passes through the one-way valve on the pipeline, and directly enters the oil phase pipeline of the hydrolysis section, and is mixed with water in a low-temperature cooling bath at 0-20°C.
  • the middle temperature is 0°C, and after rapid mixing, it enters the tubular reactor for hydrolysis reaction.
  • the reaction temperature of the tubular reactor is 30-40°C, in some embodiments, it is 30°C, and the inner diameter is 2-5mm.
  • the ultrasonic vibration is turned on to vibrate and stir the liquid in the tubular reactor, and the suction pump of the liquid separator 16 at the end is turned on. Since the inner diameter of the tubular reactor is 2-5mm, the liquid can travel slowly, and it can travel from the inlet to the outlet of the tubular reactor for about 5-10 minutes, and the hydrolysis reaction can fully occur.
  • the liquid separator 16 is used to collect the mixed liquid. It is carried out simultaneously with the acylation reaction in the reaction section.
  • the liquid separator 16 is a liquid-liquid separator. After the mixed liquid is stratified in the liquid separator 16, the oil phase can be discharged from the lower port, and the measured pH value of the oil phase is 6-7. The upper end of the liquid separator 16 is connected to the suction pump through the pipeline, and the waste gas (such as HCl) produced by hydrolysis can be sucked away, and enter the lye filling at the rear end to collect the waste gas without causing air pollution.
  • the waste gas such as HCl
  • liquid separator 16 is: as shown in Figure 5, comprises container 1604, agitator, feed port 1602, material outlet 1605 and exhaust port 1601, and agitator extends in container 1604 , the stirring device is connected and driven through an external reducer.
  • the stirring device is not fully drawn in the accompanying drawings of this case, because the stirring device is a conventional structure, which is not the focus of this case, and the stirring device is not used in this case.
  • the discharge port 1605 is connected to the bottom end of the container 1604, and the feed port 1602 and the exhaust port 1601 are jointly connected to the upper end of the container 1604, and the feed port 1602 is connected to the outlet of the tubular reactor of the hydrolysis section through a pipeline, and the exhaust port 1601
  • a suction pump is connected through a pipeline, and a control valve 1603 is provided at the feed port 1602 , the feed port 1605 and the exhaust port 1601 .
  • the control valve 1603 at the material outlet 1605 remains closed first, and then opens the control valve 1603 when the oil phase is ready to be discharged after stratification. Stratification is to separate the water and oil phases. The oil phase will be below the water. The oil phase will be discharged first by the discharge port 1605. After the oil phase is discharged, some water will inevitably be discharged. Later, it can be rectified Purify.
  • the space of the container 1604 of the liquid separator 16 is large enough, and the effective volume is 10L, so there is no need to worry that the inside of the liquid separator 16 will be filled. Wait till finally mixed liquid all flows in the liquid distributor 16 and also can not be filled.
  • the temperature resistance of the liquid distributor 16 is -80-200° C., and there is no worry that the liquid distributor 16 will be damaged due to temperature during use.
  • the one-way valve is a stainless steel sleeve one-way valve
  • the material of the internal flow channel is polytetrafluoroethylene, which has better anti-corrosion performance.
  • the flow channel in the ordinary one-way valve is made of ordinary rubber or rubber with corrosion resistance, but after testing, the corrosion resistance effect is not good, so it is replaced with polytetrafluoroethylene material, which greatly improves the anti-corrosion performance.
  • the method of the synchronous hydrolysis is to first turn on the water pump, and then immediately connect the acylation reaction liquid flowing out from the tank reactor 13 to the hydrolysis section, and continuously and stably hydrolyze simultaneously, so that the hydrolysis efficiency and material balance rate are greatly improved.
  • the acylation reaction is synchronized with the hydrolysis reaction, which avoids the temporary storage of the acylation reaction solution and reduces the overflow and pollution of HCl gas.
  • the hydrolysis reaction is an exothermic reaction, and should be mixed with water rapidly in a low-temperature cooling bath (0°C) before entering the tubular reactor, so that the hydrolysis temperature can be rapidly reduced.
  • a one-way valve is added in the front section of the pipeline where the acylation reaction liquid and water meet to prevent the liquid from flowing back from the hydrolysis section to the acylation section, avoiding the risk of channel blockage in the tubular reactor.
  • the hydrolyzed liquid can directly enter from the upper end of the liquid separator 16 for synchronous liquid separation, and the oil phase can be released from the lower port conveniently.
  • the measured pH value of the hydrolyzed oil phase is 6-7, meeting the requirement for purification.
  • the acylation reaction solution flowing out of the tubular reactor is immediately connected to the hydrolysis section, and the continuous and stable synchronous hydrolysis improves the material balance rate.
  • the most important thing is that there is no risk of blockage, and the hydrolyzed liquid can directly enter the hydrolysis section.
  • the liquid separator 16 the liquid is separated synchronously, and the oil phase can be easily discharged from the lower port for rectification and crystallization to obtain 2-methyl-6-propionylnaphthalene, and the waste gas generated can be exhausted from the upper end of the liquid separator 16 Port 1601 is drawn out; the water phase is processed through the treatment process of acylation wastewater.
  • the oil phase separated by the hydrolysis reaction is fed from the middle of the first rectification tower to the first rectification tower for rectification, the light fraction is steamed from the top of the tower, recovered after condensation, and the liquid in the bottom of the tower flows out from the bottom of the tower ;
  • the oil phase separated by the hydrolysis reaction is passed into the first rectification tower, wherein the first rectification tower is an acylation lightening tower, and the shape of the packing used is ceramics, springs, ⁇ rings or combinations thereof,
  • the light distillate at the top of the light removal tower is mainly a small amount of water that is condensed and then recycled into the advanced oxidation treatment of wastewater.
  • the side line of the light removal tower extracts nitrobenzene, which can be recycled and reused; the bottom liquid of the light removal tower is The crude product is sent to the second rectification column.
  • the pressure of the first rectifying column is 0.05KPa-10KPa, in some embodiments 0.1KPa-2KPa; the reflux ratio is (1-2):1.
  • the condensation temperature of the first rectification column is 10-20°C, in some embodiments 10-15°C.
  • the second rectification tower is a weight-removing tower, and the shape of the filler used is ceramics, springs, ⁇ rings or a combination thereof; the overhead distillation of the weight-removing tower removes 2-methyl-6-propionyl
  • the light components of isomers other than (2,6-MPN) are recovered after condensation; a small amount of heavier colored impurities are discharged from the bottom of the weight removal tower for regular collection and post-treatment, and the side line of the weight removal tower is extracted from the acylation
  • the 2,6-MPN fraction of the crude product is cooled and sent to the recrystallizer;
  • the acyl group of 2-methyl-6-acylnaphthalene in the isomers other than 2,6-MPN is methyl, ethyl , or one of isopropyl, that is, the product 2-methyl-6-acylnaphthalene is 2-methyl-6-formylnaphthalene, 2-methyl-6-acetylnaphthalene or 2-methyl-6 - Iso
  • the pressure of the second rectification column is 0.05KPa-10KPa, in some embodiments 0.1KPa-2KPa; the reflux ratio is (5-10):1.
  • the condensation temperature of the second rectification column is 50-90°C, in some embodiments 70-90°C.
  • the method for recrystallizing the crude product 2,6-MPN is as follows: methanol aqueous solution (wherein the mass ratio of methanol:water is 85:15) and the crude product 2,6-MPN by mass of 8:1 Ratio, mix in a three-necked flask equipped with stirring, thermometer, and reflux condenser, stir in a water bath at 55°C until the light yellow solid is completely dissolved, continue stirring for 20 minutes, crystallize the recrystallization solution at 10°C, and crystallize for 6 hours . After the crystallization is complete, carry out suction filtration and solid-liquid separation to obtain white fine powder 2-methyl-6-propionylnaphthalene.
  • the product is placed in a desiccator to remove the solvent, and the 2,6-MPN slurry obtained by recrystallization
  • the material is separated from the filter cake and filtrate in the product filter, the mother liquor and flushing liquid enter the methanol recovery tower, and the methanol is distilled from the top for reuse; the bottom distillate enters the sewage system; the wet filter cake is dried by a dryer, and the exhaust gas of the dryer enters Methanol recovery system, the product 2,6-MPN is obtained after drying.
  • the water phase separated in the liquid separator 16 is treated through the treatment process of acylation wastewater, and the specific process is as follows.
  • the resource recovery acylation wastewater treatment system of the embodiment of the present application includes: sedimentation tank 1702, filtration unit 1703, extraction unit 1704, rectification tower 1705, biochemical treatment unit 1706, dissolution tank 1707 and reaction Still 1708;
  • Settling tank 1702 is used to generate aluminum hydroxide precipitation;
  • the inlet of filter unit 1703 is connected with the outlet of settling tank 1702;
  • the inlet of extraction unit 1704 is connected with the filtrate outlet of filter unit 1703, and extraction unit 1704 is provided with extraction agent inlet; Rectification
  • the inlet of the tower 1705 is connected to the outlet of the extraction phase of the extraction unit 1704, the outlet of the extractant of the rectification tower 1705 is connected to the inlet of the extractant, and the outlet of the nitrobenzene of the rectification tower 1705 is connected to the organic solvent pipeline for production;
  • the inlet of the biochemical treatment unit 1706 is connected to the extraction phase.
  • the raffinate phase outlet of the unit 1704, the outlet of the biochemical treatment unit 1706 is connected to the waste water standard discharge pipeline; the outlet of the dissolution pool 1707 is connected to the filter cake outlet of the filtration unit 1703, and the dissolution pool 1707 is provided with a concentrated hydrochloric acid inlet; the inlet of the reaction kettle 1708 is connected to the dissolving Pool 1707 exit.
  • the extraction unit 1704 is a centrifugal extraction device, such as a commercially available centrifugal extraction machine.
  • the biochemical treatment unit 1706 may employ an activated sludge treatment system.
  • the activated sludge system is not the focus of this application, and a conventional activated sludge treatment system can be used.
  • a conventional activated sludge treatment system includes an aeration tank 17061, a secondary settling tank 17602, a return system, an excess sludge discharge system, an oxygen supply system, etc.
  • the extraction unit The raffinate phase outlet of 1704 is connected with the water inlet of aeration tank 17061, so that the raffinate phase can be introduced into the activated sludge treatment system, and treated according to the conventional activated sludge method to achieve standard discharge.
  • the communication mode between the various components can be realized through pipeline connection or transport by transfer vehicle according to the nature of the material.
  • valves and pumps can be installed on the corresponding pipelines as required, which are conventional technologies and not the focus of this application.
  • the acylation wastewater treatment system capable of resource recovery in the embodiment of the present application further includes a regulating tank 1701 arranged before the sedimentation tank 1702 .
  • the acylation wastewater advanced regulation tank 1701 adjusts the water quality, water quantity and pretreatment, and then enters the sedimentation tank 1702.
  • the reactor can be a stainless steel reactor or the like.
  • the sedimentation tank 1702 can be a horizontal flow sedimentation tank or a vertical flow sedimentation tank, etc.
  • the filter unit 1703 can be a plate and frame filter press.
  • the dissolution pool can be an ordinary dissolution pool
  • the rectification tower can be a packed tower or a tray tower.
  • the water phase When in use, the water phase first enters the regulating tank 1701 to adjust the water volume, balance the water quality and pre-treat it, and then enter the sedimentation tank 1702; the acylation wastewater entering the sedimentation tank 1702 passes through an alkali source, namely at least one of sodium hydroxide, potassium hydroxide or liquid ammonia. Adjust the pH to alkaline, that is, the pH is 8-10, in some embodiments the pH is around 9, the aluminum ions in the acylation wastewater are precipitated as aluminum hydroxide, and the suspension containing the aluminum hydroxide precipitate enters the filter unit 1703 for filtration , to obtain aluminum hydroxide filter cake and filtrate.
  • an alkali source namely at least one of sodium hydroxide, potassium hydroxide or liquid ammonia.
  • Adjust the pH to alkaline that is, the pH is 8-10, in some embodiments the pH is around 9
  • the aluminum ions in the acylation wastewater are precipitated as aluminum hydroxide
  • the filtrate adopts at least one of hydrochloric acid, nitric acid or sulfuric acid to adjust the pH of the filtrate to 2-3, then enters the extraction unit 1704, and under stirring conditions, is extracted and separated by the extractant in the extraction unit 1704 to form the extraction phase and
  • the raffinate phase wherein the extractant is a non-polar organic solvent, is at least one of n-heptane, n-octane, n-hexane, benzene, toluene, xylene, carbon tetrachloride, and the volume of the extractant and the filtrate The ratio is 0.5-5:1.
  • the extraction phase enters the rectification tower 1705 to realize the efficient separation of nitrobenzene and the extractant to obtain the raw material nitrobenzene and the extractant, and the extractant from the rectification tower 1705 enters the extraction unit 1704 again to participate in the extraction of the filtrate, and
  • the nitrobenzene from the rectification tower 1705 enters the organic solvent pipeline for production and is used for the production of related chemicals;
  • the raffinate phase from the extraction unit 1704 enters the biochemical treatment unit 1706 and is discharged up to standard after biochemical treatment.
  • the aluminum hydroxide filter cake coming out of the filter unit 1703 enters the dissolution tank 1707, and is heated and dissolved at 40-50°C under the action of concentrated hydrochloric acid in the dissolution tank 1707, wherein the weight ratio of the filter cake to the concentrated hydrochloric acid is 0.5-2.5:1, and then Enter the reaction kettle, and polymerize with the auxiliary agent added to the reaction kettle 1708 to obtain polyaluminum chloride, a water treatment agent.
  • the quality of the obtained polyaluminum chloride meets the national standard for polyaluminum chloride "GB/T22627-2014 Water treatment agent-polyaluminum chloride"
  • the auxiliary agent is calcium aluminate or/and magnesium aluminate, and its added amount is 2-10wt% of the weight of the filter cake.
  • the resource-recyclable acylation wastewater treatment process of the embodiment of the present application can use the resource-recoverable acylation wastewater treatment system of the embodiment of the application to treat the acylation wastewater and return its resources, but its realization
  • the system device used is not limited to the acylation wastewater treatment system that can recover resources in the embodiment of the present application.
  • the raw material reagents and equipment involved in the examples of this application are commercially available reagents and equipment; the detection methods involved in the examples of this application, etc., are all commercially available unless otherwise specified for the conventional method.
  • the suction filtration can adopt a laboratory suction filtration device assembled from a Buchner funnel, a suction filtration bottle, a rubber hose, an air suction pump, and filter paper.
