A continuous production process of N-fatty acyl amino acid compounds

Abstract

The application relates to the technical field of chemical synthesis, and particularly discloses a continuous production process of N-fatty acyl amino acid compounds, which adopts a two-stage acyl chloride feeding mode: first, 45% to 55% of total feeding amount of fatty acyl chloride is subjected to a first condensation reaction with an amino acid salt aqueous solution to obtain a material liquid 1; then, the remaining acyl chloride, the material liquid 1 and liquid alkali are subjected to a second condensation reaction, and then the target product is obtained through steps of continuous acidification, continuous liquid separation and continuous washing. Through accurate control of material ratio and liquid alkali equivalent of the two-step reaction, the hydrolysis side reaction of the fatty acyl chloride is significantly inhibited, the reaction period is shortened from the hour level to the minute level, the product conversion rate can reach more than 99%, the free fatty acid content is lower than 1%, the color is significantly better than that of the same kind of products on the market, the product has excellent industrial amplification potential, and can be widely applied to the fields of daily chemicals, medicines and food and other fields with high requirements on the purity of raw materials.

Description

Technical Field

[0001] This invention relates to the field of chemical synthesis technology, and in particular to a continuous production process for N-fatty acyl amino acid compounds. Background Technology

[0002] N-Fatty acyl amino acid surfactants are a new type of green and environmentally friendly surfactant made from renewable biomass, serving as an upgraded alternative to traditional petroleum-based surfactants. These compounds contain both carboxyl anionic and ammonium cationic groups in their molecular structure, making them typical amphoteric substances. They possess numerous excellent properties, including renewability, low irritation, low toxicity, good biocompatibility, and excellent biodegradability and environmental compatibility, making them a preferred category of surfactants for the food, pharmaceutical, and cosmetic industries. Due to the rich structural types of amino acids and the ability to control product structure by adjusting the length and number of fatty acid chains, these surfactants exhibit broad structural diversity, displaying differentiated physicochemical and biological properties, thus providing ample room for expanding their application scenarios.

[0003] Currently, the preparation processes of N-fatty acyl amino acid compounds mainly include three technical routes: The first method is the fatty acid methyl ester method. The process involves first preparing fatty acids into corresponding fatty acid methyl esters, and then obtaining the target product through an amino-ester transesterification reaction under the action of a catalyst. This process requires an additional catalyst, which not only increases the complexity of the process but also necessitates a catalyst recovery and reuse unit, thus limiting its industrial application.

[0004] The second method is the high-temperature melt direct condensation method, which uses fatty acids and amino acids as raw materials. The mixture undergoes direct dehydration condensation under high-temperature conditions, eliminating the need for acyl chloride raw materials and organic solvents. The reaction temperature is typically 180-220℃, and the catalyst (such as magnesium oxide or zinc oxide) is used at 2%-5% of the total raw material mass. The resulting product is free of inorganic salts and has a purity exceeding 98%. The disadvantages of this method are the high requirements for equipment materials, necessitating corrosion resistance, such as Hastelloy equipment, resulting in significant equipment investment. Furthermore, the high reaction temperature leads to higher energy consumption, 20%-30% higher than traditional processes. Additionally, this process is not suitable for thermally unstable amino acids, such as glutamine, which easily decomposes at high temperatures, limiting the range of applicable raw materials.

[0005] The third method is the fatty acyl chloride process, based on the Schotton-Baumann condensation reaction. It uses fatty acyl chlorides and amino acids or their salts as raw materials, undergoing a condensation reaction under alkaline conditions. This is currently the mainstream process for industrial production. Existing industrial production often employs a batch operation mode, which suffers from problems such as long reaction cycles, cumbersome operation, and poor batch-to-batch product quality stability.

[0006] To address the shortcomings of batch processes, the industry has proposed various continuous production improvement schemes. Some processes employ a continuous production method combining an amidation column and an aging unit. A portion of the reactants is recycled back to the mixing unit for further reaction, while the other portion is heated by a heater and then enters the aging unit to complete the aging reaction. The mass ratio of recycled material to material entering the aging unit is 2-20:1. After the aging reaction, N-fatty acyl amino acid salts are obtained. Although this process achieves continuous operation, the overall production efficiency is low because most of the material is in a recycling state and the proportion of material entering the aging reaction unit is low. Furthermore, the aging reaction time still requires 4-12 hours, and the reaction cycle has not been significantly shortened.

[0007] Another continuous flow preparation process divides the amidation reaction into two temperature zones, each equipped with different continuous reaction equipment: the first temperature zone is 0~15℃, using a microchannel reactor; the second temperature zone is 70~85℃, using a tubular reactor; the post-treatment stage uses membrane separation to separate the by-product salt solution and the product solution. This process has the following shortcomings: First, the continuous operation process with two temperature controls is complex and difficult to control, and the investment cost of the microchannel reaction equipment is high; second, the 70~85℃ holding temperature reaction conditions easily cause hydrolysis of fatty acyl chlorides, which not only increases the content of by-products but also wastes raw materials; third, membrane separation can only remove inorganic salts from the system and cannot effectively remove organic impurities. Excessive unreacted amino acid salts remain in the product, which can easily lead to high viscosity, color, and total solids content, making it difficult to meet the quality requirements of high-end applications such as daily chemicals.

[0008] There are also related continuous reaction processes that achieve continuous reaction through the combination of pipeline synthesis units and neutralization reactors, with product conversion rates reaching over 96%. However, this process involves pre-mixing liquid alkali with amino acid salts before reacting with fatty acyl chlorides. This results in excessively high local alkali concentrations within the system, which can easily lead to the hydrolysis of fatty acyl chlorides to generate free fatty acids. Consequently, the product contains a high level of free fatty acid impurities, making it impossible to prepare high-quality N-fatty acyl amino acid surfactants.

[0009] In summary, the existing continuous production process of fatty acyl chloride still suffers from problems such as low production efficiency, numerous side reactions of acyl chloride hydrolysis, high product impurity content, large equipment investment, and complex process control. There is an urgent need to develop a continuous process that is more suitable for large-scale industrial production. Summary of the Invention

[0010] To address the problems of low production efficiency, numerous hydrolysis side reactions of acyl chlorides, high product impurity content, large equipment investment, and complex process control in existing continuous production processes using fatty acyl chlorides, this invention provides a continuous production process for N-fatty acyl amino acid compounds.

[0011] To solve the above-mentioned technical problems, the technical solution provided by the present invention is as follows: A continuous production process for N-fatty acyl amino acid compounds includes the following steps: S1, 45% to 55% of the total feed amount of fatty acyl chloride is continuously mixed with the aqueous solution of all amino acid salts and the first condensation reaction is carried out to obtain feed solution 1; S2, the remaining fatty acyl chloride, the feed solution 1 and liquid alkali are continuously mixed and a second condensation reaction is carried out to obtain feed solution 2; S3, the liquid 2 and the acid solution are mixed and acidified to obtain liquid 3; S4, the liquid 3 is continuously separated to obtain an oil phase liquid 4 and a salt-containing phase; S5, the oil phase liquid 4 is continuously washed to obtain N-fatty acyl amino acid compounds; In S2, the molar ratio of the remaining fatty acyl chloride to the alkali in the strong alkali solution is 1:0.9 to 1:1.05.

