EA045095B1 - METHOD FOR JOINT PRODUCTION OF DIALKYL SUCCINATE AND 1,4-BUTHANEDIOL BY TWO-STAGE HYDROGENATION OF DIALKYL MALEATE - Google Patents

Description

The present invention relates to a method for the efficient co-production of dialkyl succinate and 1,4-butanediol by catalytic hydrogenation of dialkyl maleate. The method essentially consists of the following steps:

(a) hydrogenating the dialkyl maleate stream in a first reaction step over a suitable catalyst to produce dialkyl succinate;

(b) separating the effluent from the first reaction stage in a flash vessel to produce a liquid stream rich in dialkyl succinate, and dividing said stream into three parts:

i) the first part corresponding to less than 50 mol. % of the total amount obtained, then subjected to low pressure phase separation, separation of absorbed hydrogen and distillation to obtain high purity dialkyl succinate;

ii) the second part is sent to the second reaction stage;

iii) the recirculation liquid, which is cooled by a pump in the heat exchanger, is sent back to the inlet of the first reaction stage;

c) further hydrogenating said second portion of the dialkyl succinate obtained in the first reaction step in a second reaction step using another suitable catalyst to produce primarily 1,4-butanediol together with small amounts of gamma-butyrolactone and tetrahydrofuran as by-products; the effluent is subjected to cooling, phase separation and distillation to yield 1,4-butanediol, where the first reaction step occurs at an inlet temperature of 80°C to 130°C, a pressure of 30 to 80 bar and an hourly liquid space velocity of 0 .5 to 2.0;

the second stage of the reaction takes place at an inlet temperature from 160°C to 190°C, at a pressure from 30 to 80 bar and at an hourly liquid space velocity from 0.1 to 0.3;

the hydrogen/organic feed molar ratio ranges from 15 to 30 in the first reaction stage and from 30 to 60 in the second reaction stage;

the amount of dimethylsuccinate not converted in the second stage of the reaction and separated by distillation in the form of a mixture with a certain amount of gamma-butyrolactone is returned to the second stage of the reaction;

In both reaction stages, conditions such as hydrogen/organic feed ratio, pressure and temperature are such as to maintain the reactors in a mixed vapor-liquid phase.

Field of technology

The embodiments described herein generally relate to a process and equipment for producing 1,4-butanediol along with a dialkyl succinate as a by-product by reacting an alkyl maleate feedstock with hydrogen gas in a two-step reaction over suitable catalysts.

State of the art

Over the past decades, the world has seen a sharp increase in the production of plastic materials and, accordingly, plastic waste.

Consequently, biodegradable and compostable polymers are attracting widespread attention as substitutes for traditional durable plastic polymers.

In new biodegradable plastics, polybutylene succinate (PBS) is one of the most interesting components due to its excellent properties such as compostability, mechanical properties, compatibility with other polymers and processability.

Polybutylene succinate can be used in many applications, such as agricultural films in place of low-density polyethylene and food contact applications.

The article Biodegradability of Plastics, published in the Journal of Molecular Sciences on August 29, 2009, by Y. Tokiwa et al., in addition to describing the mechanism of biodegradation of polybutylene succinate and other plastics, highlights the advantages of biodegradable plastics such as increasing soil fertility, reducing the accumulation of plastic materials in the environment and reducing the cost of waste management. In addition, with the help of enzymes and microorganisms, biodegradable plastics can be processed into useful metabolites (monomers and oligomers).

Polybutylene succinate is produced using 1,4-butanediol (BDO) and succinic acid (SA) as raw materials through the esterification-polycondensation method to form water; alternatively, PBS can also be produced from butanediol and dimethyl succinate (DMS) in a transesterification-polycondensation process to produce methanol rather than water.