  • the acylation liquid is subjected to suction filtration with a suction filtration device to remove solid particles in the acylation liquid to obtain an acylation liquid, which is stored in a storage tank for 30 days, and its solid content is less than 0.2%.
  • the reaction temperature is controlled at -5°C, and they flow into the tank reactor 13 with multiple tanks connected in series, and the tank reactor 13 is placed in the second In the second constant temperature tank, the control reaction temperature is 40°C, and the reaction is stirred in the tank reactor 13.
  • the total residence time through the three tank reactors 13 is 60min, and the acylation reaction solution flowing out from the tank reactor 13
  • the hydrolysis reaction is carried out with desalted water in a tubular reactor.
  • the hydrolysis reaction method is carried out in a tubular reactor; in the hydrolysis section, the water injection pump is turned on in advance, and the deionized water is passed into the water phase pipeline of the hydrolysis section before the acylation reaction solution is passed into the hydrolysis section to control the water
  • the flow rate of water in the phase line is 10mL/min.
  • the method for rectifying the oil phase obtained by hydrolysis reaction separation comprises the following steps, as shown in Figure 6:
  • the oil phase is fed into the first rectification tower from the middle part of the first rectification tower for rectification, containing 80wt% nitrobenzene solvent and 20wt% 2-methyl-6-propionyl in the acylation reaction liquid Naphthalene crude fraction.
  • the pressure of the first rectification tower is 0.1kPa, the rectification temperature is 45-130°C, nitrobenzene is distilled from the top of the tower at 45°C, the reflux ratio is 1:1, the condensation temperature at the top of the tower is 10°C, the gas phase at the top of the tower After nitrobenzene is condensed, the nitrobenzene solvent C is recovered from the top of the tower, and the temperature of the bottom liquid B is 130° C., which flows out from the bottom of the tower. After being rectified by the first rectifying tower, 98% of the nitrobenzene solvent C is recovered, and the bottom liquid B contains 78% of crude 2-methyl-6-propionyl naphthalene.
  • the discharge temperature of the still bottom liquid B in step (1) is 130° C., and it is pumped to the second rectification tower while it is hot for further rectification.
  • the pressure of the second rectification tower is 1kPa, the rectification temperature is 110-150°C, the reflux ratio is 5:1, the condensation temperature is 75°C, the isomerization of 17wt% 2-methyl-6-propionylnaphthalene
  • Body light fraction D is steamed from the top of the tower at 114°C and collected after condensation; 80wt% of the product E2-methyl-6-propionylnaphthalene is extracted from the second rectifying tower at 136°C; 3% of the tower
  • the still weight component F enters the bottom of the tower, is heated by the reboiler at the bottom of the tower, and then pumped to the tower still material collection tank while it is hot.
  • the theoretical plate number of the first rectifying tower is 30, and the feeding position is located at the 15th theoretical plate; the theoretical plate number of the second rectifying tower is 40, and the feeding position is located at the 25th theoretical plate, and the product measuring line adopts It is located at the 22nd theoretical board.
  • the method for recrystallizing the crude product 2,6-MPN is as follows: methanol aqueous solution (wherein the mass ratio of methanol: water is 85:15) and the crude product 2,6-MPN in a mass ratio of 8:1, in a stirring , a thermometer, and a reflux condenser in a three-necked flask, stirred in a 55°C water bath until the light yellow solid was completely dissolved, continued stirring for 20 minutes, and crystallized the recrystallization solution at 10°C for 6 hours. After the crystallization is complete, carry out suction filtration and solid-liquid separation to obtain white fine powder 2-methyl-6-propionylnaphthalene.
  • the product is placed in a desiccator to remove the solvent, and the 2,6-MPN slurry obtained by recrystallization
  • the material is separated from the filter cake and filtrate in the product filter, the mother liquor and flushing liquid enter the methanol recovery tower, and the methanol is distilled from the top for reuse; the bottom distillate enters the sewage system; the wet filter cake is dried by a dryer, and the exhaust gas of the dryer enters Methanol recovery system, the product 2,6-MPN is obtained after drying.
  • the acylation solution prepared in the method of Example 1 is mixed with the prepared 2-methylnaphthalene raw material solution 14, and the acylation reaction is carried out in a microchannel and still reactor, followed by hydrolysis and vacuum distillation 1. Obtain 2-methyl-6-propionylnaphthalene product after recrystallization, the yield is 80%, and the purity is 98.0%.
  • Example 2 Compared with Example 1, the data in Example 2 did not use a suction filtration device to carry out suction filtration treatment, and all undissolved solid particles in the acylation solution were removed. There is a risk of blocking the microchannel reactor 12 in the acylation reaction stage, thereby reducing the product yield of 2-methyl-6-propionylnaphthalene after recrystallization. In addition, the protective atmosphere and control parameters are also extremely important in the process of preparing the acylation liquid 15.
  • Example 1 and Example 4-7 in the table the influence of different acylation reaction conditions on the acylation reaction process is shown, wherein in Example 4-7 and Comparative Example 2-8, it is the same as that of Example 1 Compare, in the acylation process, respectively change the temperature that the three-way mixer 11 and the microchannel reactor 12 are placed in the first constant temperature tank, the kettle reactor 13 is placed in the second constant temperature tank to control the reaction temperature, and the stills connected in series
  • the number of formula reactor 13 and residence time and 2-methylnaphthalene: propionyl chloride: the ratio relationship of AlCl 3 thus draws that different acylation reaction conditions all influence the selectivity of 2-methylnaphthalene and obtain after recrystallization Product yield of 2-methyl-6-propionylnaphthalene.
  • the three-way mixer 11 and the microchannel reactor 12 are placed in the first constant temperature tank and the low-temperature mixing can not only remove the heat at a low temperature when the raw material liquid 14 and the acylation liquid 15 are mixed violently exothermic, but also make the acylation reaction at a low temperature
  • the microreactor zone is controlled by kinetics, and the reaction is fast, so that 60%-70% of 2-methylnaphthalene is converted into acylated products; React for a period of time in the device 13 so that the unreacted 2-methylnaphthalene will continue to react, improve the conversion rate of the reaction, and simultaneously reduce the cost of using the microchannel reactor 12 in industry.
  • the precise control of the temperature of the first constant temperature tank and the second constant temperature tank ensures a high conversion rate of the reaction and a high selectivity of the reaction.
  • embodiment 1 is compared with comparative example 2
  • changing the temperature of the first constant temperature tank can have a strong impact on the selectivity of 2-methylnaphthalene
  • embodiment 1 is compared with comparative example 3
  • changing the second constant temperature tank The temperature can seriously affect the selectivity of 2-methylnaphthalene and the product yield of 2-methyl-6-propionylnaphthalene.
  • Example 1 comparing Example 1 with Comparative Example 4, it can be seen that utilizing 5 tank reactors 13 and a total residence time of 90min, it is almost the same as the data of Example 1, but the experimental time is prolonged.
  • embodiment 4-7 and comparative example 8 it can be seen that in comparative example 8, the microchannel reactor 12 reaction that only utilizes internal diameter to be 3mm is not combined with tank reactor 13, causing the selectivity of 2-methylnaphthalene to be relatively high.
  • the microchannel reactor 12 has high specific surface area, high safety performance, fast temperature response, and is mainly based on diffusion and mass transfer, it can quickly remove heat from the reaction system and ensure that the reactants are at the optimum reactant temperature.
  • the reaction is carried out under the environment, which is conducive to the increase of the selectivity and yield of the reactant, but only using the microchannel reactor 12 causes the acylation reaction to take a long time, which is not conducive to the hydrolysis in the later stage.
  • the acylation reaction solution flowing out from the tank reactor in comparative example 9 was carried out intermittent tank type hydrolysis, which compared with Example 1, did not directly pass the acylation reaction solution flowing out from the tank reactor 13 into the pipe
  • the synchronous hydrolysis reaction is carried out directly in the type reactor, that is, after the acylation reaction liquid is obtained first, there is a certain time interval before the hydrolysis reaction is carried out, and water is directly added to the storage tank for hydrolysis.
  • the acylation reaction liquid cannot be hydrolyzed in time, and it is easy to place Hydrolysis reaction with water in the air will cause HCl gas to overflow, pollute the air, and is not conducive to the regulation of hydrolysis reaction time, temperature and material balance rate, and will increase water consumption by 2-3 times.
  • Example 10 compared with Example 1, the acylation reaction liquid and desalted water flowing out from the kettle reactor were hydrolyzed in a continuous synchronous hydrolysis reactor, and the acylation reaction was terminated and the obtained hydrolyzate was not used in Example 1.
  • the liquid separator in the liquid separator is used for liquid separation, and the oil phase and the water phase are separated after 24 hours of static separation, which not only requires a long liquid separation time of at least 24 hours, but also pollutes the environment with the exhaust gas generated during the liquid separation process.
  • Examples 8-12 exemplarily demonstrate that the hydrolyzate enters the liquid separator 16, and the aqueous phase solution after liquid separation is processed through the treatment process of acylation wastewater.
  • the pH of the filtrate was adjusted to 2 with 18wt% hydrochloric acid, and the filtrate was transferred into a separatory funnel filled with 500 mL of n-heptane, shaken and allowed to stand for 30 min to separate layers.
  • the supernatant (extract phase) is transferred to a packed column for rectification and recovery (the number of plates is 15, the reflux ratio is 0.06, the temperature at the top of the tower is 98°C, and the temperature at the bottom of the tower is 176°C), to obtain nitro
  • the main indicators of the obtained product polyaluminum chloride are: the mass fraction of aluminum chloride is 29%, the basicity is 55%, and it is insoluble in water The mass fraction of the substance is 0.2, the pH value is 4.5, the mass fraction of iron is 1.0%, and arsenic, lead, chromium, mercury and cadmium are not detected, and all indicators of the product polyaluminum chloride are in line with the national standard "GB/T22627 -Requirements of 2014 Water Treatment Agent - Polyaluminum Chloride.
  • the pH of the filtrate was adjusted to 2.5 with 50 wt% sulfuric acid, and the filtrate was transferred into a separatory funnel filled with 600 mL of n-heptane, shaken and allowed to stand for 30 min to separate layers.
  • the supernatant (extract phase) is transferred to a rectification tower for rectification and recovery (the number of plates is 15, the reflux ratio is 0.06, the temperature at the top of the tower is 98°C, and the temperature at the bottom of the tower is 176°C), to obtain nitro
  • the main indicators of the obtained product polyaluminum chloride are: the mass fraction of aluminum chloride is 30%, the basicity is 56%, and it is insoluble in water The mass fraction of the substance is 0.3, the pH value is 5.0, the mass fraction of iron is 1.0%, and arsenic, lead, chromium, mercury and cadmium are not detected, and all the indicators of the product polyaluminum chloride are in line with the national standard "GB/T22627 -Requirements of 2014 Water Treatment Agent - Polyaluminum Chloride.
  • the supernatant (extract phase) is transferred to a rectification tower for rectification and recovery (the number of plates is 15, the reflux ratio is 0.06, the temperature at the top of the tower is 98°C, and the temperature at the bottom of the tower is 176°C), to obtain nitro
  • the solution in the lower layer of the separatory funnel (raffinate phase) was transferred to the activated sludge treatment system for further treatment for 12 hours.
  • the main indicators of the obtained product polyaluminum chloride are: the mass fraction of aluminum chloride is 31%, the basicity is 62%, and it is insoluble in water The mass fraction of the substance is 0.4, the pH value is 6.5, the mass fraction of iron is 2.0%, arsenic, lead, chromium, mercury and cadmium are not detected, and all the indicators of the product polyaluminum chloride are in line with the national standard "GB/T22627 -Requirements of 2014 Water Treatment Agent - Polyaluminum Chloride.
  • the embodiment of the present application is basically the same as embodiment 18, except that the extractant is carbon tetrachloride.
  • the embodiment of the present application is basically the same as that of embodiment 18, except that the extractant is a mixture of n-octane and xylene in a volume ratio of 1:1.