[0012] Compared to existing technologies, the continuous production process for N-fatty acyl amino acid compounds provided by this invention employs a unique two-stage fatty acyl chloride feeding design. First, 45%–55% of the total fatty acyl chloride feed is reacted with the entire amino acid salt aqueous solution in a first condensation reaction. In this first step, excess amino acid salt maintains a weakly alkaline environment in the reaction system (pH stable at 9–11), promoting preferential amidation of the fatty acyl chloride and effectively suppressing hydrolysis side reactions, thus reducing the formation of free fatty acid impurities. In the second step, the remaining fatty acyl chloride and a strong alkaline solution are added, with the molar ratio of alkali to alkali controlled at 1:0.9–1:1.05, further reducing the concentration of free alkali in the reaction system and inhibiting hydrolysis of fatty acyl chloride at its source. Subsequent continuous acidification and separation, along with continuous washing processes, remove byproduct inorganic salts, unreacted amino acids, and water-soluble impurities from the system, ultimately yielding the N-fatty acyl amino acid product.

[0013] This invention enables continuous production of N-fatty acyl amino acid compounds with high production efficiency, reducing the reaction cycle from hours to minutes, ensuring consistent product quality across batches, and minimizing impurities. Testing shows that the product conversion rate can reach over 99%, the free fatty acid content is less than 1%, and the color is low. This meets the high-end application requirements of daily chemicals, pharmaceuticals, and other fields, and is suitable for industrial-scale production.

[0014] As a specific embodiment of the present invention, the continuous production process of the N-fatty acyl amino acid compound specifically includes the following steps: S1, 45% to 55% of the total fatty acyl chloride feed amount is transported to a static mixer for mixing with all the amino acid salt aqueous solution, and then enters the first tubular reactor for the first condensation reaction to obtain feed liquid 1; S2, the remaining fatty acyl chloride, the feed solution 1 and the strong alkali solution are transported to a static mixer for mixing, and then enter the second tubular reactor for the second condensation reaction to obtain feed solution 2; S3, the feed liquid 2 and the acid solution are transported to a static mixer for mixing, and then enter the third tubular reactor for acidification reaction to obtain feed liquid 3; S4, the liquid 3 is fed into a continuous liquid separator to continuously separate the oil phase liquid 4 and the salt-containing phase; S5, the oil phase liquid 4 is continuously washed with water through a continuous washing device to obtain N-fatty acyl amino acid compounds.

[0015] Furthermore, the fatty acyl chloride is a C8~C18 fatty acyl chloride.

[0016] In some specific embodiments of the present invention, the fatty acyl chloride is lauroyl chloride, cocoyl chloride, oleyl chloride, stearoyl chloride, or myristoyl chloride.

[0017] Furthermore, the amino acid salt is at least one selected from sarcosine, glycine, glutamate, alanine, aspartate, or arginine.

[0018] Furthermore, in S1, the mass concentration of the amino acid salt aqueous solution is 20%~30%.

[0019] Furthermore, in S1, the temperature of the first tubular reactor is 30℃~60℃, and the residence time of the material in the first tubular reactor is 0.3min~2min.

[0020] In some specific embodiments of the present invention, before the material in step S1 enters the static mixer, it is first preheated by metering pumps, with fatty acyl chloride and amino acid salt aqueous solution being transported separately to a tubular reaction heat exchanger at a set ratio. The preheating temperature is controlled at 30℃~60℃. The inner diameter of the tubular heat exchanger is 2mm~4mm, and the total tube length is 1m~5m. The preheated material then enters the static mixer for mixing, and then enters the first tubular reactor for the first condensation reaction. The inner diameter of the first tubular reactor is 2mm~4mm, and the total tube length is 5m~10m.

[0021] Furthermore, in S2, the mass concentration of the strong alkaline solution is 10% to 30%.

[0022] Specifically, in S2, the strong alkaline solution is an aqueous solution of sodium hydroxide or an aqueous solution of potassium hydroxide.

[0023] Furthermore, in S2, the temperature of the second tubular reactor is 30℃~60℃, and the residence time of the material in the second tubular reactor is 0.3min~2min.

[0024] In some specific embodiments of the present invention, before the material in step S2 enters the static mixer, the fatty acyl chloride, liquid alkali, and feed liquid 1 are first preheated by metering pumps in a set ratio to a tubular reaction heat exchanger. The preheating temperature is controlled at 30℃~60℃. The inner diameter of the tubular heat exchanger is 2mm~4mm, and the total tube length is 1m~5m. The preheated material then enters the static mixer for mixing, and then enters the second tubular reactor for the second condensation reaction. The inner diameter of the second tubular reactor is 2mm~4mm, and the total tube length is 10m~15m.

[0025] Furthermore, in S3, the acid solution is hydrochloric acid or sulfuric acid.

[0026] Specifically, in S3, the mass concentration of the hydrochloric acid is 30%~36%, and the mass concentration of the sulfuric acid is 30%~98%.

[0027] Furthermore, in S3, the pH of the acidification reaction is 1~4, the temperature is 40℃~70℃, and the residence time of the material in the third tubular reactor is 0.3min~1.5min.

[0028] In some specific embodiments of the present invention, the inner diameter of the third tubular reactor is 2mm to 4mm, and the total length of the tube is 5m to 10m.

[0029] Furthermore, in S4, the temperature of the continuous liquid separation device is 40℃~70℃.

[0030] In some specific embodiments of the present invention, in S4, continuous liquid separation specifically refers to: Oil and water phase materials are pre-placed inside the continuous liquid separator and allowed to stand to form an initial state in which the oil phase is on the upper layer and the water phase is on the lower layer. The feed liquid 3 is continuously fed into the oil-water interface of the continuous liquid separator, and the liquid separation temperature is controlled at 40℃~70℃. The separated oil phase and water phase are continuously pumped out, and the sum of the oil phase extraction speed and the water phase extraction speed is controlled to be equal to the feed speed of liquid 3, so as to keep the original oil and water volume of the liquid in the device basically unchanged. The first batch of liquid 3 entering the continuous liquid separator is not separated and is used as the pre-placed oil-water two-phase material. Subsequent feeds will then begin continuous liquid separation.

[0031] Specifically, the continuous liquid separation device can employ a four-necked flask, which is in contact with a heat-conducting medium to control the separation temperature within a set range. A feed pipe is inserted into one neck of the four-necked flask, with its outlet located at a preset oil-water phase interface height, ensuring that the feed position of the liquid 3 is directly opposite the oil-water phase interface of the pre-set substrate. The other two necks of the four-necked flask serve as the oil phase outlet and the water phase outlet, respectively. An oil outlet pipe is inserted into the oil phase outlet, with its opening located in the middle of the upper oil phase. A bottom-insertion pipe extending to the bottom of the four-necked flask is inserted into the water phase outlet for extracting the lower water phase. The oil outlet pipe and the bottom-insertion pipe are connected to the oil phase metering pump and the water phase metering pump, respectively, via pipelines.

[0032] In some specific embodiments of the present invention, the feed flow rate of liquid 3 is controlled to be 51 g / min to 144 g / min, the oil phase discharge flow rate is 15 g / min to 20 g / min, and the aqueous phase discharge flow rate is 36 g / min to 124 g / min.

[0033] Furthermore, in S5, a countercurrent washing method is adopted, wherein the mass flow rate of the oil phase liquid 4 is 1 to 4 times the mass flow rate of the water phase; and the washing temperature is 40℃ to 70℃.