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Abstract

一种酰化液及利用酰化液连续合成酰基萘的工艺方法，属于化工合成技术领域，包括利用配置系统(100)配置酰化液，并将酰化液与含2-甲基萘的原料液，按照2-甲基萘：酰化剂：Lewis催化剂为1：1.3：1.5混合形成酰化反应液，酰化反应液依次进入微通道反应器((12)和多釜串联的釜式反应器(13)中进行酰化反应；后进行水解、精馏及结晶，进而连续合成酰基萘。

Description

一种酰化液及利用酰化液连续合成酰基萘的工艺方法
相关申请的交叉引用
本申请基于申请号为202111075055.5、申请日为2021年09月14日的中国专利；申请号为202111093577.8、申请日为2021年09月17日的中国专利；申请号为202210095938.0、申请日为2022年01月26日的中国专利；申请号为202111638857.2、申请日为2021年12月29日的中国专利、申请号为202210530956.7、申请日为2022年05月16日的中国专利以及申请号为202210135569.3、申请日为2022年02月14日的中国专利申请提出，并要求该六项中国专利申请的优先权，将该六项中国专利申请的全部内容在此引入本申请作为参考。
技术领域
本申请涉及化工合成技术领域，具体而言，涉及一种酰化液及利用酰化液连续合成酰基萘的工艺方法。
背景技术
2,6-萘二甲酸是特种高端聚酯聚萘二甲酸乙二醇酯(PEN)的重要单体，且是合成多种高性能聚萘酯、聚氨酯以及液晶聚酯树脂的关键单体，特别是与乙二醇反应制得的聚萘二甲酸二乙酯(PEN)，在各方面物理化学性能均较目前广泛使用的聚对苯二甲酸二乙酯(PET)优越，在纤维、薄膜、包装容器和电子元件等领域有着广泛的应用前景。工业化生产2，6-萘二甲酸(2,6-NDCA)的方法可分为两类一类是以石油基的苯系化合物为原料的传统工艺方法，如BP-Amoco、日本三菱化学公司、Chevron及芬兰Optatech采用此法，此法步骤多，生产成本高；另一类是以煤基的萘系化合物为原料的改进工艺方法，以萘和甲基萘为原料，进行2,6-二甲基萘(2,6-DMN)的合成，一步制得DMN产品，再经异构化、分离提纯即可得到2,6-DMN，Exxon-Mobil公司采用此法，此法步骤简洁，成本相对较低。
其中工业化路线和半工业化路线为：
(1)邻二甲苯烷基化、环化、脱氢和异构化生产2,6-二甲基萘，再氧化生产萘二甲酸。此反应过程复杂，5步合成路线长，技术要求高，反应条件比较苛刻，总收率为36％左右。
(2)萘经液相烷基化生产2,6二异丙基萘，再氧化制2,6-萘二甲酸。由萘在酸处理过的丝光沸石催化作用下与丙烯反应，反应的混合物用精馏的方法得到粗二烷基萘(异构体的混合物)，再经过进一步分离得到2,6-二烷基萘；经空气液相氧化制得2,6-萘二甲酸。此种方法反应工艺步骤多，烷基化和提纯是该技术的关键。当前，2,6-NDCA的工业化合成路线主要有两条，一是2-甲基萘乙酰化后得到2-甲基-6-乙酰基萘，再进一步氧化制得2,6-NDCA。二是2,6-二烷基萘氧化制2,6-NDCA，包括2,6-二甲基萘、2,6-二乙基萘、2,6-二异丙基萘的液相催化氧化。前者虽然具有收率和选择性高，工艺相对简单的优点，但制备2-甲基-6-乙酰基萘所用催化剂比较昂贵，且难以回收，成本高、反应物毒性大、污染大，因此生产规模也较小。而后者通过萘的烷基化后，采用Co-Mn-Br催化剂体系，以低级脂肪酸为溶剂，在钛材反应器内反应，温度200℃左右，压力3.0MPa左右，其反应条件温和，且易于控制。
其中以2-甲基萘为原料经过酰基化反应生成2-甲基-6-丙酰基萘后，再氧化得到2,6-萘二甲酸，该合成方法原料来源丰富、副反应少、产物容易精制，同采用2,6-二烷基萘氧化制备2,6-萘二甲酸的合成路线相比要容易很多。而在以2-甲基萘为原料的酰基化反应中，高稳定性、高均一性的酰化液是非常关键的。酰基化反应中采用Lewis酸作为催化剂生成亲电络合物，容易进入被酰化物的对位，因此反应的选择性较高。但由于催化剂和酰化剂活性较高，易与水发生反应并变质，接触湿度较大的空气时会产生大量白色酸雾，同时催化剂反应后会生成沉淀，故酰化液在空气中较不稳定。另外，反应需要在均相中进行，溶液中未溶解的催化剂与变质生成的颗粒物均会影响酰化液反应的效率，影响最终产品的收率与纯度。
现有技术中利用2-甲基萘为原料，丙酰氯为酰化剂，三氯化铝为催化剂，硝基苯为溶剂，在常温常压下进行傅克酰基化反应制备2-甲基-6-丙酰基萘。其中酰基化反应结束后需要将反应水解淬灭，并将油相水解至pH为6-7，再经过减压精馏、精馏、重结晶等纯化方法得到纯度较高的2-甲基-6-丙酰基萘。其中该反应的反应温度具有一定的特点，两个物料需要在低温混合先反应，再将温度升高反应一段时间。低温混合的目的有两个，一是，两个物料混合时会剧烈放热，需低温移热；二是，低温区酰化反应受动力学控制，60％-70％的原料(2-甲基萘)已经反应。现已报道的酰基化反应采用间歇的釜式反应器或连续的微通道反应器，但是釜式反应器换热面积小，反应温度不易控制，停留时间不一致、产品质量不易稳定；而微通道反应器有堵塞的风险，且放大成本较高，操作较复杂。因此酰化反应采用间歇的釜式反应器的缺点是换热面积小，反应温度不易控制，反应时间较长，停留时间不一致，产品质量不易稳定，产物收率较低；酰化反应采用连续的微通道反应器的缺点是有堵塞的风险，放大成本较高，操作较复杂；该反应的反应温度具有一定的特点，酰化液和原料液两个物料需要在低温混合先反应，再将温度升高反应一段时间，现有装置未完全满足该反应特点的需要。
此外，该反应酰基化反应后，含有2-甲基-6-酰基萘的酰化反应液的水解过程对得到高纯度和高收率的2-甲基-6-酰基萘有着重要作用。具体是水解可以使酰化反应停止，水解掉三氯化铝和酰化剂，溶解反应中产生的HCl气体，使产生2-甲基-6-酰基萘保留在油相中。水解后油相的pH值提高至6-7，以消除酸性对产物的影响，在减压精馏升温提纯时，酰基在弱酸中氧化易发生缩合和沥青化，生成焦油，降低产物的收率和纯度。这就要求在水解酰化反应液时，做到同步连续水解为得到纯度较高的2-甲基-6-丙酰基萘做好保障。
现有技术中水解酰化反应液时，出现以下三种方案，第一种是先将酰化反应液流入混合器保存，和水混合后再进入后端的水解段的微通道反应器，无法做到酰化反应和水解反应同步进行，并且在酰化反应液与水接触时会放出大量的热，无法及时将温度冷却下来；第二种是将反应出来的产品通入水解反应器，将反应混合液倒入装有冰块的大烧杯中，倾倒过程中不断机械搅拌，倾倒完毕后加入精馏水继续搅拌半个小时使酰化产物水解完全。该方案是在间歇式水解反应，时间较长，操作较复杂，水解效果一般；第三种是将酰化反应液加入到乙醇水溶液中淬灭，分层将得到的硝基苯相减压精馏回收溶剂硝基苯，得到粗产物2-甲基-6-乙酰基萘。该方案是乙醇水溶液水解，并未突出其优势。
总结下来，现有技术有以下缺点：间歇或半连续式水解，没有做到同步连续水解，先得到酰基化反应液后，到进行水解反应前有一定的时间间隔；酰基化反应液不能及时水解，放置时易与空气中的水发 生水解反应，自身还会有HCl气体溢出，污染空气；水解反应会产生Al(OH)3乳状物，采用微通道反应器进行水解反应时，酰化反应液与水接触时会出现堵塞管路的情况；该水解反应是个放热反应，瞬间放出的热使水解段在管路中产生压力，同时会有少量的酸气挥发；水解得到的液体，体积较大，较浑浊，导致分液不易操作，时间较长。
最后，酰基化反应后废水主要为高毒性、高含盐的难生物降解的有机废水(表1)，极难于处理。可以说，在当前日趋严格的环境监管现状下，酰基化反应产生的难生物降解有机废水是限制Friedel-Crafts酰基化反应工业化的主要症结所在。
表1酰基化废水水质
Figure PCTCN2022115478-appb-000001
因此，研发针对酰基化反应产生的难生物降解有机废水的处理技术十分必要。
发明内容
本申请旨在至少在一定程度上解决相关技术中的技术问题之一，提出了一种酰化液及利用酰化液连续合成酰基萘的工艺方法，针对现有技术中酰化液不稳定且均一性差以及利用酰化液制备酰基萘反应效率低，损耗高，水解过程不稳定且易阻塞管路和酰基萘收率低纯度低的技术缺陷，提出一种新的酰化液，并利用酰化液提高合成酰基萘的反应速率以及水解稳定性，从而提高酰基萘的收率；此外本申请实施例同时对酰化液合成酰基萘后的水解废水进行处理，实现绿色环保生产。
有鉴于此，本申请实施例提出了一种酰化液，利用配置系统进行配置，所述配置系统包括：
釜体，所述釜体限定出第一腔室，所述釜体具有与所述第一腔室连通的第一加料口、第二加料口、搅拌口和出液口；
加料口盖，所述加料口盖可拆卸地密封盖装在所述第一加料口上；
加料泵，所述加料泵与所述第二加料口密封相连，以便向所述第二加料口内添加酰化剂；和
搅拌器，所述搅拌器的一部分密封插装在所述搅拌口内，所述搅拌器包括搅拌轴和搅拌叶片，所述搅拌叶片设置在所述搅拌轴上，所述搅拌轴的至少一部分和所述搅拌叶片中的每一者设置在所述第一腔室内；
利用配置系统配制酰化液包括以下步骤：
a：在惰性气体保护下称量Lewis催化剂；
b：将溶剂和步骤a称量的所述Lewis催化剂通过所述第一加料口加入所述第一腔室内并利用所述搅拌器搅拌均匀得到混合溶液；
c：利用所述加料泵向步骤b得到的所述混合溶液中加入酰化剂，得到酰化液。
在一些实施例中，所述釜体包括：
第一壳体，所述出液口设置在所述第一壳体上，所述第一壳体的上端敞开；
上盖，所述上盖密封盖装在所述第一壳体上，所述第一壳体和所述上盖限定出所述第一腔室，所述第一加料口、所述第二加料口和所述搅拌口均设置在所述上盖上；和
夹套，所述夹套套设在所述第一壳体上，所述夹套具有供换热介质进入的换热介质第一进口和供换热介质流出的换热介质第一出口。
在一些实施例中，所述配置系统还包括加热装置，所述加热装置具有换热介质第二进口和换热介质第二出口，所述换热介质第一出口与所述换热介质第二进口相连，所述换热介质第一进口 与所述换热介质第二出口相连；
温度传感器，所述釜体具有与所述第一腔室连通的测温口，所述温度传感器的一部分密封插装在所述测温口内，所述温度传感器的检测端设置在所述第一腔室内；和
控制器，所述控制器与所述加热装置和所述温度传感器中的每一者相连，以便所述控制器根据所述温度传感器检测到的温度控制所述加热装置。
在一些实施例中，所述配置系统还包括抽滤装置，所述抽滤装置包括：
过滤装置，所述过滤装置包括第二壳体和过滤膜，所述第二壳体限定出第二腔室，所述过滤膜设置在所述第二腔室内，所述过滤膜将所述第二腔室分隔成第一部分和第二部分，所述第二壳体上设有与所述第二腔室连通的过滤装置进口和过滤装置出口，所述过滤装置进口对应所述第一部分设置，所述过滤装置出口对应所述第二部分设置，所述过滤装置进口与所述出液口相连；
储液罐，所述储液罐包括第三壳体，所述第三壳体限定出第三腔室，所述第三壳体上设有储液罐进口和储液罐出口，所述储液罐进口与所述过滤装置出口连通；和
抽真空泵，所述抽真空泵具有第一抽真空口，所述第三壳体上设有与所述第三腔室连通的第二抽真空口，所述第一抽真空口与所述第二抽真空口连通。
在一些实施例中，所述第一壳体、所述第二壳体和所述第三壳体为一体式结构，其中所述出液口设置在所述第一壳体的底部，所述过滤装置进口设置在所述第二壳体的顶部，所述过滤装置设置在所述釜体的底部，所述过滤装置出口设置在所述第二壳体的底部，所述储液罐进口设置在所述第三壳体的顶部，所述储液罐设置在所述过滤装置的底部。
在一些实施例中，所述溶剂为硝基苯，步骤b为：
首先将所述硝基苯通过所述第一加料口加入所述第一腔室内；
然后将步骤a称量的所述Lewis催化剂加入所述硝基苯中；
之后利用换热介质将所述硝基苯和所述Lewis催化剂加热至50℃-60℃，同时利用所述搅拌器在200rpm-400rpm下搅拌。
在一些实施例中，利用配置系统配制酰化液还包括步骤d，在惰性气体氛围下，利用所述抽滤装置对步骤c得到酰化液进行抽滤处理，以便得到过滤后酰化液。
在一些实施例中，步骤c中，所述酰化剂的加入速度为每秒3滴-10滴，所述搅拌器的搅拌速度为200rpm-400rpm。
在一些实施例中，所述酰化剂与所述Lewis催化剂的摩尔质量比为(1.1-1.5):(1.3-1.7)；和/或
所述Lewis催化剂与所述溶剂的摩尔质量比为(1.3-1.7)：5。
在一些实施例中提出了一种连续合成酰基萘的方法，利用上述任一所述的酰化液进行合成包括如下步骤：
S1：将含2-甲基萘的原料液和上述任一实施例中的所述的酰化液混合，其中所述2-甲基萘：所述酰化剂：所述Lewis催化剂为1：1.3：1.5形成酰化反应液；
S2：所述酰化反应液依次进入微通道反应器和多釜串联的釜式反应器中进行酰化反应并同步进行水解以及后续通过精馏及结晶得到2-甲基-6-丙酰基萘。
在一些实施例中，所述微通道反应器设有若干个并相互并联。
在一些实施例中，步骤S1中，将所述原料液和所述酰化液通过注射器注入三通型混合器里进行混合。
在一些实施例中，所述三通型混合器为T型混合器或Y型混合器。
在一些实施例中，所述三通型混合器和所述微通道反应器放置于第一恒温槽中，温度控制在-5-0℃，所述釜式反应器放置于第二恒温槽中，温度控制在30-50℃。
在一些实施例中，所述釜式反应器设有2-4个且相互串联。
在一些实施例中，所述酰化反应液在所述釜式反应器中的总停留时间为50-80min。
在一些实施例中，所述微通道反应器和所述釜式反应器的材质均为耐强酸腐蚀材料。
在一些实施例中，所述微通道反应器的内径为0.5-3.175毫米。
在一些实施例中，所述酰化液中的酰化剂为丙酰氯、乙酰氯、乙酸酐、丙酸酐中的任一种。
在一些实施例中，所述水解方法为在水解段中，将酰化反应后的酰化反应液通入到水解段之前将水通入水解段的水相管路，在水流出水解段出口时，再将所述酰化反应液通入到所述水解段 的油相管路，并与水在低温冷浴器中迅速混合后进入管式反应器进行水解反应，水解后的混合液从所述管式反应器的出口流出，并采用分液器收集混合液进行水相和油相的分离；在此阶段，水始终保持流通状态直至混合液完全排出，实现水解反应与酰化反应同步进行。
在一些实施例中，通过开启注水泵向所述水解段的所述水相管路通水，控制所述水相管路内水的流速为3-15mL/min。
在一些实施例中，所述酰化反应液经过管路上的单向阀，直接进入水解段的油相管路。
在一些实施例中，所述单向阀为不锈钢卡套单向阀，内部流道的材质为聚四氟乙烯。
在一些实施例中，所述低温冷浴器的温度为0-20℃。
在一些实施例中，所述管式反应器的反应温度为30-40℃。
在一些实施例中，所述分液器包括容器、搅拌装置、进料口、出料口和排气口，所述搅拌装置伸入所述容器内并与外置的减速器连接进行驱动，所述出料口连通于所述容器的底端，所述进料口和所述排气口共同连通于所述容器上端；所述进料口通过管路接通水解段的所述管式反应器的出口，所述排气口通过管线接入一个抽气泵；所述进料口、所述出料口和所述排气口处均设有控制阀。
在一些实施例中，水解段的所述管式反应器中进行水解反应时，同时所述抽气泵将产生的废气抽出。
在一些实施例中，连续合成酰基萘的方法还包括步骤S3：酰基化废水处理；具体包括以下步骤
(1)调节所述分液器分离出的所述水相的pH至碱性，获得含有氢氧化铝沉淀的悬浮液，并将所述悬浮液过滤得到氢氧化铝滤饼和滤液；
(2)将所述滤液调节pH至酸性，随后转入萃取剂中搅拌、静置，获得分层的萃余相和萃取相；
(3)将所述萃余相经生化处理后达标排放；
(4)将所述萃取相精馏分离后获得硝基苯和所述萃取剂，将所述硝基苯作为有机溶剂回用，将所述萃取剂重新用于所述滤液萃取；
(5)将所述滤饼加入浓盐酸加热溶解，添加助剂后聚合得到聚合氯化铝。
在一些实施例中，步骤(1)中，调节所述分液器分离出的所述水相的pH至碱性的方法包括：采用碱原料调节所述水相的pH至8-10，所述碱原料为氢氧化钠、氢氧化钾或液氨中的一种或两种以上。
在一些实施例中，步骤(2)中，将所述滤液调节pH至酸性，包括：采用酸调节滤液的pH至2-3，酸为盐酸、硝酸或硫酸中的一种或两种以上。
在一些实施例中，步骤(2)中，所述萃取剂为非极性有机溶剂，且所述萃取剂与滤液的体积比为0.5-5:1。
在一些实施例中，所述萃取剂为正庚烷、正辛烷、正己烷、苯、甲苯、二甲苯、四氯化碳中的至少一种。
在一些实施例中，步骤(3)中，生化处理采用活性污泥法；步骤(5)中，加热溶解的温度为40-50℃，滤饼与浓盐酸的重量比为0.5-2.5:1。
在一些实施例中，步骤(5)中，所述助剂的添加量为滤饼干重的2-10wt％，所述助剂为铝酸钙或/和铝酸镁。
在一些实施例中，精馏包括以下步骤：
(1)将所述油相从第一精馏塔中部送料至第一精馏塔中进行精馏，轻馏分从塔顶蒸出，经冷凝后回收，塔釜液从塔底部流出；
(2)将步骤(1)中的塔釜液趁热泵送至第二精馏塔中进一步精馏，第二精馏塔的轻馏分从塔顶蒸出，经冷凝后收集；产品从第二精馏塔测线采出；塔釜重组分进入塔底，并由塔底再沸器加热后，将塔釜重组分趁热泵送至塔釜料收集罐中。
在一些实施例中，所述第一精馏塔的压力为0.05KPa-210KPa；回流比为(1-2)：1。
在一些实施例中，其中，所述第一精馏塔的冷凝温度为10-220℃。
在一些实施例中，所述第二精馏塔的压力为0.05KPa-210KPa；回流比为(5-210)：1。
在一些实施例中，所述第二精馏塔的冷凝温度为50-290℃。
在一些实施例中，所述2-甲基萘为从洗油中提取，纯度为99.0-99.9％；其中提取方法为：
I洗油通过精馏分离得到甲基萘富集馏分；
II将所述甲基萘富集馏分通入共沸精馏塔进行共沸精馏，得到共沸馏出物；
III将所述共沸馏出物通入分离器得到2-甲基萘粗品；
IV将所述2-甲基萘粗品通入多个并联设置的间歇式熔融结晶器进行2-甲基萘的结晶提纯。
本申请实施例中配制得到的酰化液，溶液均相性好，稳定性高能够用于微通道反应器，并且利用酰化液进行酰化反应制备2-甲基-6-丙酰基萘时，酰化反应液立刻接入水解段，连续稳定的同步水解，提高了物料平衡率且没有管路堵塞的风险，水解出来的液体可以直接进入分液器中，同步分液，油相可以方便收集，产生的废气可以从分液器的上端排气口抽出。此外，同步水解出的油相pH值为6-7，达到提纯的要求。
附图说明
本申请上述的和/或附加的方面和优点从下面结合附图对实施例的描述中将变得明显和容易理解，其中：
图1是本申请实施例提供的配置系统的结构示意图；
图2是图1中釜体、过滤装置和储液罐处的内部结构示意图；
图3是图2中pH计的结构示意图的结构示意图；
图4是本申请实施例提供的焊接系统的结构示意图；
图5是本申请实施例提供的打孔系统的结构示意图；
图6为根据本申请实施例提供的精馏工艺流程图；
图7为根据本申请实施例提供的可资源回收的分液器分离出的水相的处理工艺流程图；
图8为常规活性污泥处理系统图；
图9为本申请实施例8中利用图7的处理工艺时分液器分离出的水相中COD的变化图；
图10为本申请实施例9中利用图7的处理工艺时分液器分离出的水相中COD的变化图；
图11为本申请实施例10中利用图7的处理工艺时分液器分离出的水相中COD的变化图；
图12为本申请实施例中从洗油中提取高纯度2-甲基萘的工艺流程图；
图13为图12中的共沸精馏塔的精馏柱的结构示意图；
图14为图12中的间歇式熔融结晶器的结构示意图；
图15为本申请图12中的结晶板的结构示意图；
其中100、配置系统；
1、釜体；101、第一加料口；102、第二加料口；103、搅拌口；104、出液口；1041、第一控制阀；105、测温口；106、pH计口；107、第一壳体；108、上盖；109、夹套；1091、换热介质第一进口；1092、换热介质第一出口；110、第一腔室；
2、加料口盖；
3、搅拌器；301、搅拌轴；302、搅拌叶片；
4、pH计；
5、温度传感器；
6、加料泵；600、第一软管；
7、过滤装置；701、第二壳体；7011、过滤装置进口；7012、过滤装置出口；7013、第一部分；7014、第二部分；702、过滤膜；
8、储液罐；801、第三壳体；8011、储液罐进口；8012、储液罐出口；8013、第二抽真空口；
9、抽真空泵；900、第二软管；901、第一抽真空口；
10、玻璃罩；1001、配合面；
11、三通型混合器；
12、微通道反应器；
13、釜式反应器；
14、原料液；
15、酰化液；
16、分液器；1601、排气口；1602、进料口；1603、控制阀；1604、容器；1605、出料口；
1701、调节池；1702、沉淀池；1703、过滤单元；1704、萃取单元；1705、精馏塔；1706、生化处理单元；17061、曝气池；17602、二沉池；1707、溶解池；1708、反应釜；
1801、洗油；1802、萘馏分轻质油；1803、甲基萘富集馏分；1804、重质馏分油；1805、共沸馏出物；1806、残油；1807、共沸剂混合物；1808、甲基萘粗品；1809、水；1810、共沸剂；1811、加热冷却介质；1812、残余母液；1813、甲基萘产品；1814、控温介质出口；1815、控温介质入口；1816、物料进口；1817、物料出口；1818、结晶板控温介质入口；1819、结晶板控温介质出口；1820、凸起；1821、凹槽；1822、螺旋形凹槽；
V1、洗油储罐；V2、常压精馏塔；V3、共沸精馏塔；V4、超声静态混合器；V5、分离器；V6、精馏塔；V7、数字式可控温油浴；V8、间歇式熔融结晶器。
具体实施方式
为了能够更清楚地理解本申请的上述目的、特征和优点，下面结合附图和具体实施方式对本申请进行进一步的详细描述。需要说明的是，在不冲突的情况下，本申请的实施例及实施例中的特征可以相互组合。
在下面的描述中阐述了很多具体细节以便于充分理解本申请，但是，本申请还可以采用其他不同于在此描述的其他方式来实施，因此，本申请的保护范围并不受下面公开的具体实施例的限制。
如图1至图3所示，本申请实施例的用于配制酰化液的配置系统100包括釜体1、加料口盖2、加料泵6和搅拌器3。釜体1限定出第一腔室110，釜体1具有与第一腔室110连通的第一加料口101、第二加料口102、搅拌口103和出液口104。
可选地，第一加料口101、第二加料口102和搅拌口103中的每一者设置在出液口104的上方。
由此，用于配制酰化液的溶剂和Lewis催化剂可以通过第一加料口101加入第一腔室110内，酰化剂可以通过第二加料口102加入第一腔室110内。
加料口盖2可拆卸地密封盖装在第一加料口101上。在需要向釜体1内加料(溶剂和Lewis催化剂)的情况下，从第一加料口101上取下加料口盖2，加料完成后，及时将加料口盖2密封盖装在第一加料口101上，从而减少溶剂和Lewis催化剂与外界环境的接触时间，同时减少第一腔室110通过第一加料口101暴露在外界环境中的时间。
加料泵6与第二加料口102密封相连，以便向第二加料口102内添加酰化剂。由此，在利用加料泵6向第一腔室110内加料过程中以及加料完成后，可以有效避免酰化剂与外界环境接触，同时避免第一腔室110通过第二加料口102暴露在外界环境中。
搅拌器3的一部分密封插装在搅拌口103内，搅拌器3包括搅拌轴301和搅拌叶片302，搅拌叶片302设置在搅拌轴301上，搅拌轴301的至少一部分和搅拌叶片302中的每一者设置在第一腔室110内。
由此，利用搅拌器3可以对第一腔室110内的物料进行充分搅拌混合，使Lewis催化剂充分溶解，提高酰化液的配制效率；并且搅拌器3的一部分密封插装在搅拌口103内，使得在搅拌器3工作过程中，搅拌口103处始终保持密封，避免第一腔室110通过搅拌口103暴露在外界环境中。
利用本申请实施例的用于配制酰化液的配置系统100实施的用于配制酰化液的方法包括以下步骤：
a、在惰性气体保护下称量Lewis催化剂；
b、将溶剂和步骤a称量的Lewis催化剂通过第一加料口101加入第一腔室110内，利用搅拌器3搅拌，以便步骤a称量的Lewis催化剂溶解而得到混合溶液；
c、利用加料泵6向步骤b得到的混合溶液中加入酰化剂，并利用搅拌器3搅拌，以便得到酰化液。
本申请实施例的用于配制酰化液的方法，称量Lewis催化剂时，在惰性气体保护下进行称量，有效保证了Lewis催化剂的活性，不受天气环境的影响，避免了夏天湿度较大时酸雾的形成；本 申请实施例的用于配制酰化液的方法中，在惰性环境称量Lewis催化剂后加入到溶剂中，并盖装上加料口盖2，通过液封的方式减少了Lewis催化剂与空气的接触时间，使得在配制酰化液的过程中，可以有效减少甚至Lewis催化剂和酰化剂与水分等接触，进而使得配制出酰化液的稳定性和均一性均较好；本申请实施例的用于配制酰化液的方法中，由于搅拌器3的搅拌作用，从而可以使Lewis催化剂充分快速溶解在溶剂中，有利于缩短酰化液的配制时间，提高酰化液的配制效率；本申请实施例的用于配制酰化液的方法配制得到的酰化液，溶液均相性好，稳定性高能够用于微通道反应器，并且提高了反应效率。
可选地，步骤a中，Lewis催化剂选自AlCl 3、BF 3、ZnCl 2或FeCl 3中的至少一种。本申请实施例中对Lewis催化剂没有特别限制，能够用于合成2-甲基-6-丙酰基萘的Lewis催化剂均可以采用本申请实施例的用于配制酰化液的方法进行酰化液的配制。酰化剂可以为酰化剂选自乙酰化剂、丙酰化剂或丁酰化剂中的至少一种。
可选地，第二加料口102为磨口，第二加料口102与宝塔接头密封相连，蠕动泵的出料管与宝塔接头密封相连。
可选地，加料泵6为蠕动泵。
由此，可以通过调节蠕动泵的流速，来控制向釜体1的第一腔室110内加入的酰化剂的流速，从而方便控制加入第一腔室110内的酰化剂的量，有利于进一步提高配制出的酰化液的稳定性和均一性。
可选地，加料泵6与第二加料口102通过第一软管600相连。第一软管600可以为聚四氟乙烯管。
在一些实施例中，搅拌轴301和搅拌叶片302的材质为金属，搅拌器3包括轴防腐蚀层和叶片防腐蚀层，轴防腐蚀层包覆在搅拌轴301的一部分上，叶片防腐蚀层包覆在搅拌叶片302上。
在一些实施例中，搅拌轴301和搅拌叶片302的材质为不锈钢，轴防腐蚀层和叶片防腐蚀层的材质为聚四氟乙烯。
搅拌轴301和搅拌叶片302采用金属材质，可以有效保证搅拌轴301和搅拌叶片302的具有足够的结构强度。在搅拌轴301上包覆轴防腐蚀层、在搅拌叶片302上包覆叶片防腐蚀层，则可以避免搅拌轴301和搅拌叶片302的金属材质部分与溶剂、Lewis催化剂和酰化剂接触，不仅可以避免搅拌轴301和搅拌叶片302腐蚀，而且可以避免腐蚀产物进入酰化液而影响配制的酰化液的质量。
可选地，上述步骤c中，酰化剂的加入速度为每秒3滴-10滴，搅拌器3的搅拌速度为200rpm-400rpm。
在一些实施例中，如图2所示，釜体1包括第一壳体107、上盖108和夹套109，出液口104设置在第一壳体107上，第一壳体107的上端敞开。上盖108密封盖装在第一壳体107上，第一壳体107和上盖108限定出第一腔室110，第一加料口101、第二加料口102和搅拌口103中的每一者设置在上盖108上。夹套109套设在第一壳体107上，夹套109具有供换热介质进入的换热介质第一进口1091和供换热介质流出的换热介质第一出口1092。