[0034] In practical implementation, the continuous washing can be achieved using a multi-stage series centrifugal extractor, with 3 to 5 centrifugal extractors operating in series. Specifically, a benchtop centrifugal extractor with a drum diameter of 50mm (e.g., Shandong Liancui LC-50 model) can be selected. During operation, the oil phase feed liquid 4 is sequentially connected to the heavy phase outlet and light phase inlet of each centrifuge, forming a counter-current channel. The washing water enters from the heavy phase inlet of the first centrifuge. Utilizing the centrifugal force field to enhance two-phase mixing and rapid phase separation can effectively shorten the extraction time and improve separation efficiency.

[0035] Alternatively, a multi-stage countercurrent scrubbing tower can be used, employing 2 to 4 extraction towers connected in series. Specifically, glass or stainless steel extraction towers with a height of 50-100 cm and a diameter of 20 mm can be selected. To improve mass transfer efficiency, the scrubbing tower is filled with structured or random packing materials, typically including glass beads, θ rings, or glass springs. The packing material breaks down the dispersed phase, forming fine droplets, thereby significantly increasing the contact surface area between the oil and water phases. During operation, the oil phase feed liquid 4 enters from the bottom of the tower, and the scrubbing water enters from the top. Mass transfer is achieved through sufficient contact within the packing layer. The purified oil phase feed liquid 5 overflows from the top of the tower, while the water phase rich in impurities is discharged from the bottom.

[0036] It should be noted that when the target product is an N-fatty acyl amino acid salt compound, after obtaining the refined oil phase solution 5 in step S5, the neutralization step S6 is continued: the oil phase solution 5 and the strong alkali solution are transported to the static mixer by a metering pump for mixing, and then enter the fourth tubular reactor for continuous neutralization reaction, thus obtaining an aqueous solution of the N-fatty acyl amino acid salt compound.

[0037] The strong alkaline solution is preferably an aqueous solution of sodium hydroxide or potassium hydroxide, with a mass concentration controlled at 10%~30%. The molar ratio of N-fatty acyl amino acids in the oil phase feed liquid 5 to the alkali metal in the liquid alkali is controlled at 1:1~1:1.1. The neutralization reaction temperature is controlled at 50~80℃, and the residence time of the material in the fourth tubular reactor is 0.6~2 min to ensure that the neutralization reaction is complete.

[0038] If the target product is an N-fatty acyl amino acid free acid, step S6 can be omitted, and the oil phase liquid 5 obtained in step S5 can be directly used as the final product. This process route flexibility expands the applicability of the present invention in different application scenarios.

[0039] In the practical industrial application of this invention, the process consisting of steps S1 to S6 possesses a high degree of modular operational flexibility. Each unit operation can be connected in series to achieve continuous production from start to finish, or the feed liquid can be temporarily stored in an intermediate storage tank after any unit operation is completed, or after several unit operations are combined. This combination of continuous and intermittent operation is particularly beneficial for segmented maintenance, online cleaning, and flexible production scheduling of the production unit. For example, when maintenance is required on the downstream separation and purification section (such as the S4 liquid separation equipment or the S5 washing equipment), the upstream reaction section (S1-S3) can continue to operate and store the feed liquid in the tank area, thereby avoiding a complete shutdown of the entire unit and significantly improving the stability of production operation and equipment utilization.

[0040] Furthermore, this invention allows for flexible adjustment of the product output form according to the needs of the end market. For mainstream liquid detergent products in the daily chemical industry, which are typically supplied as salt-type products, step S6 can be performed to prepare N-fatty acyl amino acid sodium or potassium salts by switching the type of liquid alkali (such as sodium hydroxide or potassium hydroxide aqueous solution). However, for some high-end applications with strict requirements on ionic strength, or in cases where solid-state storage is required, step S6 can be omitted, and high-purity N-fatty acyl amino acid free acid can be directly used as the final product.

[0041] In summary, this invention provides a continuous production process for N-fatty acyl amino acid compounds. This process employs a two-stage acyl chloride feeding method: first, 45%–55% of the total feed weight of fatty acyl chloride is reacted with an amino acid salt aqueous solution to form feed solution 1; then, the remaining acyl chloride, feed solution 1, and liquid alkali undergo a second condensation reaction, followed by continuous acidification, continuous separation, and continuous washing to obtain the target product. This invention significantly suppresses the hydrolysis side reaction of fatty acyl chloride by precisely controlling the material ratio and liquid alkali equivalent in the two-step reaction. The reaction cycle is shortened from hours to minutes, the product conversion rate can reach over 99%, the free fatty acid content is less than 1%, and the color is significantly superior to similar commercially available products. It possesses excellent industrial scale-up potential and can be widely applied in fields such as daily chemicals, pharmaceuticals, and food where the purity of raw materials is critical. Detailed Implementation

[0042] To make the objectives, technical solutions, and advantages of this invention clearer, the invention will be further described in detail below with reference to embodiments. It should be understood that the specific embodiments described herein are merely illustrative and not intended to limit the invention.

[0043] To better illustrate the present invention, further examples are provided below.

[0044] Unless otherwise specified, all raw materials used in the following examples and comparative examples are commercially available analytical grade; the conversion rate was determined by liquid chromatography; the free fatty acid content was determined according to GB / T 6365 standard, QPJSW007-2021; and the color was determined according to the platinum-cobalt colorimetric method (GB / T 3143) GB / T 22295-2008.