由此，在进行釜体1的加工时，第一壳体107和上盖108可以分别单独加工，方便釜体1的加工。在进行Lewis催化剂的溶解操作时，换热介质可以通过换热介质第一进口1091流入夹套109内，并通过第一壳体107与第一腔室110内的物料换热，之后换热介质通过换热介质第一出口1092流出夹套109。从而实现对第一腔室110内物料的加热，使得第一腔室110内的温度保持在适合Lewis催化剂溶剂的预设温度，而使Lewis催化剂快速溶解在溶剂中，从而大大缩短酰化液的配制时间，提高酰化液的配制效率。
换热介质可以采用水、油等液体。
在一些实施例中，溶剂为硝基苯，上述步骤b为：
首先将硝基苯通过第一加料口101加入第一腔室110内；
然后将步骤a称量的Lewis催化剂加入硝基苯中；
之后利用夹套109内的换热介质将硝基苯和Lewis催化剂加热至50℃-60℃，同时利用搅拌器3在200rpm-400rpm下搅拌，以便步骤a称量的Lewis催化剂溶解而得到混合溶液。
由此，使得第一腔室110内的温度保持在适合Lewis催化剂溶解的预设温度，搅拌器3的搅拌速度也保持在适合Lewis催化剂溶解的搅拌速度，而使Lewis催化剂快速溶解在硝基苯中，从 而大大缩短酰化液的配制时间，提高酰化液的配制效率。
在一些实施例中，酰化液配制系统进一步包括加热装置、温度传感器5和控制器。加热装置具有换热介质第二进口和换热介质第二出口，换热介质第一出口1092与换热介质第二进口相连，换热介质第一进口1091与换热介质第二出口相连。
釜体1具有与第一腔室110连通的测温口105，温度传感器5的一部分密封插装在测温口105内，温度传感器5的检测端设置在第一腔室110内。控制器与加热装置和温度传感器5中的每一者相连，以便控制器根据温度传感器5检测到的温度控制加热装置。
由此，被加热装置加热后的换热介质通过换热介质第二出口流至换热介质第一进口1091，换热介质通过换热介质第一进口1091流至夹套109内，利用夹套109内的换热介质对第一腔室110内的物料进行加热；之后夹套109内的换热介质通过换热介质第一出口1092流出夹套109，并通过换热介质第二进口流回加热装置而被加热装置加热，实现换热介质在加热装置和釜体1之间的循环流动。
利用温度传感器5可以对第一腔室110内的物料温度进行实时检测，温度传感器5检测的第一腔室110内物料的温度传递给控制器，以便利用控制器控制加热装置。在一些实施例中，在温度传感器5检测到的温度高于预设温度的情况下，控制器控制加热装置停止加热，以免第一腔室110内的物料高于预设温度；在温度传感器5检测到的温度低于预设温度的情况下，控制器控制加热装置开始加热，以免第一腔室110内的物料低于预设温度。使得Lewis催化剂溶解过程中，第一腔室110内的温度保持在适合Lewis催化剂溶解的预设温度，有利于提高Lewis催化剂的溶解速度，提高酰化液的配制效率。
可选地，控制器的型号为DSC350。
上述步骤c中，上述步骤b得到的混合溶液的温度不高于60℃。
本申请实施例的用于配制酰化液的方法中，在步骤b中Lewis催化剂溶解在溶剂中后，可以无需对混合溶液进行降温处理，直接加入酰化剂配制酰化液，不仅降低了能耗而且缩短了酰化液的配制时间，提高了配制效率。
当然，在步骤b中Lewis催化剂溶解在溶剂中后，也可以在利用加料泵6向第一腔室110内添加酰化剂的同时，通过换热介质第一进口1091向夹套109内通入换热介质(例如，冷却水或冷却剂)，实现对混合溶液的冷却。
如图1和图2所示，用于配制酰化液的配置系统100进一步包括pH计4，釜体1具有与第一腔室110连通的pH计口106，pH计4的一部分密封插装在pH计口106内，pH计4的检测端设置在第一腔室110内。
由此，利用pH计4可以对第一腔室110内物料的酸碱度进行实时检测，有利于提高酰化液的配制效率，提高配制出的酰化液的质量。
在一些实施例中，pH计口106为磨口。由此，方便实现pH计口106的密封。
例如，如图3所示，pH计4的外部设有玻璃罩10，玻璃罩10具有锥形的配合面1001，配合面1001与pH计口106密封配合。
可选地，pH计4的型号为SIN-PH6.3-5022-AL/Y。
在一些实施例中，进一步包括抽滤装置，抽滤装置包括过滤装置7、储液罐8和抽真空泵9。
过滤装置7包括第二壳体701和过滤膜702，第二壳体701限定出第二腔室，过滤膜702设置在第二腔室内，过滤膜702将第二腔室分隔成第一部分7013和第二部分7014，第二壳体701上设有与第二腔室连通的过滤装置进口7011和过滤装置出口7012，第一部分7013邻近过滤装置进口7011设置，第二部分7014邻近过滤装置出口7012设置，过滤装置进口7011与出液口104相连。
储液罐8包括第三壳体801，第三壳体801限定出第三腔室，第三壳体801上设有储液罐进口8011和储液罐出口8012，储液罐进口8011与过滤装置出口7012连通。抽真空泵9具有第一抽真空口901，第三壳体801上设有与第三腔室连通的第二抽真空口8013，第一抽真空口901与第二抽真空口8013连通。
可选地，用于配制酰化液的方法还包括步骤d，将上述步骤c得到酰化液在惰性气体氛围下进行抽滤处理，去除固体颗粒。本申请用于配制酰化液的方法中，将配制后的酰化液采用惰性氛围下抽滤处理，将溶液中未溶解的固体颗粒全部去除，进一步提高了溶液的均相性。并且抽真空 泵9使过滤装置7内形成负压，使得利用釜体1配制好的酰化液在该负压作用下，快速的经过滤装置7的过滤膜702，并流入储液罐8内储存，从而可以进一步提高酰化液的整体配制效率。
可选地，如图1所示，第一抽真空口901和第二抽真空口8013通过第二软管900相连。
在一些实施例中，在进行抽滤操作时，可以通过第一加料口101与惰性气体气源相连，利用第一加料口101向酰化液配制系统中充入惰性气体，从而使得配制后的酰化液在惰性氛围下进行抽滤处理。或者，在向第一加料口101加入溶剂之前，便通过第一加料口101向第一壳体107、第二壳体701和第三壳体801内通入惰性气体，以便使酰化液的整个配制过程均在惰性氛围下进行。
在一些实施例中，第一壳体107、第二壳体701和第三壳体801为一体式结构。其中，出液口104设置在第一壳体107的底部，过滤装置进口7011设置在第二壳体701的顶部，过滤装置7设置在釜体1的下部。过滤装置出口7012设置在第二壳体701的底部，储液罐进口8011设置在第三壳体801的顶部，储液罐8设置在过滤装置7的下部。
由此，釜体1内配制好的酰化液可以直接靠自身重力经出液口104流出釜体1，并经过滤装置进口7011进入过滤装置7内，不需要在出液口104和过滤装置进口7011之间，以及过滤装置出口7012和储液罐进口8011之间额外设置泵送酰化液的液泵。有利于简化用于配制酰化液的配置系统100的整体结构，降低用于配制酰化液的配置系统100的制造和运行成本。另外，在进行用于配制酰化液的配置系统100的组装时，可以省去出液口104与过滤装置进口7011以及过滤装置出口7012和储液罐进口8011的连接，方便用于配制酰化液的配置系统100的组装。
可选地，第一壳体107、夹套109、第二壳体701和第三壳体801的材质为高硼硅玻璃。由此，方便观察釜体1、过滤装置7和储液罐8内的情况。
在一些实施例中，用于配制酰化液的配置系统100进一步包括第一控制阀1041，第一控制阀1041设置在出液口104上，以便控制出液口104的通断。
由此，在开启第一控制阀1041的情况下，配制好的酰化液通过出液口104流出釜体1，在关闭第一控制阀1041的情况下，釜体1密封，配制好的酰化液保存在釜体1内。
在一些实施例中，用于配制酰化液的配置系统100进一步包括第二控制阀，第二控制阀设置在储液罐出口8012上，以便控制储液罐出口8012的通断。
由此，响应于开启第二控制阀，过滤后酰化液通过储液罐出口8012流出储液罐8，在关闭第二控制阀的情况下，储液罐8密封，过滤后酰化液保存在储液罐8内。
在一些实施例中，酰化剂、Lewis催化剂与溶剂的摩尔质量比为(1.1-1.5):(1.3-1.7):5；和/或Lewis催化剂与溶剂的摩尔质量比为(1.3-1.7)：5。换言之，酰化剂、Lewis催化剂与溶剂的摩尔质量满足：酰化剂、Lewis催化剂与溶剂的摩尔质量比为(1.1-1.5):(1.3-1.7):5，且Lewis催化剂与溶剂的摩尔质量比为(1.3-1.7)：5；或者，酰化剂、Lewis催化剂与溶剂的摩尔质量比为(1.1-1.5):(1.3-1.7):5；或者，Lewis催化剂与溶剂的摩尔质量比为(1.3-1.7)：5。
本申请实施例的用于配制酰化液的方法中，优化了各物质的配比，能够使原料得到充分利用，降低了生产成本。
如图12所示，在一些实施例中，将配置好的酰化液进行连续合成酰基萘，其包括，
原料液14的配制：以酰化液为5L为例，与5L的酰化液在常温下装有搅拌装置的反应器内加入700g的硝基苯和284g的2-甲基萘，制成原料液14。其中2-甲基萘为从洗油中提取，纯度为99.0-99.9％；其中2-甲基萘从洗油中提取的方法为：
I洗油通过精馏分离得到甲基萘富集馏分；
II将甲基萘富集馏分通入共沸精馏塔进行共沸精馏，得到共沸馏出物；
III将共沸馏出物通入分离器得到2-甲基萘粗品；
IV将2-甲基萘粗品通入多个并联设置的间歇式熔融结晶器进行2-甲基萘的结晶提纯。
在一些实施例中，步骤I洗中，洗油通过常压精馏塔进行精馏分离，得到萘馏分轻质油、甲基萘富集馏分和重质馏分油，其中重质馏分油的一部分循环回常压精馏塔，另一部分回收利用。萘馏分轻质油从常压精馏塔的塔顶得到，甲基萘富集馏分从常压精馏塔的侧线得到，重质馏分油从常压精馏塔的塔底得到，塔顶温度为210-225℃，侧线温度为235-260℃，塔底温度为290-310℃。
步骤II中，将甲基萘富集馏分和共沸剂按照比例混合，混合物加热到一定温度后进入共沸精 馏塔进行共沸精馏。其中共沸剂为单一化合物共沸剂或混合物型共沸剂，其中单一化合物共沸剂为乙二醇、二甘醇、乙醇胺、二乙二醇、N-甲基甲酰胺中的任一种，混合物型共沸剂为庚烷与乙醇胺的混合物或庚烷和乙二醇的混合物；共沸精馏塔的塔顶温度为150-175℃，塔顶压力为1-4KPa，回流比为5-15；塔底温度为220-245℃，塔底压力为10-15KPa。
基于采用单一化合物共沸剂，甲基萘富集馏分与共沸剂混合的质量比为1-3:1；响应于采用混合物型共沸剂，混合物型共沸剂中的庚烷与乙醇胺或庚烷与乙二醇的质量比为0.2-1:1，甲基萘富集馏分与共沸剂混合的质量比为1:0.8-2。步骤III中，将共沸馏出物和水通入超声静态混合器中进行超声混合，然后通入分离器进行油水静置分离，分离器的操作温度为50-80℃，静置时间为0.2-1小时，从分离器分离出水相和油相，其中，水相为共沸剂和水，水相进入精馏塔进行共沸剂和水的蒸馏分离，分离后的共沸剂返回至步骤II，和甲基萘富集馏分混合，循环利用；分离后的水返回和共沸馏出物混合，循环利用；油相为2-甲基萘粗品，油相进入间歇式熔融结晶器进行2-甲基萘的结晶提纯。
步骤IV中，间歇式熔融结晶器设有2个，晶体生长过程初温为35-40℃，降温速率为3-8℃/h，终温为8-12℃，恒温时间为0.5-1h；发汗过程升温速率为2-6℃/h，终温为30-32℃；熔融过程升温速率为2-8℃/h，终温为45-50℃，终温恒温时间0.5-1h。
在一些实施例中，2-甲基萘为从洗油中提取的方法为：洗油1801加入洗油储罐V1中，经过搅拌混合后，用泵将洗油1801送入常压精馏塔V2，洗油1801通过精馏分离，塔顶得到萘馏分轻质油1802，可以作为提取萘的原料，也可以循环回常压精馏塔V2；侧线得到甲基萘富集馏分1803，塔底采出物为重质油馏分1804，重质油馏分1804的一部分循环回常压精馏塔V2，另一部分可送往煤焦油加工厂去进一步分离苊、氧芴、工业芴等产品；常压蒸馏塔V2塔顶温度为210-225℃，侧线温度为235-260℃，塔底温度为290-310℃。将上述得到的甲基萘富集馏分1803和共沸剂1810按照一定比例混合，混合物加热到一定温度后进入共沸精馏塔V3进行共沸精馏。共沸精馏塔V3的塔顶得到的共沸馏出物1805进入分离器进行共沸剂1810的回收利用，塔底得到残油1806，将残油1806外甩，作为提取其它精细化学品的原料。
由于共沸剂1810和水1809互溶性强，本申请采用水萃取的方法回收共沸剂1810，步骤(2)得到的共沸馏出物1805和水1809混合后，首先在超声静态混合器V4中进行超声混合，然后进入分离器V5进行油水静置分离。从分离器分离出水相和油相，其中，水相为共沸剂混合物1807，包括共沸剂1810和水1809，油相为2-甲基萘粗品。
从分离器V5出来的水相进入精馏塔V6进行共沸剂1810和水1809的蒸馏分离，分离后从塔底得到的共沸剂1810返回步骤(2)去和甲基萘富集馏分1803混合，循环利用；分离后从塔顶得到的水返回和共沸馏出物1805混合，循环利用。从分离器V5出来的2-甲基萘粗品去熔融结晶提纯。此外，采用两个间歇式熔融结晶器V8并联的方式，交替进行2-甲基萘的结晶提纯，达到连续化结晶提纯的目标。在间歇式熔融结晶器V8中，甲基萘粗品1808在间歇式熔融结晶器V8中经过晶体生长、发汗、熔融等工艺过程，最终实现结晶提纯的目的。甲基萘粗品1808经过熔融结晶提纯后得到高纯度的2-甲基萘产品1813，在晶体生长和发汗过程排出的液体是残余母液1812，残余母液1812可作为提取1-甲基萘的原料。经过结晶提纯后得到的2-甲基萘产品1813纯度为99.0-99.9％，产品收率70％-90％。
其中上述2-甲基萘为从洗油中提取的方法中，常压精馏塔V2为填料塔，填料高度相当于理论塔板数为30-50层的不锈钢填料塔，塔底的重质油馏分1804一部分循环和洗油混合后进入常压精馏塔V2，一部分外甩，外甩和循环的比例为2-5:1，得到的甲基萘富集馏分1803中2-甲基萘的含量为50-70％。共沸精馏塔V3采用同心管精密分馏柱如图13所示，分馏柱由两根经精巧设计和精密校准的同心管融合而成，内管表面设有螺旋形凹槽1822，能使垂直上升的蒸汽与同心环形间隙中的液体薄膜之间高效传质传热，材质为玻璃或不锈钢，分馏柱理论塔板数为80-120。熔融结晶器为充满式正方体型静态结晶器如图14所示，外形为正方体，设有控温介质出口品1814、控温介质入口1815、物料进口1816和物料出口1817，熔融结晶器内设有多组相互平行的结晶板，如图15所示，每组结晶板上布置有盘旋状的凸起1820，结晶时晶层在凸起1820上生长，不易脱落；发汗过程中，晶层从结晶板的生长面脱落，进入凸起1820间的凹槽1821内，并保持与换热面接触，防止晶层发汗过程中易脱落；结晶板的内部流通加热、冷凝介质(即控温介质)，控温介质通过结晶板控温介质入口1818流入，通过结晶板控温介质出口1819流出，循环至下一组结晶 板，最终通过熔融结晶器的控温介质出口品1814流出。结晶板的材质为不锈钢、有机玻璃或其他合金。每个间歇式熔融结晶器V8均连接数字式可控温油浴V7，数字式可控温油浴V7与间歇式熔融结晶器V8之间传输加热冷却介质1811。
因此本申请的实施例中一改相关技术中利用苯系化合物为原料的传统工艺方法和以煤基的萘系化合物为原料的改进工艺方法中以萘和甲基萘为原料进行2,6-二甲基萘合成的方法，利用煤焦油洗油中提取高纯度2-甲基萘，实现提高2-甲基萘分离精制工艺中生产效率、产品纯度和收率；降低了污染和设备腐蚀概率，从而实现从原料2-甲基萘作为原料液与酰化液进行酰化、水解等合成得到2-甲基-6-丙酰基萘，达到了采用非石油基的路线实现2-甲基-6-丙酰基萘的工业生产的目的。
其中将配置好的酰化液和原料液14进行酰化反应如图4所示：用计量泵分别吸取原料液14和酰化液15，其中酰化液15的(计量泵)流速为158-162g/min，原料液14的(计量泵)流速为82-84g/min，即按照摩尔比为2-甲基萘：丙酰氯：AlCl 3为1：1.3：1.5。将原料液14和酰化液15通过注射器注入三通型混合器11里进行混合，混合后进入微通道反应器12反应，5-10min后酰化反应液从微通道反应器12的出口流出，直接流入多釜串联的釜式反应器13中，在釜式反应器13中搅拌反应，经过2-4个釜式反应器13，酰化反应液在多釜串联的釜式反应器13中的总停留时间为50-80min，从釜式反应器13中流出的酰化反应液进行同步水解、并经过精馏后得到2-甲基-6-丙酰基萘。
釜式反应器13采用多釜串联反应器实现连续反应，通过调节阀来控制流速，进而控制物料在釜式反应器13的停留时间。微通道反应器12设有若干个并相互并联，酰化液15和原料液14的流速与并联的数量成正比。
其中，三通型混合器11和微通道反应器12放置于第一恒温槽中，温度控制在-5-0℃，在一些实施例中为控制三通型混合器11和微通道反应器12的温度为-5℃；釜式反应器13放置于第二恒温槽中，温度控制在30-50℃，在一些实施例中控制釜式反应器13的温度为35℃。
在一些实施例中，所有并联支路的微通道反应器12的原料液14采用同一个原料液泵输送，酰化液15采用同一个酰化液泵输送。
在一些实施例中，并联的微通道反应器12与多釜串联的釜式反应器13采用直接相连的方式。
在一些实施例中，三通型混合器11为内径为0.5mm-3mm的T型混合器或Y型混合器。
在一些实施例中，微通道反应器12和釜式反应器13的材质均为耐强酸腐蚀材料，且完全密封隔绝空气。
在一些实施例中，微通道反应器12的内径为0.5-3.175毫米。
在一些实施例中，酰化液15中的酰化剂为丙酰氯、乙酰氯、乙酸酐、丙酸酐中的任一种。
在一些实施例中，酰化反应液直接通入水阶段进行水解的方法为：其中在水解段中，提前开启注水泵，在酰化反应液通入到水解段之前就将水(去离子水)通入水解段的水相管路，控制水相管路内水的流速为3-15mL/min，最佳为10mL/min。在水流出水解段出口的情况下，再将酰化反应液经过管路上的单向阀，直接进入水解段的油相管路，与水在低温冷浴器0-20℃，在一些实施例中为0℃，并迅速混合后进入管式反应器进行水解反应，管式反应器的反应温度为30-40℃，在一些实施例中为30℃，内径为2-5mm。同时开启超声震荡对管式反应器内的液体震荡搅拌，并开启末端的分液器16的抽气泵。由于管式反应器内径为2-5mm，液体可以缓慢行进，从管式反应器的进口到出口可行进5-10分钟左右，可以充分发生水解反应。
在管式反应器的出口流出浑浊浅褐色液体(水解后的混合液)的情况下，采用分液器16收集混合液，在此阶段，水始终保持流通状态，直至混合液完全排出，水解反应与反应段的酰化反应同步进行。
分液器16为液液分液器，混合液在分液器16内分层后，油相可从下口放出，测得油相pH值为6-7。分液器16上端通过管线接入抽气泵，可将水解产生的废气(如HCl)抽走，进入后端的碱液灌，收集废气，不会造成空气污染。
在一些实施例中，分液器16的具体结构为：如图5所示，包括容器1604、搅拌装置、进料口1602、出料口1605和排气口1601，搅拌装置伸入容器1604内，搅拌装置通过外置的减速器连接并驱动。搅拌装置在本案的附图中没有完全画出，因为搅拌装置就是常规的结构，不是本案的重点，而且本案也没有用到搅拌装置。出料口1605连通于容器1604底端，进料口1602和排气口1601共同连通于容器1604上端，进料口1602通过管路接通水解段的管式反应器的出口，排气口 1601通过管线接入一个抽气泵，进料口1602、出料口1605和排气口1601处均设有控制阀1603。出料口1605处的控制阀1603先保持关闭状态，直到分层准备排出油相时再打开控制阀1603。分层是为了将水和油相分层，油相会在水的下方，出料口1605会将油相最先排出，油相排完之后不可避免会连带排出一些水，后期可以通过精馏进行提纯。
这里需要强调的是，分液器16的容器1604空间足够大，有效容积为10L，无需担心分液器16内会被填满，管式反应器内的通道内径很细小，液体行进缓慢，哪怕等到最后混合液全部流入到分液器16内也不会被装满。分液器16的耐受温度为-80-200℃，也不会担心分液器16在使用过程中由于温度原因而被损坏。
在一些实施例中，单向阀为不锈钢卡套单向阀，内部流道的材质为聚四氟乙烯，防腐性能更好。普通单向阀内的流道是普通橡胶材质或具有耐腐蚀性的橡胶，但经过试验，其耐腐蚀的效果都不佳，于是更换成聚四氟乙烯材质，大幅度提高防腐性能。
该同步水解的方法是先将水泵打开，再将从釜式反应器13中流出的酰化反应液立刻接入水解段，连续稳定的同步水解，水解效率和物料平衡率都大为提高。
本案中酰化反应与水解反应同步，避免了酰化反应液的临时储存，减少了HCl气体的溢出和污染。
该水解反应是个放热反应，应与水在低温冷浴器(0℃)中迅速混合后再进入管式反应器，可将水解的温度迅速降低。
采用管式反应器进行水解反应时，在酰化反应液与水汇合的管路前段加入单向阀，防止液体从水解段倒流回酰化段，避免了管式反应器通道堵塞的风险。
水解出来的液体可以直接从分液器16的上端进入，同步分液，油相可以方便的从下口放出，测得水解出的油相pH值为6-7，达到提纯的要求。
该水解方法是将从管式反应器中流出的酰化反应液立刻接入水解段，连续稳定的同步水解，提高了物料平衡率，重要的是没有堵塞的风险，水解出来的液体可以直接进入分液器16中，同步分液，油相可以方便的从下口放出进行精馏及结晶处理得到2-甲基-6-丙酰基萘，产生的废气可以从分液器16的上端排气口1601抽出；水相经过酰基化废水的处理工艺进行处理。
在一些实施例中提出了水解反应分离得到的油相进行精馏的方法，如图6所示：
(1)将水解反应分离得到的油相从第一精馏塔中部送料至第一精馏塔中进行精馏，轻馏分从塔顶蒸出，经冷凝后回收，塔釜液从塔底部流出；
在一些实施例中，水解反应分离得到的油相通入第一精馏塔中，其中第一精馏塔中为酰化脱轻塔，所用填料的形状为陶瓷、弹簧、θ环或其组合，脱轻塔的塔顶部的轻馏分主要为少量水经冷凝后回收进入废水高级氧化处理，脱轻塔的侧线采出硝基苯，可回收再次利用；脱轻塔釜底流出的塔釜液为粗产品送入第二精馏塔。
在一些实施例中，第一精馏塔的压力为0.05KPa-10KPa，在一些实施例中为0.1KPa-2KPa；回流比为(1-2)：1。
在一些实施例中，第一精馏塔的冷凝温度为10-20℃，在一些实施例中为10-15℃。
(2)将步骤(1)中的塔釜液趁热泵送至第二精馏塔中进一步精馏，第二精馏塔的轻馏分从塔顶蒸出，经冷凝后收集；产品从第二精馏塔测线采出；塔釜重组分进入塔底，并由塔底再沸器加热后，将塔釜重组分趁热泵送至塔釜料收集罐中。
在一些实施例中，第二精馏塔为脱重塔，所用填料的形状为陶瓷、弹簧、θ环或其组合；脱重塔的塔顶馏出除2-甲基-6-丙酰基萘(2,6-MPN)以外的同分异构体轻组分经冷凝后回收；脱重塔的釜底排出少量较重有色杂质定期进行收集后处理，其中脱重塔的侧线采出酰化段粗产品2,6-MPN馏份冷却后送至重结晶器；2,6-MPN以外的同分异构体中2-甲基-6-酰基萘的酰基基团为甲基、乙基、或异丙基中的一种，即产品2-甲基-6-酰基萘为2-甲基-6-甲酰基萘、2-甲基-6-乙酰基萘或2-甲基-6-异丙基萘。
在一些实施例中，第二精馏塔的压力为0.05KPa-10KPa，在一些实施例中为0.1KPa-2KPa；回流比为(5-10)：1。
在一些实施例中，第二精馏塔的冷凝温度为50-90℃，在一些实施例中为70-90℃。
在一些实施例中，粗产品2,6-MPN进行重结晶的方法为：将甲醇水溶液(其中甲醇：水的质量比为85:15)和粗产品2,6-MPN按8:1的质量比例，在装有搅拌、温度计、回流冷凝管的三口 烧瓶内混合，55℃水浴中搅拌至淡黄色固体全溶，继续搅拌20分钟后，将重结晶液在10℃下结晶出来，结晶6小时。结晶完全后用进行抽滤，固液分离，可得出白色细小粉末2-甲基-6-丙酰基萘，该产物置于干燥器中以除去溶剂，重结晶得到的2,6-MPN浆料在产品过滤机分离滤饼与滤液，母液、冲洗液进入甲醇回收塔，顶部馏出甲醇再次利用；釜底馏出物进入污水系统；湿滤饼采用干燥机进行干燥，干燥机排气进入甲醇回收系统，干燥后得到产品2,6-MPN。
此外，分液器16中分离的水相经过酰基化废水的处理工艺进行处理具体工艺如下所示。
如图7所示，本申请实施例的可资源回收的酰基化废水处理系统，包括：沉淀池1702、过滤单元1703、萃取单元1704、精馏塔1705、生化处理单元1706、溶解池1707和反应釜1708；沉淀池1702用于生成氢氧化铝沉淀；过滤单元1703的入口连通沉淀池1702的出口；萃取单元1704的入口连通过滤单元1703的滤液出口，萃取单元1704设有萃取剂入口；精馏塔1705的入口连通萃取单元1704的萃取相出口，精馏塔1705的萃取剂出口连通萃取剂入口，精馏塔1705的硝基苯出口连通生产用有机溶剂管线；生化处理单元1706的入口连通萃取单元1704的萃余相出口，生化处理单元1706的出口连通废水达标排放管线；溶解池1707的出口连通过滤单元1703的滤饼出口，溶解池1707设有浓盐酸入口；反应釜1708的入口连通溶解池1707的出口。
在一些实施例中，萃取单元1704为离心萃取装置，比如市售的离心萃取机等。
在一些实施例中，生化处理单元1706可以采用活性污泥处理系统。活性污泥系统不属于本申请的重点，采用常规活性污泥处理系统即可。如图8所示，常规的活性污泥处理系统包括曝气池17061、二沉池17602、回流系统、剩余污泥排放系统、供氧系统等，在本申请的一些实施例中，将萃取单元1704的萃余相出口与曝气池17061的进水口连通，即可将萃余相引入活性污泥处理系统，按常规的活性污泥法进行处理，实现达标排放。
需要说明的是，本申请实施例的可资源回收的酰基化废水处理系统，各构件之间的连通方式，可以根据物料的性质，选择通过管线连通或通过转运车运送等实现，在用管线连通的情况下，可根据需要在相应管线上安装阀门和泵，这些均属于常规技术，不属于本申请的重点。
在一些实施例中，为了调节水量、均衡水质并对酰基化废水进行预处理，本申请实施例的可资源回收的酰基化废水处理系统，还包括设在沉淀池1702之前的调节池1701。酰基化废水先进调节池1701调节水质、水量并预处理，再进入沉淀池1702。
在一些实施例中，反应釜可以采用不锈钢反应釜等。沉淀池1702可以采用平流沉淀池或竖流沉淀池等，过滤单元1703可以采用板框压滤机。溶解池可以采用普通的溶解池，精馏塔可以采用填料塔或板式塔等。
本申请实施例的可资源回收的酰基化废水处理系统的工作原理为：