[0045] Example 1 This embodiment provides a continuous production process for sodium lauroyl sarcosinate, including the following steps: S1, 554.55g of lauroyl chloride and 1885.48g of a 30% sodium sarcosinate aqueous solution were pumped at a set flow rate to a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 1m for preheating. The lauroyl chloride feed rate was 15.00g / min, and the sodium sarcosinate aqueous solution feed rate was 51.00g / min. The preheating temperature was set at 30℃. After the two preheated materials were mixed by a static mixer, they entered the first tubular reactor with an inner diameter of 3mm and a tube length of 6m for reaction. The reaction temperature was controlled at 30℃, and the residence time of the materials in the reactor was 0.8min. The pH value of the output liquid was measured to be 9. The system ran continuously for 37.0min, and liquid 1, totaling 2440.03g, was collected. S2, 563.08g of lauroyl chloride, 2440.03g of solution 1 obtained in step S1, and 563.09g of 19% caustic soda were pumped into a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 1m for preheating via a metering pump at a set flow rate. The feed rates for lauroyl chloride, solution 1, and caustic soda were 15.00g / min, 65.00g / min, and 15.00g / min, respectively. The preheating temperature was set to 30℃. After preheating, the three materials were mixed in a static mixer and then entered a second tubular reactor with an inner diameter of 4mm and a total pipe length of 12m for reaction. The reaction temperature was controlled at 30℃, and the residence time of the materials in the reactor was 1.9min. The pH value of the solution was measured to be 9. The system ran continuously for 37.5min, and solution 2, totaling 3566.20g, was collected. S3, the 3566.20g of liquid 2 obtained in step S2 and 891.55g of 30% hydrochloric acid solution are delivered by a metering pump at a set flow rate. The feed rate of liquid 2 is 60.00g / min and the feed rate of hydrochloric acid solution is 15.00g / min. After the two materials are mixed by a static mixer, they enter a third tubular reactor with an inner diameter of 4mm and a total pipe length of 5m for acidification and heat preservation. The temperature is controlled at 55℃ and the residence time of the material in the reactor is 0.9min. The pH value of the liquid 2 is detected to be 3. The system runs continuously for 59.5min, and liquid 3, totaling 4457.74g, is collected. S4. The liquid 3 (4457.74g) obtained in step S3 is fed into a vertical gravity separator with an effective volume of 2L. 1L of stable bottom material composed of oil phase and water phase is placed in the separator beforehand. The feed port is set near the oil-water separation interface (the allowable interface fluctuation range is ±2cm). Liquid 3 is cut in through the feed port. The oil phase overflows from the top of the separator and the water phase is discharged from the bottom. The separation temperature is controlled at 55℃. The feed rate of liquid 3 is 90.00g / min, the discharge rate of oil phase (liquid 4) is 27.70g / min, and the discharge rate of water phase is 62.31g / min. The system runs continuously for 49.5min. A total of 1371.45g of oil phase liquid 4 and 3086.29g of water phase are collected. S5, the oil phase liquid 4 (1371.45g) obtained in step S4 is countercurrently washed with 685.72g of deionized water. The washing equipment is a packed tower with an inner diameter of 2cm and a height of 50cm, filled with θ-ring packing. The washing temperature is controlled at 55℃. The countercurrent method is adopted. The oil phase liquid 4 enters from the bottom of the tower (feed rate 30.00g / min), and the deionized water enters from the top of the tower (feed rate 15.00g / min). The residence time of the material in the tower is 1.7min. The system runs continuously for 45.7min, and the washed oil phase liquid 5, totaling 1371.32g, is collected. S6. The oil phase liquid 5 (1371.32g) obtained in step S5 and 2194.33g of 10% NaOH aqueous solution are transported by a metering pump. The feed rate of the oil phase liquid 5 is 15.00g / min, and the feed rate of the liquid alkali is 24.00g / min. After the two materials are mixed by a static mixer, they enter a fourth tubular reactor with an inner diameter of 2mm and a total pipe length of 10m for neutralization reaction. The neutralization temperature is controlled at 60℃, and the residence time of the materials in the reactor is 0.9min. After the reaction is completed, a sodium lauroyl sarcosinate aqueous solution is obtained with a product concentration of 41.31%. After adding water to make a 30% aqueous solution (4910.42g), the product yield is 99.28% based on lauroyl chloride, the free fatty acid content is 0.58%, and the colorimetric value is 10 by the platinum-cobalt colorimetric method (the colorimetric value of commercially available similar products is usually 20).

[0046] Example 2 This embodiment provides a continuous production process for sodium cocoyl glycinate, including the following steps: S1, 412.35g of cocoyl chloride and 1463.85g of 28% sodium glycinate aqueous solution were pumped into a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 5m for preheating via a metering pump at a set flow rate. The feed rate of cocoyl chloride was 20.00g / min, and the feed rate of sodium glycinate aqueous solution was 71.00g / min. The preheating temperature was set at 60℃. After the two preheated materials were mixed by a static mixer, they entered the first tubular reactor with an inner diameter of 2mm and a tube length of 10m for reaction. The reaction temperature was controlled at 60℃, and the residence time of the materials in the reactor was 0.4min. The pH value of the output liquid was measured to be 11. The system ran continuously for 20.6min, and liquid 1, totaling 1876.20g, was collected. In step S2, 419.56g of cocoyl chloride, 1876.20g of solution 1 obtained in step S1, and 567.91g of 15% caustic soda were pumped to a tubular reaction heat exchanger with an inner diameter of 3mm and a length of 1m for preheating via a metering pump at a set flow rate. The feed rates for cocoyl chloride, solution 1, and caustic soda were 17.00g / min, 76.00g / min, and 23.00g / min, respectively. The preheating temperature was set to 60℃. After preheating, the three materials were mixed in a static mixer and then entered a second tubular reactor with an inner diameter of 3mm and a total length of 10m for reaction. The reaction temperature was controlled at 60℃, and the residence time of the materials in the reactor was 0.8min. The pH value of the solution was measured to be 11. The system ran continuously for 24.7min, and solution 2, totaling 2863.67g, was collected. S3, the 2863.67g of liquid 2 obtained in step S2 and 695.46g of 30% hydrochloric acid solution are delivered by a metering pump at a set flow rate. The feed rate of liquid 2 is 70.00g / min and the feed rate of hydrochloric acid solution is 17.00g / min. After the two materials are mixed by a static mixer, they enter a third tubular reactor with an inner diameter of 4mm and a total pipe length of 4m for acidification and heat preservation. The temperature is controlled at 70℃, and the residence time of the material in the reactor is 0.9min. The pH value of the liquid 2 is detected. The system runs continuously for 41min, and liquid 3, totaling 3559.13g, is collected. S4. The liquid 3 (3559.13g) obtained in step S3 is fed into a vertical gravity separator with an effective volume of 2L. 1L of stable bottom material composed of oil phase and water phase is placed in the separator beforehand. The feed port is set near the oil-water separation interface (the allowable interface fluctuation range is ±2cm). Liquid 3 is cut in through the feed port. The oil phase overflows from the top of the separator and the water phase is discharged from the bottom. The separation temperature is controlled at 70℃. The feed rate of liquid 3 is 54.00g / min, the discharge rate of oil phase (liquid 4) is 14.80g / min, and the discharge rate of water phase is 39.20g / min. The system runs continuously for 66min. A total of 976.80g of oil phase liquid 4 and 2582.33g of water phase are collected. S5, the oil phase liquid 4 (976.80g) obtained in step S4 is countercurrently washed with 488.40g of deionized water. The washing equipment consists of three centrifugal extractors (model: Shandong Liancui LC-50) with a drum diameter of 50mm connected in series. The washing temperature is controlled at 55℃, and the countercurrent method is adopted. The feed rate of oil phase liquid 4 is 30.00g / min, the feed rate of deionized water is 15.00g / min, and the total residence time is 6.6min. The system runs continuously for 32.6min, and the washed oil phase liquid 5, totaling 976.05g, is collected. S6. The oil phase liquid 5 (976.05g) obtained in step S5 and 1073.65g of 15% NaOH aqueous solution are fed through a metering pump. The feed rate of the oil phase liquid 5 is 20.00g / min, and the feed rate of the liquid alkali is 22.00g / min. After the two materials are mixed in a static mixer, they enter a fourth tubular reactor with an inner diameter of 2mm and a total pipe length of 7m for neutralization reaction. The neutralization temperature is controlled at 80℃, and the residence time of the materials in the reactor is 0.6min. After the reaction is completed, an aqueous solution of sodium cocoyl glycinate is obtained with a product concentration of 51.19%. After adding water to make a 30% aqueous solution (3499.335g), the product yield is 94.48% based on cocoyl chloride, the free fatty acid content is 0.61%, and the colorimetric value is 10 according to the platinum-cobalt colorimetric method (the colorimetric value of commercially available similar products is 15-20).

[0047] Example 3 This embodiment provides a continuous production process for oleoylsarcosine, including the following steps: S1, 366.04g of oleoyl chloride and 1464.16g of a 20% sodium sarcosinate aqueous solution were preheated by metering pumps at a set flow rate to a tubular reaction heat exchanger with an inner diameter of 3mm and a length of 4m. The feed rate of oleoyl chloride was 15.00g / min, and the feed rate of sodium sarcosinate aqueous solution was 60.00g / min. The preheating temperature was set at 40℃. After preheating, the two materials were mixed by a static mixer and then entered the first tubular reactor with an inner diameter of 4mm and a tube length of 10m for reaction. The reaction temperature was controlled at 40℃, and the residence time of the materials in the reactor was 2.0min. The pH value of the output liquid was measured to be 10.5. The system ran continuously for 24.5min, and liquid 1, totaling 1830.20g, was collected. S2, 386.66g of oleyl chloride, 1830.20g of the feed solution 1 obtained in step S1, and 386.67g of 14% caustic soda were pumped into a tubular reaction heat exchanger with an inner diameter of 4mm and a length of 1m for preheating via a metering pump at a set flow rate. The feed rate of oleyl chloride was 15.00g / min, the feed rate of feed solution 1 was 71.00g / min, and the feed rate of caustic soda was 15.00g / min. The preheating temperature was set to 40℃. After preheating, the three materials were mixed by a static mixer and then entered a second tubular reactor with an inner diameter of 4mm and a total pipe length of 10m for reaction. The reaction temperature was controlled at 40℃, and the residence time of the materials in the reactor was 1.5min. The pH value of the feed solution was detected to be 11. The system ran continuously for 25.8min, and feed solution 2, totaling 2603.52g, was collected. S3, the 2603.52g of liquid 2 obtained in step S2 and 411.08g of 32% hydrochloric acid solution are delivered by a metering pump at a set flow rate. The feed rate of liquid 2 is 95.00g / min and the feed rate of hydrochloric acid solution is 15.00g / min. After the two materials are mixed by a static mixer, they enter a third tubular reactor with an inner diameter of 3mm and a total pipe length of 10m for acidification and heat preservation. The temperature is controlled at 40℃ and the residence time of the material in the reactor is 0.7min. The pH value of the liquid 2 is detected. The system runs continuously for 27.4min, and liquid 3, totaling 3014.60g, is collected. S4. The liquid 3 (3014.60g) obtained in step S3 is fed into a vertical gravity separator with an effective volume of 2L. 1L of stable bottom material composed of oil phase and water phase is placed in the separator beforehand. The feed port is set near the oil-water separation interface (the allowable interface fluctuation range is ±2cm). Liquid 3 is cut in through the feed port. The oil phase overflows from the top of the separator and the water phase is discharged from the bottom. The separation temperature is controlled at 50℃. The feed rate of liquid 3 is 80.00g / min, the discharge rate of oil phase (liquid 4) is 23.45g / min, and the discharge rate of water phase is 56.55g / min. The system runs continuously for 37.7min. A total of 883.65g of oil phase liquid 4 and 2130.95g of water phase are collected. S5, the oil phase liquid 4 (883.65g) obtained in step S4 is washed countercurrently with 589.10g of deionized water. The washing equipment consists of 5 centrifugal extractors (model: Shandong Liancui LC-50) with a drum diameter of 50mm connected in series. The washing temperature is controlled at 40℃. The countercurrent method is adopted. The feed rate of oil phase liquid 4 is 30.00g / min, the feed rate of deionized water is 20.00g / min, and the total residence time is 9.8min. The system runs continuously for 29.5min, and 883.64g of oleylsarcosine product is collected. The yield based on oleyl chloride is 99.94%, the free fatty acid impurity content is 0.89%, and the Gardner color is 1-2 (the color of commercially available similar products is 5).

[0048] Example 4 This embodiment provides a continuous production process for sodium cocoyl alanine, comprising the following steps: S1, 488.01g of cocoyl chloride and 1952.02g of a 26% sodium alanine aqueous solution were pumped into a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 1m for preheating via a metering pump at a set flow rate. The feed rate of cocoyl chloride was 18.00g / min, and the feed rate of sodium alanine aqueous solution was 72.00g / min. The preheating temperature was set at 50℃. After the two preheated materials were mixed by a static mixer, they entered the first tubular reactor with an inner diameter of 2mm and a tube length of 8m for reaction. The reaction temperature was controlled at 50℃, and the residence time of the materials in the reactor was 0.3min. The pH value of the output liquid was measured to be 9. The system ran continuously for 27.1min, and liquid 1, totaling 2440.03g, was collected. S2, 488.00g of cocoyl chloride, 2440.03g of liquid solution 1 obtained in step S1, and 488.00g of 18% caustic soda were pumped to a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 1m for preheating via a metering pump at a set flow rate. The feed rate of cocoyl chloride was 18.00g / min, the feed rate of liquid solution 1 was 90.00g / min, and the feed rate of caustic soda was 18.00g / min. The preheating temperature was set to 50℃. After preheating, the three materials were mixed by a static mixer and then entered a second tubular reactor with an inner diameter of 4mm and a total pipe length of 15m for reaction. The reaction temperature was controlled at 50℃, and the residence time of the materials in the reactor was 1.8min. The pH value of the liquid solution was detected to be 9. The system ran continuously for 27.1min, and liquid solution 2, totaling 3416.02g, was collected. S3, the 3416.02g of liquid 2 obtained in step S2 and 788.31g of 32% sulfuric acid solution are delivered by a metering pump at a set flow rate. The feed rate of liquid 2 is 65.00g / min, and the feed rate of sulfuric acid solution is 15.00g / min. After the two materials are mixed by a static mixer, they enter a third tubular reactor with an inner diameter of 2mm and a total pipe length of 8m for acidification and heat preservation. The temperature is controlled at 55℃, and the residence time of the material in the reactor is 0.4min. The pH value of the liquid 2 is detected. The system runs continuously for 52.6min, and liquid 3, totaling 4204.33g, is collected. S4. The liquid 3 (4204.33g) obtained in step S3 is fed into a vertical gravity separator with an effective volume of 2L. 1L of stable bottom material composed of oil phase and water phase is placed in the separator beforehand. The feed port is set near the oil-water separation interface (the allowable interface fluctuation range is ±2cm). Liquid 3 is cut in through the feed port. The oil phase overflows from the top of the separator and the water phase is discharged from the bottom. The separation temperature is controlled at 55℃. The feed rate of liquid 3 is 93.00g / min, the discharge rate of oil phase (liquid 4) is 26.75g / min, and the discharge rate of water phase is 66.25g / min. The system runs continuously for 45.2min. A total of 1209.10g of oil phase liquid 4 and 2995.24g of water phase are collected. S5, the oil phase liquid 4 (1209.10g) obtained in step S4 is washed countercurrently with 1209.10g of deionized water. The washing equipment consists of four centrifugal extractors (model: Shandong Liancui LC-50) with a drum diameter of 50mm connected in series. The washing temperature is controlled at 55℃, and the countercurrent method is adopted. The feed rate of oil phase liquid 4 is 20.00g / min, the feed rate of deionized water is 20.00g / min, and the total residence time is 9.8min. The system runs continuously for 60.5min, and the washed oil phase liquid 5, totaling 1209.08g, is collected. S6. The oil phase liquid 5 (1209.08g) obtained in step S5 and 1511.37g of 12% NaOH aqueous solution were transported by a metering pump. The feed rate of the oil phase liquid 5 was 16.00g / min, and the feed rate of the liquid alkali was 20.00g / min. After the two materials were mixed by a static mixer, they entered a fourth tubular reactor with an inner diameter of 4mm and a total pipe length of 5m for neutralization reaction. The neutralization temperature was controlled at 60℃, and the residence time of the materials in the reactor was 1.9min. After the reaction was completed, an aqueous solution of sodium cocoyl alanine was obtained with a product concentration of 47.74%. After adding water to make a 35% aqueous solution (3710.67g), the product yield was 98.59% based on cocoyl chloride, the free fatty acid content was 0.66%, and the colorimetric value was 10 by the platinum-cobalt colorimetric method (the colorimetric value of commercially available similar products is 20).