使用时，水相先进入调节池1701调节水量、均衡水质并预处理，再进入沉淀池1702；进入沉淀池1702的酰基化废水经碱源即氢氧化钠、氢氧化钾或液氨中至少一种调节pH至碱性，即pH在8-10，在一些实施例中pH在9左右，酰基化废水中的铝离子沉淀为氢氧化铝，含有氢氧化铝沉淀的悬浮液进入过滤单元1703过滤，得到氢氧化铝滤饼和滤液。其中滤液采用盐酸、硝酸或硫酸中的至少一种调节滤液的pH至2-3后，进入萃取单元1704，在搅拌的条件下，经萃取单元1704中的萃取剂萃取分层，形成萃取相和萃余相，其中萃取剂为非极性有机溶剂，为正庚烷、正辛烷、正己烷、苯、甲苯、二甲苯、四氯化碳中的至少一种，且萃取剂与滤液的体积比为0.5-5:1。
随后萃取相进入精馏塔1705实现硝基苯和萃取剂的高效分离，得到原料硝基苯和萃取剂，从精馏塔1705出来的萃取剂重新进入萃取单元1704，参与对滤液的萃取，而从精馏塔1705出来的硝基苯进入生产用有机溶剂管线，用于相关化学品的生产；从萃取单元1704出来的萃余相进入生化处理单元1706，经生化处理后达标排放。从过滤单元1703出来的氢氧化铝滤饼进入溶解池1707，在溶解池1707中浓盐酸的作用下40-50℃加热溶解，其中滤饼与浓盐酸的重量比为0.5-2.5:1，随后进入反应釜，与加入反应釜1708的助剂聚合得到水处理药剂聚合氯化铝，所得聚合氯化铝质量满足聚合氯化铝国家标准《GB/T22627-2014水处理剂-聚氯化铝》其中助剂为铝酸钙或/和铝酸镁，其添加量为滤饼干重的2-10wt％。
需要说明的是，本申请实施例的可资源回收的酰基化废水处理工艺，可以借助本申请实施例的可资源回收的酰基化废水处理系统进行酰基化废水的处理并回其资源，但其实现所借助的系统装置并不局限于本申请实施例的可资源回收的酰基化废水处理系统。
下面结合具体实施例来说明本申请实施例的可资源回收的酰基化废水处理工艺。
本申请实施例中所涉及到的原料试剂和设备，如无特殊说明，均为可通过商业途径获得的试剂和设备；本申请实施例中所涉及到的检测方法等，如无特殊说明，均为常规方法。
本申请以下实施例均在实验室条件下进行，抽滤可以采用由布氏漏斗、抽滤瓶、胶管、抽气泵、滤纸组装成的实验室抽滤装置。
下面结合实施例详细描述本申请实施例的用于配制酰化液的方法。
实施例1
<酰化液的配置>
向三口烧瓶中加入600.4g硝基苯，在氮气保护下称量200.41g氯化铝，加入第一腔室110的硝基苯中，利用搅拌器3搅拌，利用换热介质加热至60℃，300rpm下搅拌溶解得到混合溶液，在混合溶液温度为50℃的条件下，利用加料泵6滴加120.13g丙酰化剂，滴加速度为1秒5滴，300rpm下搅拌，滴加完毕后，得到配制后的酰化液。该酰化液在氮气保护下，利用抽滤装置进行抽滤处理，除去酰化液中固体颗粒得到酰化液，并在储液灌中储存30天，其固含量小于0.2％。
<酰化液连续合成酰基萘>
用计量泵分别吸取实施例1中的原料液14和酰化液15，其中酰化液15的(计量泵)流速为160g/min，原料液14的(计量泵)流速为83g/min，即按照摩尔比为2-甲基萘：丙酰氯：AlCl 3＝1：1.3：1.5进行配料。将两种物料通过注射器注入内径为3mm的三通型混合器11里进行混合，混合后进入内径为3mm的微通道反应器12反应，5min后酰化反应液从微通道反应器12的出口流出，其中三通型混合器11和微通道反应器12放置于第一恒温槽中，控制反应温度为-5℃，流入多釜串联的釜式反应器13中，釜式反应器13放置于第二恒温槽中，控制反应温度为40℃，在釜式反应器13中搅拌反应，经过3个釜式反应器13的总停留时间为60min，从釜式反应器13中流出的酰化反应液与除盐水在管式反应器中进行水解反应。
即在管式反应器中进行水解反应方法；在水解段中，提前开启注水泵，在酰化反应液通入到水解段之前就将去离子水通入水解段的水相管路，控制水相管路内水的流速为10mL/min。在水流出水解段出口的情况下，再将从釜式反应器13中流出的酰化反应液经过管路上的单向阀，直接进入水解段的油相管路，与水在低温冷浴器中0℃迅速混合后进入管式反应器，30℃下进行水解反应，同时开启超声震荡对管式反应器内的液体震荡搅拌，并开启末端的分液器16的抽气泵。终止酰化反应并将得到的水解液进入分液器16，在分液器16中80℃下搅拌，4h后即可完成油相和水相的分层，分液后的油相溶液待精馏，水相溶液经过废水高级氧化处理后，送入园区污水处理厂，该酰化反应的2-甲基萘的转化率大于99.0％，选择性为88.0％。
水解反应分离得到的油相进行精馏的方法，包括以下步骤,如图6所示：
(1)油相从第一精馏塔中部送料至第一精馏塔中进行精馏，该酰化反应液中含有80wt％硝基苯溶剂和20wt％的2-甲基-6-丙酰基萘粗馏分。第一精馏塔的压力为0.1kPa、精馏温度为45-130℃，硝基苯在45℃从塔顶蒸出，回流比为1:1，塔顶冷凝温度为10℃，塔顶气相硝基苯经冷凝后从塔顶回收硝基苯溶剂C，釜底液B的温度为130℃，从塔底流出。经过第一精馏塔精馏后，98％的硝基苯溶剂C被回收，釜底液B中含有78％的粗2-甲基-6-丙酰基萘。
(2)步骤(1)中的釜底液B出料温度为130℃，将其趁热泵送至第二精馏塔中进一步精馏。第二精馏塔的压力为1kPa、精馏温度为110-150℃，回流比为5:1，冷凝温度为75℃，17wt％的2-甲基-6-丙酰基萘的同分异构体轻馏分D在114℃从塔顶蒸出，经冷凝后收集；80wt％的产品E2-甲基-6-丙酰基萘在136℃从第二精馏塔测线采出；3％的塔釜重组分F进入塔底，并由塔底再沸器加热后趁热泵送至塔釜料收集罐中。
第一精馏塔的理论塔板数为30，进料位置位于第15块理论板；第二精馏塔的理论塔板数为40，进料位置位于第25块理论板，产品测线采出位于第22块理论板处。
粗产品2,6-MPN进行重结晶的方法为：将甲醇水溶液(其中甲醇：水的质量比为85:15)和粗产品2,6-MPN按8:1的质量比例，在装有搅拌、温度计、回流冷凝管的三口烧瓶内混合，55℃水浴中搅拌至淡黄色固体全溶，继续搅拌20分钟后，将重结晶液在10℃下结晶出来，结晶6小时。结晶完全后用进行抽滤，固液分离，可得出白色细小粉末2-甲基-6-丙酰基萘，该产物置于干燥器中以除去溶剂，重结晶得到的2,6-MPN浆料在产品过滤机分离滤饼与滤液，母液、冲洗液进入甲醇回收塔，顶部馏出甲醇再次利用；釜底馏出物进入污水系统；湿滤饼采用干燥机进行干燥，干燥机排气进入甲醇回收系统，干燥后得到产品2,6-MPN。
采用实施例1方法中配制的酰化液与配制好的2-甲基萘原料液14混合，在微通道与釜式联用反应器中进行酰基化反应，反应后进行水解、减压精馏、重结晶后得到2-甲基-6-丙酰基萘产品，收率为80％,纯度为98.0％。
其中实施例2至实施例7中，以及对比例1至对比例9中，除了按照表1调整酰化液的配制、酰化水解反应条件和油相精馏过程以外，其余与实施例1相同。各实施例以及对比例中得到的酰化液固含量、酰化反应的2-甲基萘的转化率，选择性以及重结晶后得到2-甲基-6-丙酰基萘产品收率和纯度结果如下表1所示。
表1各实施例和对比例中不同阶段性能检测结果
Figure PCTCN2022115478-appb-000002
Figure PCTCN2022115478-appb-000003
Figure PCTCN2022115478-appb-000004
Figure PCTCN2022115478-appb-000005
Figure PCTCN2022115478-appb-000006
Figure PCTCN2022115478-appb-000007
Figure PCTCN2022115478-appb-000008
如表格1所示，实施例2中的数据与实施例1相较未利用抽滤装置进行抽滤处理，将酰化液中未溶解的固体颗粒全部去除，实施例2中酰化液在后期酰化反应阶段有阻塞微通道反应器12的风险，从而降低了重结晶后得到2-甲基-6-丙酰基萘产品收率。此外，配制酰化液15的过程中保护气氛和控制参数也极为重要，例如通过对比例1和实施例1对比可知，在称量氯化铝时不采用氮气保护下进行，可观察到瓶口冒出白烟，氯化铝颜色由浅黄色变成白色，说明氯化铝表面存在失活，因此由于AlCl 3催化剂遇空气中的水分易失活，导致酰化反应阶段中的2-甲基萘转化不完全。另外，在酰化反应阶段中由实施例1和对比例5-7的对比可知，丙酰氯的量加入较少也会导致2-甲基萘转化不完全，而丙酰氯的量和催化剂的量加入较多的情况下并不会影响2-甲基萘转化，综上，丙酰氯的添加量较少，催化剂的量较少及催化温度不适宜均会导致2-甲基萘转化不完全。
如表格中的实施例1和实施例4-7所示，展示了不同的酰化反应条件对酰化反应过程的影响，其中实施例4-7及对比例2-8中与实施例1相比较，在酰化过程中分别改变三通型混合器11和微通道反应器12放置于第一恒温槽中的温度、釜式反应器13放置于第二恒温槽中控制反应温度、串联的釜式反应器13的个数以及停留时间以及2-甲基萘：丙酰氯：AlCl 3的比值关系从而得出了不同的酰化反应条件均影响2-甲基萘的选择性和重结晶后得到2-甲基-6-丙酰基萘产品收率。
其中三通型混合器11和微通道反应器12放置于第一恒温槽中低温混合不仅可以在原料液14和酰化液15混合时剧烈放热时低温移热，而且使得酰化反应在低温微反应器区受动力学控制，反应较快，使得60％-70％的2-甲基萘转化为酰化产物；而在釜式反应器13升高温度，可将混合液在釜式反应器13中反应一段时间从而将未反应2-甲基萘的继续反应，提高反应的转化率，并同时在工业上可降低全部使用微通道反应器12的成本。通过第一恒温槽和第二恒温槽温度的精准控制保证反应较高的转化率和反应较高的选择性。例如表中将实施例1与对比例2对比可知，改变第一恒温槽的温度可严重影响2-甲基萘的选择性，而通过实施例1与对比例3对比可知，改变第二恒温槽的温度可严重影响2-甲基萘的选择性和2-甲基-6-丙酰基萘产品收率。
而将实施例1与对比例4对比可知，利用5个釜式反应器13并且总停留时间为90min，其与实施例1的数据几乎相同，只是延长了实验时间。通过实施例4-7与对比例8的对比可知，对比例8中只利用内径为3mm的微通道反应器12反应未与釜式反应器13联用，造成2-甲基萘的选择性较低，得出虽然微通道反应器12比表面积高、安全性能高、温度响应快、以扩散传质为主，并可以快速将热量从反应体系移除，并且保证反应物在最佳反应物温度下进行反应，有利于反应物的选择性和收率的提高，但只利用微通道反应器12造成酰化反应的时间较长，不利于后期的水解。
对比例9中从釜式反应器中流出的酰化反应液进行间歇釜式水解，其与实施例1相较，并未将从釜式反应器13中流出的酰化反应液直接通入管式反应器中直接进行同步水解反应，即先得到酰化反应液后，到进行水解反应前有一定的时间间隔，在储罐中直接加水进行水解，酰化反应液不能及时水解，放置时易与空气中的水发生水解反应，自身还会有HCl气体溢出，污染空气，且不利于调控水解反应时间、温度及物料平衡率，且还会增加2-3倍水的消耗量。对比例10中相较于实施例1从釜式反应器中流出的酰化反应液与除盐水在连续同步水解反应器进行水解反应，终止酰化反应并将得到的水解液不利用实施例1中的分液器进行分液，而利用静止分离24h后将油相和水相分层，不仅需要至 少24小时的较长分液时间，而且分液过程中产生的废气污染环境。
此外实施例8-12中示例性展示了水解液进入分液器16，分液后的水相溶液经过酰基化废水的处理工艺进行处理具体工艺。
实施例8
取500mL酰基化废水(COD：6530mg/L，硝基苯：3200mg/L，氯离子：3000mg/L，铝离子：3000mg/L)于烧杯中，加入8g氢氧化钠，调节pH至9.0，反应0.5h，得到含有氢氧化铝沉淀的悬浊液；随后将该悬浊液抽滤后得到滤饼和滤液，滤饼自然晾干。用18wt％的盐酸调节滤液pH至2，并将滤液转入装有500mL正庚烷的分液漏斗中，摇匀后静置30min分层。上清液(萃取相)转入填料塔中精馏回收(塔板数为15块，回流比为0.06，塔顶温度为98℃，塔底温度为176℃)，得到纯度99％的硝基苯原料和纯度98％的萃取剂正庚烷，硝基苯回收率达到90％，萃取剂正庚烷的损失量仅为3％。分液漏斗下层溶液(萃余相)转入活性污泥处理系统中进一步处理10h。将10000mg干燥的滤饼加入盛有5000mL浓盐酸的烧杯中于40℃溶解后，在70℃和0.5MPa下与500mg铝酸钙反应，得到产品聚合氯化铝。
经检测，经活性污泥处理系处理后，出水中COD＝55mg/L(如图9所示)，硝基苯浓度＝0.5mg/L，均满足排放标准《GB8978-1996污水综合排放标准》(COD<100mg/L，硝基苯<1.0mg/L)的要求；所获得的产品聚合氯化铝的主要指标为：氯化铝质量分数为29％，盐基度为55％，水不溶物的质量分数为0.2，pH值为4.5，铁质量分数为1.0％，砷、铅、铬、汞和镉均未检出，产品聚合氯化铝的各项指标均符合国家标准《GB/T22627-2014水处理剂-聚氯化铝》的要求。
实施例9
取500mL酰基化废水(COD：6900mg/L，硝基苯：2700mg/L，氯离子：3000mg/L，铝离子：3000mg/L)于烧杯中，加入7.5g氢氧化钾，调节pH至9.2，反应0.5h，得到含有氢氧化铝沉淀的悬浊液；随后将该悬浊液抽滤后得到滤饼和滤液，滤饼自然晾干。用50wt％的硫酸调节滤液pH至2.5，并将滤液转入装有600mL正庚烷烷的分液漏斗中，摇匀后静置30min分层。上清液(萃取相)转入精馏塔中精馏回收(塔板数为15块，回流比为0.06，塔顶温度为98℃，塔底温度176℃)，得到纯度99％的硝基苯原料和纯度98％的萃取剂正庚烷，硝基苯回收率达到92％，萃取剂正庚烷的损失量仅为4％。分液漏斗下层溶液(萃余相)转入活性污泥处理系统中进一步处理12h。将10000mg干燥的滤饼加入盛有6000mL浓盐酸的溶解槽中45℃溶解后，在80℃和0.7MPa下与400mg铝酸镁反应，得到产品聚合氯化铝。