[0049] Example 5 This embodiment provides a continuous production process for sodium cocoyl sarcosinate, including the following steps: S1, 418.50g of cocoyl chloride and 1813.52g of 25% sodium sarcosinate aqueous solution were pumped into a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 1m for preheating via a metering pump at a set flow rate. The feed rate of cocoyl chloride was 15.00g / min, and the feed rate of sodium sarcosinate aqueous solution was 65.00g / min. The preheating temperature was set at 55℃. After the two preheated materials were mixed by a static mixer, they entered the first tubular reactor with an inner diameter of 2mm and a tube length of 8m for reaction. The reaction temperature was controlled at 55℃, and the residence time of the materials in the reactor was 0.4min. The pH value of the output liquid was measured to be 9. The system ran continuously for 28min, and liquid 1, totaling 2232.03g, was collected. S2, 465.00g of cocoyl chloride, 2232.03g of liquid solution 1 obtained in step S1, and 527.00g of 16% caustic soda were pumped into a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 1m for preheating via a metering pump at a set flow rate. The feed rate of cocoyl chloride was 15.00g / min, the feed rate of liquid solution 1 was 72.00g / min, and the feed rate of caustic soda was 17.00g / min. The preheating temperature was set at 55℃. After preheating, the three materials were mixed by a static mixer and then entered a second tubular reactor with an inner diameter of 2mm and a total pipe length of 15m for reaction. The reaction temperature was controlled at 55℃, and the residence time of the materials in the reactor was 0.6min. The pH value of the liquid solution was detected as 9. The system ran continuously for 31min, and liquid solution 2, totaling 3224.02g, was collected. S3, the 3224.02g of liquid 2 obtained in step S2 and 537.34g of 36% hydrochloric acid solution are delivered by a metering pump at a set flow rate. The feed rate of liquid 2 is 90.00g / min and the feed rate of hydrochloric acid solution is 15.00g / min. After the two materials are mixed by a static mixer, they enter a third tubular reactor with an inner diameter of 2mm and a total pipe length of 5m for acidification and heat preservation. The temperature is controlled at 65℃, and the residence time of the material in the reactor is 1.3min. The pH value of the liquid 2 is detected to be 2.5. The system runs continuously for 35.8min, and liquid 3, totaling 3761.38g, is collected. S4. The liquid 3 (3761.38g) obtained in step S3 is fed into a vertical gravity separator with an effective volume of 2L. 1L of stable bottom material composed of oil phase and water phase is placed in the separator beforehand. The feed port is set near the oil-water separation interface (the allowable interface fluctuation range is ±2cm). Liquid 3 is cut in through the feed port. The oil phase overflows from the top of the separator and the water phase is discharged from the bottom. The separation temperature is controlled at 55℃. The feed rate of liquid 3 is 100.00g / min, the discharge rate of oil phase (liquid 4) is 29.11g / min, and the discharge rate of water phase is 70.89g / min. The system runs continuously for 37.6min. A total of 1094.54g of oil phase liquid 4 and 2666.82g of water phase are collected. S5, the oil phase liquid 4 (1094.54g) obtained in step S4 is countercurrently washed with 547.27g of deionized water. The washing equipment is a packed tower with an inner diameter of 2cm and a height of 100cm, filled with glass spring packing. The washing temperature is controlled at 55℃. The countercurrent method is adopted. The oil phase liquid 4 enters from the bottom of the tower (feed rate 30.00g / min), and the deionized water enters from the top of the tower (feed rate 15.00g / min). The residence time of the material in the tower is 5min. The system runs continuously for 36.5min, and the washed oil phase liquid 5, totaling 1094.50g, is collected. S6. The oil phase liquid 5 (1094.50g) obtained in step S5 and 1702.55g of 10% NaOH aqueous solution were transported by a metering pump. The feed rate of the oil phase liquid 5 was 18.00g / min, and the feed rate of the liquid alkali was 28.00g / min. After the two materials were mixed by a static mixer, they entered a fourth tubular reactor with an inner diameter of 3mm and a total pipe length of 10m for neutralization reaction. The neutralization temperature was controlled at 70℃, and the residence time of the materials in the reactor was 1.7min. After the reaction was completed, an aqueous solution of sodium cocoyl sarcosinate was obtained with a product concentration of 42.03%. After adding water to make a 30% aqueous solution (3918.85g), the product yield was 98.60% based on cocoyl chloride, the free fatty acid content was 0.55%, and the colorimetric value was 10 by the platinum-cobalt colorimetric method (the colorimetric value of commercially available similar products is 20).

[0050] Example 6 This embodiment provides a continuous production process for stearoylsarcosine, including the following steps: S1, 469.05g of stearoyl chloride and 1407.15g of 25% sodium sarcosinate aqueous solution were pumped into a tubular reaction heat exchanger with an inner diameter of 3mm and a length of 1m for preheating via a metering pump at a set flow rate. The feed rate of stearoyl chloride was 20.00g / min, and the feed rate of sodium sarcosinate aqueous solution was 60.00g / min. The preheating temperature was set at 60℃. After the two preheated materials were mixed by a static mixer, they entered the first tubular reactor with an inner diameter of 3mm and a tube length of 6m for reaction. The reaction temperature was controlled at 60℃, and the residence time of the materials in the reactor was 0.7min. The pH value of the output liquid was measured to be 11. The system ran continuously for 23.5min, and liquid 1, totaling 1876.20g, was collected. S2, 469.05g of stearoyl chloride, 1876.20g of solution 1 obtained in step S1, and 351.78g of 17% caustic soda were pumped to a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 1m for preheating via a metering pump at a set flow rate. The feed rate of stearoyl chloride was 20.00g / min, the feed rate of solution 1 was 80.00g / min, and the feed rate of caustic soda was 15.00g / min. The preheating temperature was set to 60℃. After preheating, the three materials were mixed by a static mixer and then entered a second tubular reactor with an inner diameter of 3mm and a total pipe length of 15m for reaction. The reaction temperature was controlled at 60℃, and the residence time of the materials in the reactor was 1.2min. The pH value of the solution was detected to be 11. The system ran continuously for 23.5min, and solution 2, totaling 2697.03g, was collected. S3, the 2697.03g of liquid 2 obtained in step S2 and 539.41g of 30% hydrochloric acid solution are delivered by a metering pump at a set flow rate. The feed rate of liquid 2 is 80.00g / min and the feed rate of hydrochloric acid solution is 16.00g / min. After the two materials are mixed by a static mixer, they enter a third tubular reactor with an inner diameter of 3mm and a total pipe length of 10m for acidification and heat preservation. The temperature is controlled at 60℃ and the residence time of the material in the reactor is 0.8min. The pH value of the liquid 2 is detected. The system runs continuously for 33.7min, and liquid 3, totaling 3236.44g, is collected. S4. The liquid 3 (3236.44g) obtained in step S3 is fed into a vertical gravity separator with an effective volume of 2L. 1L of stable bottom material composed of oil phase and water phase is placed in the separator beforehand. The feed port is set near the oil-water separation interface (the allowable interface fluctuation range is ±2cm). Liquid 3 is cut in through the feed port. The oil phase overflows from the top of the separator and the water phase is discharged from the bottom. The separation temperature is controlled at 60℃. The feed rate of liquid 3 is 51.00g / min, the discharge rate of oil phase (liquid 4) is 17.35g / min, and the discharge rate of water phase is 33.65g / min. The system runs continuously for 63.5min. A total of 1101.72g of oil phase liquid 4 and 2134.72g of water phase are collected. S5, the oil phase liquid 4 (1101.72g) obtained in step S4 is countercurrently washed with 734.48g of deionized water. The washing equipment is a packed tower with an inner diameter of 2cm and a height of 100cm, filled with glass bead packing. The washing temperature is controlled at 70℃. The countercurrent method is adopted. The oil phase liquid 4 enters from the bottom of the tower (feed rate 30.00g / min), and the deionized water enters from the top of the tower (feed rate 20.00g / min). The residence time of the material in the tower is 4min. The system runs continuously for 36.7min, and 1101.02g of stearoylsarcosine product is collected, with a yield of 99.95%. The free fatty acid impurity content is 0.59%, and the Gardner color is 1 (the color of commercially available similar products is 4-5).

[0051] Example 7 This embodiment provides a continuous production process for potassium lauroyl sarcosinate, including the following steps: S1, 554.55g of lauroyl chloride and 2459.32g of sodium sarcosinate aqueous solution with a mass concentration of 23% are fed into a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 1m via a metering pump for preheating. The feed rate of lauroyl chloride is 16.00g / min, and the feed rate of sodium sarcosinate aqueous solution is 72.00g / min. The preheating temperature is set at 55℃. After the two preheated materials are mixed by a static mixer, they enter a first tubular reactor with an inner diameter of 4mm and a tube length of 9m for reaction. The reaction temperature is controlled at 55℃, the residence time of the material in the reactor is 1.5min, the pH value of the discharged liquid is detected to be 9, the discharge rate is 88.00g / min, the running time is 34.2min, and the mass of liquid 1 is 3013.87g. In step S2, lauroyl chloride (547.97g), the solution 1 obtained in step S1 (3013.87g), and a 15% KOH aqueous solution (684.97g) are pumped to a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 1m for preheating via a metering pump at a set flow rate. The feed rate of lauroyl chloride is 16.00g / min, the feed rate of solution 1 is 88.00g / min, and the feed rate of KOH aqueous solution is 20.00g / min. The preheating temperature is set to 55℃. After preheating, the three materials are mixed in a static mixer and then enter a second tubular reactor with an inner diameter of 2mm and a total pipe length of 10m for reaction. The reaction temperature is controlled at 55℃, the residence time of the materials in the reactor is 0.3min, the pH value of the discharged solution is 10, the discharge rate is 124.00g / min, the running time is 34.2min, and the mass of solution 2 is 4246.82g. In step S3, the feed solution 2 (4246.82g) obtained in step S2 and a 30% hydrochloric acid solution (684.97g) are pumped together using a metering pump at a set flow rate. The feed rate of feed solution 2 is 124.00g / min, and the feed rate of hydrochloric acid solution is 20.00g / min. After the two materials are mixed in a static mixer, they enter a third tubular reactor with an inner diameter of 4mm and a total pipe length of 7m for acidification and heat preservation. The temperature is controlled at 65℃, the residence time of the materials in the reactor is 0.7min, the pH value of the discharged feed solution is detected to be 2, the discharge rate is 144.00g / min, the running time is 34.2min, and the mass of feed solution 3 is 4931.787g. S4. The liquid 3 (4931.787g) obtained in step S3 is fed into a vertical gravity separator with an effective volume of 2L. A 1L stable substrate composed of oil and water phases is pre-placed in the separator. The inlet is located near the oil-water separation interface (allowable interface fluctuation range of ±2cm). Liquid 3 is fed in through the inlet, the oil phase overflows from the top of the separator, and the water phase is discharged from the bottom. The separation temperature is controlled at 55℃. The feed rate of liquid 3 is 144.00g / min, the discharge rate of the oil phase (liquid 4) is 40.00g / min, and the discharge rate of the water phase is 104.00g / min. The running time is 34.2min, and the mass of the oil phase liquid 4 is 1370.4g. S5, the oil phase liquid 4 (1370.4g) obtained in step S4 is countercurrently washed with an equal amount of deionized water (1370.4g). The washing equipment is a packed tower with an inner diameter of 2cm and a height of 100cm, filled with glass spring packing. The washing temperature is controlled at 55℃, and a countercurrent method is adopted. The oil phase liquid 4 enters from the bottom of the tower (feed rate 40.00g / min), and the deionized water enters from the top of the tower (feed rate 40.00g / min). The residence time of the material in the tower is 0.8min, the discharge rate is 40.00g / min, the running time is 34.2min, and the mass of oil phase liquid 5 (1370.4g) is... S6, the oil phase feed solution 5 (1370.4g) obtained in step S5 and a 10% KOH aqueous solution (2876.88g) are delivered via a metering pump. The molar ratio of KOH in the alkali solution to N-fatty acyl amino acids in feed solution 5 is controlled at 1.02:1. After the two materials are mixed in a static mixer, they enter a fourth tubular reactor with an inner diameter of 4mm and a total pipe length of 7m for neutralization reaction. The neutralization temperature is controlled at 70℃, and the feed rate of oil phase feed solution 5 is 40.00g / kg / min. The feed rate of KOH was 84.00 g / min, and the residence time of the material in the reactor was 0.8 min. After the reaction was completed, an aqueous solution of potassium lauroyl sarcosinate was obtained with a product concentration of 36.42%. After adding 909.07 g of water to make a 30% aqueous solution (5156.35 g), the product yield was 99.96%, the free fatty acid content was 0.49%, and the colorimetric value was 10 by the platinum-cobalt colorimetric method (the colorimetric value of commercially available similar products is 20).