经检测，经活性污泥处理系处理后，出水中COD＝50mg/L(如图10所示)，硝基苯浓度＝0.5mg/L，均满足排放标准《GB8978-1996污水综合排放标准》(COD<100mg/L，硝基苯<1.0mg/L)的要求；所获得的产品聚合氯化铝的主要指标为：氯化铝质量分数为30％，盐基度为56％，水不溶物的质量分数为0.3，pH值为5.0，铁质量分数为1.0％，砷、铅、铬、汞和镉均未检出，产品聚合氯化铝的各项指标均符合国家标准《GB/T22627-2014水处理剂-聚氯化铝》的要求。
实施例10
取500mL酰基化废水(COD：7200mg/L，硝基苯：4200mg/L，氯离子：3000mg/L，铝离子：2500mg/L)于烧杯中，加入8.2g氢氧化钾，调节pH至9.0，反应0.5h，得到含有氢氧化铝沉淀的悬浊液；随后将该悬浊液抽滤后得到滤饼和滤液，滤饼自然晾干。用15wt％的盐酸调节滤液pH至3，并将滤液转入装 有600mL正庚烷的分液漏斗中，摇匀后静置30min分层。上清液(萃取相)转入精馏塔中精馏回收(塔板数为15块，回流比为0.06，塔顶温度为98℃，塔底温度176℃)，得到纯度99％的硝基苯原料和纯度99％的萃取剂正庚烷，硝基苯回收率达到88％，萃取剂正庚烷的损失量仅为6％。分液漏斗下层溶液(萃余相)转入活性污泥处理系统中进一步处理12h。将8000mg干燥的滤饼加入盛有5000mL浓盐酸的溶解槽中于50℃溶解后，在80℃和0.7MPa下与450mg铝酸镁反应，得到产品聚合氯化铝。
经检测，经活性污泥处理系处理后，出水中COD＝55mg/L(如图11所示)，硝基苯浓度＝0.5mg/L，均满足排放标准《GB8978-1996污水综合排放标准》(COD<100mg/L，硝基苯<1.0mg/L)的要求；所获得的产品聚合氯化铝的主要指标为：氯化铝质量分数为31％，盐基度为62％，水不溶物的质量分数为0.4，pH值为6.5，铁质量分数为2.0％，砷、铅、铬、汞和镉均未检出，产品聚合氯化铝的各项指标均符合国家标准《GB/T22627-2014水处理剂-聚氯化铝》的要求。
实施例11
本申请的实施例与实施例18基本相同，不同之处在于，萃取剂为四氯化碳。
实施例12
本申请的实施例与实施例18基本相同，不同之处在于，萃取剂为正辛烷和二甲苯按体积比1:1混合的混合物。

Claims (40)

  1. 一种酰化液，其中，利用配置系统进行配置，所述配置系统包括：
    釜体，所述釜体限定出第一腔室，所述釜体具有与所述第一腔室连通的第一加料口、第二加料口、搅拌口和出液口；
    加料口盖，所述加料口盖可拆卸地密封盖装在所述第一加料口上；
    加料泵，所述加料泵与所述第二加料口密封相连，以便向所述第二加料口内添加酰化剂；和搅拌器，所述搅拌器的一部分密封插装在所述搅拌口内，所述搅拌器包括搅拌轴和搅拌叶片，所述搅拌叶片设置在所述搅拌轴上，所述搅拌轴的至少一部分和所述搅拌叶片中的每一者设置在所述第一腔室内；
    利用配置系统进行配置所述酰化液包括以下步骤：
    a、在惰性气体保护下称量Lewis催化剂；
    b将溶剂和步骤a称量的所述Lewis催化剂通过所述第一加料口加入所述第一腔室内并利用所述搅拌器搅拌均匀得到混合溶液；
    c、利用所述加料泵向步骤b得到的所述混合溶液中加入酰化剂，得到酰化液。
  2. 根据权利要求1所述的酰化液，其中，所述釜体包括：
    第一壳体，所述出液口设置在所述第一壳体上，所述第一壳体的上端敞开；
    上盖，所述上盖密封盖装在所述第一壳体上，所述第一壳体和所述上盖限定出所述第一腔室，所述第一加料口、所述第二加料口和所述搅拌口均设置在所述上盖上；和
    夹套，所述夹套套设在所述第一壳体上，所述夹套具有供换热介质进入的换热介质第一进口和供换热介质流出的换热介质第一出口。
  3. 根据权利要求1所述的酰化液，其中，所述配置系统还包括加热装置，所述加热装置具有换热介质第二进口和换热介质第二出口，所述换热介质第一出口与所述换热介质第二进口相连，所述换热介质第一进口与所述换热介质第二出口相连；
    温度传感器，所述釜体具有与所述第一腔室连通的测温口，所述温度传感器的一部分密封插装在所述测温口内，所述温度传感器的检测端设置在所述第一腔室内；和
    控制器，所述控制器与所述加热装置和所述温度传感器中的每一者相连，以便所述控制器根据所述温度传感器检测到的温度控制所述加热装置。
  4. 根据权利要求2所述的酰化液，其中，所述配置系统还包括抽滤装置，所述抽滤装置包括：
    过滤装置，所述过滤装置包括第二壳体和过滤膜，所述第二壳体限定出第二腔室，所述过滤膜设置在所述第二腔室内，所述过滤膜将所述第二腔室分隔成第一部分和第二部分，所述第二壳体上设有与所述第二腔室连通的过滤装置进口和过滤装置出口，所述过滤装置进口对应所述第一部分设置，所述过滤装置出口对应所述第二部分设置，所述过滤装置进口与所述出液口相连；
    储液罐，所述储液罐包括第三壳体，所述第三壳体限定出第三腔室，所述第三壳体上设有储液罐进口和储液罐出口，所述储液罐进口与所述过滤装置出口连通；和
    抽真空泵，所述抽真空泵具有第一抽真空口，所述第三壳体上设有与所述第三腔室连通的第二抽真空口，所述第一抽真空口与所述第二抽真空口连通。
  5. 根据权利要求4所述的酰化液，其中，所述第一壳体、所述第二壳体和所述第三壳体为一体式结构，其中所述出液口设置在所述第一壳体的底部，所述过滤装置进口设置在所述第二壳体的顶部，所述过滤装置设置在所述釜体的底部，所述过滤装置出口设置在所述第二壳体的底部，所述储液罐进口设置在所述第三壳体的顶部，所述储液罐设置在所述过滤装置的底部。
  6. 根据权利要求1-5任一所述的酰化液，其中，步骤b中溶剂为硝基苯，步骤b具体为：
    首先将所述硝基苯通过所述第一加料口加入所述第一腔室内；
    然后将步骤a称量的所述Lewis催化剂加入所述第一腔室内所述硝基苯中；
    利用换热介质将所述硝基苯和所述Lewis催化剂共同加热至50℃-60℃，同时利用所述搅拌器在200rpm-400rpm下搅拌。
  7. 根据权利要求4所述的酰化液，其中，还包括步骤d，在惰性气体氛围下，利用所述抽滤装置对步骤c得到酰化液进行抽滤处理，以便得到过滤后酰化液。
  8. 根据权利要求1所述的酰化液，其中，步骤c中，所述酰化剂的加入速度为每秒3滴-10 滴，所述搅拌器的搅拌速度为200rpm-400rpm。
  9. 根据权利要求1所述的酰化液，其中，所述酰化剂与所述Lewis催化剂的摩尔质量比为(1.1-1.5):(1.3-1.7)；和/或
    所述Lewis催化剂与所述溶剂的摩尔质量比为(1.3-1.7)：5。
  10. 一种连续合成酰基萘的方法，其中，包括如下步骤：
    S1：将含2-甲基萘的原料液和如权利要求1-9中任一所述的酰化液混合，其中所述2-甲基萘：所述酰化剂：所述Lewis催化剂按照摩尔比为1：1.3：1.5形成酰化反应液；
    S2：所述酰化反应液依次进入微通道反应器和多釜串联的釜式反应器中进行酰化反应并同步进行水解以及后续通过精馏及结晶得到2-甲基-6-丙酰基萘。
  11. 根据权利要求10所述的方法，其中，所述微通道反应器设有若干个并相互并联。
  12. 根据权利要求10所述的方法，其中，步骤S1中，将所述原料液和所述酰化液通过注射器注入三通型混合器里进行混合。
  13. 根据权利要求12所述的方法，其中，所述三通型混合器为T型混合器或Y型混合器。
  14. 根据权利要求12所述的方法，其中，所述三通型混合器和所述微通道反应器放置于第一恒温槽中，温度控制在-5-0℃，所述釜式反应器放置于第二恒温槽中，温度控制在30-50℃。
  15. 根据权利要求10所述的方法，其中，所述釜式反应器设有2-4个且相互串联。
  16. 根据权利要求10所述的方法，其中，所述酰化反应液在所述釜式反应器中的总停留时间为50-80min。
  17. 根据权利要求10所述的方法，其中，所述微通道反应器和所述釜式反应器的材质均为耐强酸腐蚀材料。
  18. 根据权利要求10所述的方法，其中，所述微通道反应器的内径为0.5-3.175毫米。
  19. 根据权利要求10所述的方法，其中，所述酰化液中的酰化剂为丙酰氯、乙酰氯、乙酸酐、丙酸酐中的任一种。
  20. 根据权利要求10所述的方法，其中，所述水解方法为在水解段中，将酰化反应后的酰化反应液通入到水解段之前将水通入水解段的水相管路，在水流出水解段出口的情况下，再将所述酰化反应液通入到所述水解段的油相管路，并与水在低温冷浴器中迅速混合后进入管式反应器进行水解反应，水解后的混合液从所述管式反应器的出口流出，并采用分液器收集混合液进行水相和油相的分离；在此阶段，水始终保持流通状态直至混合液完全排出，实现水解反应与酰化反应同步进行。
  21. 根据权利要求20所述的方法，其中，通过开启注水泵向所述水解段的所述水相管路通水，控制所述水相管路内水的流速为3-15mL/min。
  22. 根据权利要求20所述的方法，其中，所述酰化反应液经过管路上的单向阀，直接进入水解段的油相管路。
  23. 根据权利要求22所述的方法，其中，所述单向阀为不锈钢卡套单向阀，内部流道的材质为聚四氟乙烯。
  24. 根据权利要求20所述的方法，其中，低温冷浴器的温度为0-20℃。
  25. 根据权利要求20所述的方法，其中，所述管式反应器的反应温度为30-40℃。
  26. 根据权利要求20所述的方法，其中，所述分液器包括容器、搅拌装置、进料口、出料口和排气口，所述搅拌装置伸入所述容器内并与外置的减速器连接进行驱动，所述出料口连通于所述容器的底端，所述进料口和所述排气口共同连通于所述容器上端；所述进料口通过管路接通水解段的所述管式反应器的出口，所述排气口通过管线接入一个抽气泵；所述进料口、所述出料口和所述排气口处均设有控制阀。
  27. 根据权利要求26所述的方法，其中，水解段的所述管式反应器中进行水解反应时，同时所述抽气泵将产生的废气抽出。
  28. 根据权利要求20所述的方法，其中，还包括步骤S3：酰基化废水处理；具体包括以下步骤
    (1)调节所述分液器分离出的所述水相的pH至碱性，获得含有氢氧化铝沉淀的悬浮液，并将所述悬浮液过滤得到氢氧化铝滤饼和滤液；
    (2)将所述滤液调节pH至酸性，随后转入萃取剂中搅拌、静置，获得分层的萃余相和萃取 相；
    (3)将所述萃余相经生化处理后达标排放；
    (4)将所述萃取相精馏分离后获得硝基苯和所述萃取剂，将所述硝基苯作为有机溶剂回用，将所述萃取剂重新用于所述滤液萃取；
    (5)将所述滤饼加入浓盐酸加热溶解，添加助剂后聚合得到聚合氯化铝。
  29. 根据权利要求28所述的方法，其中，步骤(1)中，调节所述分液器分离出的所述水相的pH至碱性的方法包括：采用碱原料调节所述水相的pH至8-10，所述碱原料为氢氧化钠、氢氧化钾或液氨中的一种或两种以上。
  30. 根据权利要求28所述的方法，其中，步骤(2)中，将所述滤液调节pH至酸性，包括：采用酸调节滤液的pH至2-3，其中酸为盐酸、硝酸或硫酸中的一种或两种以上。
  31. 根据权利要求28所述的方法，其中，步骤(2)中，所述萃取剂为非极性有机溶剂，且所述萃取剂与滤液的体积比为0.5-5:1。
  32. 根据权利要求31所述的方法，其中，所述萃取剂为正庚烷、正辛烷、正己烷、苯、甲苯、二甲苯、四氯化碳中的至少一种。
  33. 根据权利要求28所述的方法，其中，步骤(3)中，生化处理采用活性污泥法；步骤(5)中，加热溶解的温度为40-50℃，滤饼与浓盐酸的重量比为0.5-2.5:1。
  34. 根据权利要求28所述的方法，其中，步骤(5)中，所述助剂的添加量为滤饼干重的2-10wt％，所述助剂为铝酸钙或/和铝酸镁。
  35. 根据权利要求20所述的方法，其中，精馏包括以下步骤：
    (1)将所述油相从第一精馏塔中部送料至第一精馏塔中进行精馏，轻馏分从塔顶蒸出，经冷凝后回收，塔釜液从塔底部流出；
    (2)将步骤(1)中的塔釜液趁热泵送至第二精馏塔中进一步精馏，第二精馏塔的轻馏分从塔顶蒸出，经冷凝后收集；产品从第二精馏塔测线采出；塔釜重组分进入塔底，并由塔底再沸器加热后，将塔釜重组分趁热泵送至塔釜料收集罐中。
  36. 根据权利要求35所述的方法，其中，所述第一精馏塔的压力为0.05KPa-10KPa；回流比为(1-2)：1。
  37. 根据权利要求35或36所述的方法，其中，所述第一精馏塔的冷凝温度为10-20℃。
  38. 根据权利要求35所述的方法，其中，所述第二精馏塔的压力为0.05KPa-10KPa；回流比为(5-10)：1。
  39. 根据权利要求35所述的方法，其中，所述第二精馏塔的冷凝温度为50-90℃。
  40. 根据权利要求10-39任一所述的方法，其中，所述2-甲基萘为从洗油中提取，纯度为99.0-99.9％；其中提取方法为：
    I洗油通过精馏分离得到甲基萘富集馏分；
    II将所述甲基萘富集馏分通入共沸精馏塔进行共沸精馏，得到共沸馏出物；
    III将所述共沸馏出物通入分离器得到2-甲基萘粗品；
    IV将所述2-甲基萘粗品通入多个并联设置的间歇式熔融结晶器进行2-甲基萘的结晶提纯。
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