[0052] Comparative Example 1 This comparative example provides a preparation process for sodium cocoyl sarcosinate, comprising the following steps: S1, 432.15g of 20% caustic soda solution was mixed with 1152.40g of 20% sodium sarcosinate aqueous solution to obtain a mixed caustic solution. The mixed caustic solution and 432.15g of cocoyl chloride were then pumped via a metering pump at a set flow rate to a tubular reaction heat exchanger with an inner diameter of 2mm and a length of 1m for preheating. The feed rate of cocoyl chloride was 15.00g / min, and the feed rate of the mixed caustic solution was 55.00g / min. The preheating temperature was set to 30℃. After preheating, the two materials were mixed in a static mixer and then entered a tubular reactor with an inner diameter of 2mm and a tube length of 6m for reaction. The reaction temperature was controlled at 30℃, and the residence time of the materials in the reactor was 0.3min. The pH value of the discharged solution was measured to be 9. The system ran continuously for 28.8min, and 2016.70g of solution 2 was collected. S2, the feed solution 2 (2016.70g) obtained in step S1 and 355.89g of 25% hydrochloric acid solution are delivered by a metering pump at a set flow rate. The feed rate of feed solution 2 is 85.00g / min, and the feed rate of hydrochloric acid solution is 15.00g / min. After the two materials are mixed by a static mixer, they enter a tubular reactor with an inner diameter of 2mm and a total pipe length of 10m for acidification and heat preservation. The temperature is controlled at 55℃, and the residence time of the material in the reactor is 0.4min. The pH value of the feed solution is detected to be 3. The system runs continuously for 23.7min, and feed solution 3, totaling 2372.59g, is collected. S3, the liquid 3 (2372.59g) obtained in step S2 is fed into a vertical gravity separator with an effective volume of 2L. 1L of stable bottom material composed of oil phase and water phase is placed in the separator beforehand. The feed port is set near the oil-water separation interface (the allowable interface fluctuation range is ±2cm). Liquid 3 is cut in through the feed port. The oil phase overflows from the top of the separator and the water phase is discharged from the bottom. The separation temperature is controlled at 55℃. The feed rate of liquid 3 is 66.00g / min, the discharge rate of oil phase (liquid 4) is 15.00g / min, and the discharge rate of water phase is 51.00g / min. The system runs continuously for 36.0 min, and a total of 539.22g of oil phase liquid 4 and 1833.36g of water phase are collected. S4. The oil phase liquid 4 (539.22g) obtained in step S3 is countercurrently washed with 539.22g of deionized water. The washing equipment is a packed tower with an inner diameter of 2cm and a height of 50cm, filled with θ-ring packing. The washing temperature is controlled at 55℃. The countercurrent method is adopted. The oil phase liquid 4 enters from the bottom of the tower (feed rate 30.00g / min), and the deionized water enters from the top of the tower (feed rate 30.00g / min). The residence time of the material in the tower is 1min. The system runs continuously for 18.0min, and the washed oil phase liquid 5, totaling 539.24g, is collected. In step S5, the oil phase solution 5 (539.24 g) obtained in step S4 and 899.70 g of a 10% NaOH aqueous solution were fed through a metering pump at a rate of 15.00 g / min and a rate of 25.00 g / min, respectively. After being mixed in a static mixer, the two streams entered a tubular reactor with an inner diameter of 2 mm and a total pipe length of 8 m for neutralization. The neutralization temperature was controlled at 50°C, and the residence time of the materials in the reactor was 0.7 min. After the reaction, a sodium cocoyl sarcosinate aqueous solution was obtained with a product concentration of 40.32%. After adding water to prepare a 30% aqueous solution (1932.51 g), the product yield was 99.97%, the free fatty acid content was 1.56%, and the colorimetric value was 20 (the colorimetric value of commercially available similar products is 20).

[0053] It should be noted that the 1L stable substrate composed of oil and water phases pre-placed in step S4 of Examples 1 to 7 above was pre-prepared using the same batch of feed liquid 3 to be processed in the first test of this process: First, all discharge valves of the empty vertical gravity separator were closed, and the feed valve was opened to pump feed liquid 3 into the separator at a set flow rate until the liquid level reached the requirement. Then, the feed was stopped and the mixture was allowed to stand for 10 to 30 minutes until the oil and water phases were completely separated, forming a stable stratified system with an upper layer of N-fatty acyl amino acid product oil phase and a lower layer of acidic aqueous phase containing sodium chloride / potassium chloride. This stratified system is the pre-placed substrate for subsequent continuous separation. After that, the continuous feed and separation process was started, and the interface stability could be maintained without discharging the substrate.

[0054] In summary, the N-fatty acyl amino acid compound production process provided by this invention has high production efficiency, produces high-quality N-fatty acyl amino acid compounds with free fatty acid content below 1%, low color, and significantly reduced production costs. It can significantly enhance the market competitiveness of enterprises and has broad application prospects.

[0055] The above description is only a preferred embodiment of the present invention and is not intended to limit the present invention. Any modifications, equivalent substitutions or improvements made within the spirit and principles of the present invention should be included within the protection scope of the present invention.

Claims

1. A continuous production process for N-fatty acyl amino acid compounds, characterized in that, Includes the following steps: S1, 45% to 55% of the total feed amount of fatty acyl chloride is continuously mixed with the aqueous solution of all amino acid salts and the first condensation reaction is carried out to obtain feed solution 1; S2, the remaining fatty acyl chloride, the feed solution 1 and the strong alkali solution are continuously mixed and a second condensation reaction is carried out to obtain feed solution 2; S3, the liquid 2 and the acid solution are mixed and acidified to obtain liquid 3; S4, the liquid 3 is continuously separated to obtain an oil phase liquid 4 and a salt-containing phase; S5, the oil phase liquid 4 is continuously washed to obtain N-fatty acyl amino acid compounds; In S2, the molar ratio of the remaining fatty acyl chloride to the alkali in the strong alkali solution is 1:0.9 to 1:1.

05.

2. The continuous production process for N-fatty acyl amino acid compounds as described in claim 1, characterized in that, Specifically, the steps include the following: S1, 45% to 55% of the total fatty acyl chloride feed amount is transported to a static mixer for mixing with all the amino acid salt aqueous solution, and then enters the first tubular reactor for the first condensation reaction to obtain feed liquid 1; S2, the remaining fatty acyl chloride, the feed solution 1 and the strong alkali solution are transported to a static mixer for mixing, and then enter the second tubular reactor for the second condensation reaction to obtain feed solution 2; S3, the feed liquid 2 and the acid solution are transported to a static mixer for mixing, and then enter the third tubular reactor for acidification reaction to obtain feed liquid 3; S4, the liquid 3 is fed into a continuous liquid separator to continuously separate the oil phase liquid 4 and the salt-containing phase; S5, the oil phase liquid 4 is continuously washed with water through a continuous washing device to obtain N-fatty acyl amino acid compounds.

3. The continuous production process for N-fatty acyl amino acid compounds as described in claim 1 or 2, characterized in that, The fatty acyl chloride is a C8~C18 fatty acyl chloride; and / or The amino acid salt is at least one of sarcosine, glycine, glutamate, alanine, aspartate, or arginine.

4. The continuous production process for N-fatty acyl amino acid compounds as described in claim 2, characterized in that, The molar ratio of total fatty acyl chloride to amino acid salt is 1:1 to 1:1.1; and / or In S1, the mass concentration of the amino acid salt aqueous solution is 20%~30%; and / or In S1, the temperature of the first tubular reactor is 30℃~60℃, and the residence time of the material in the first tubular reactor is 0.3min~2min.

5. The continuous production process for N-fatty acyl amino acid compounds as described in claim 2, characterized in that, In S2, the mass concentration of the strong alkali solution is 10%~30%; and / or In S2, the strong alkaline solution is an aqueous solution of sodium hydroxide or potassium hydroxide; and / or In S2, the temperature of the second tubular reactor is 30℃~60℃, and the residence time of the material in the second tubular reactor is 0.3min~2min.

6. The continuous production process for N-fatty acyl amino acid compounds as described in claim 2, characterized in that, In S3, the acid solution is hydrochloric acid or sulfuric acid.

7. The continuous production process for N-fatty acyl amino acid compounds as described in claim 2, characterized in that, In S3, the pH of the acidification reaction is 1~4, the temperature is 40℃~70℃, and the residence time of the material in the third tubular reactor is 0.3min~1.5min.

8. The continuous production process for N-fatty acyl amino acid compounds as described in claim 2, characterized in that, In S4, the temperature of the continuous liquid separation device is 40℃~70℃.

9. The continuous production process for N-fatty acyl amino acid compounds as described in claim 2, characterized in that, In S5, a countercurrent washing method is adopted, wherein the mass flow rate of the oil phase liquid 4 is 1 to 4 times the mass flow rate of the water phase; the washing temperature is 40℃ to 70℃.

10. The continuous production process of N-fatty acyl amino acid compounds as described in claim 2, characterized in that, It also includes step S6: after the oil phase liquid obtained from washing in step S5 is mixed with liquid alkali in a static mixer, it enters the fourth tubular reactor for continuous neutralization to obtain N-fatty acyl amino acid salt compounds.