Polybutylene succinate can be produced as a pure polymer or also as a copolymer, such as polybutylene succinate cobutylene terephthalate (PBST) using coesterification with terephthalic acid (TPA), or polybutylene succinate coadipate (PBSA) using adipic acid (AA). Another option is to formulate it in a mixture with another polymer, such as a mixture of polybutylene succinate (PBS) and polybutylene adipate soterephthalate (PBAT).

Unfortunately, even despite the high interest in PBS, the high cost of raw materials greatly affects the competitiveness of PBS and limits its production. In particular, prices for both succinic acid and dimethyl succinate are very high today. Most of the succinic acid on the market today is produced by fermentation, which requires very high production costs.

Dimethyl succinate, which is currently mainly produced by esterifying biosuccinic acid with methanol in a special plant, is even more expensive.

Therefore, it would be desirable to develop new or improved methods to simplify the industrial process for obtaining raw materials for PBS products and reduce the cost of their production.

1,4-Butanediol (BDO) is a versatile chemical intermediate due to its chemical stability and the presence of terminal primary hydroxyl groups.

BDO is an important raw material for the synthesis of technically important polymers such as thermoplastic urethanes and polyesters, essentially polybutylene terephthalate (PBT), polyester plasticizers, paints, coatings and adhesives.

In recent years, the largest user of BDO has been the production of tetrahydrofuran (THF), used to produce polytetramethylene ether glycol (PTMEG), a key raw material for the production of spandex fibers. Significant quantities of BDO are used to produce gamma-butyrolactone (GBL), which, in turn, is the main raw material used to produce Nmethylpyrrolidone (NMP) and other pyrrolidones.

Another important BDO derivative is polybutylene terephthalate (PBT).

There are several methods for producing BDO: the first commercial method was the Reppe method for producing BDO from acetylene and formaldehyde. Other technologies include the production of BDO from butadiene, propylene oxide, allyl alcohol, and more recently through biotransformation.

Due to the availability and low cost of raw materials, butane/maleic anhydride based processes have tended to dominate the market in the last thirty years and this trend will continue in the near future.

As described in the same applicant's patent EP 2782893 B1, butane/maleic anhydride based processes comprise two main groups.

The first group includes all methods based on the esterification of maleic anhydride with methanol, ethanol or other alcohols and the subsequent hydrogenation of the dialkyl ether, usually in the form of dimethyl maleate (DMM), to butanediol and (or), depending on pressure and temperature conditions, on a catalyst to tetrahydrofuran and gamma-butyrolactone.

The second group, on the contrary, includes methods based on the direct hydrogenation of maleic anhydride or maleic acid to BDO and (or) THF and GBL. Even though these processes seem theoretically simpler, with fewer processing steps (neither purification of maleic anhydride nor esterification is required), their industrial application has not been successful, likely due to the high cost of the catalyst and the complex materials of construction required to work with maleic acid at high temperature.

Looking again at the processes of the first group, based on the esterification of maleic anhydride, a further distinction can be made between processes based on vapor phase hydrogenation and processes based on mixed vapor-liquid phase hydrogenation.

The new method for the joint production of BDO and DMS presented in the present invention relates to methods for the hydrogenation of diesters of maleic anhydride in a mixed phase.

Regardless of the vapor or mixed phase, the reaction of dimethyl maleate with 1,4 butanediol is believed to proceed through at least two intermediates, namely

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CH—s—o—sn ₃

CH—C—O—CH ₃

CH ₂ - C-O-CH3

CH ₂ -C-O-CH3 sn ₂ -c sn ₂ -cn ₂

Dehydration reaction

Hydrogenolysis reaction

Hydrogenation reaction

Tetrahydrofuran hydrogenolysis reaction

CH ₂ -CH ₂ (AHr (gas) = -30.31 Kcal/mol) (AHr (gas) = 0.60 Kcal/mol) (AHr (gas) = -11.39 Kcal/mol)

R (gas) = 0.17 Kcal/mol)

In conclusion, dimethyl succinate, 1,4-butanediol, tetrahydrofuran and gamma-butyrolactone are reaction products that are converted to each other by the addition/removal of hydrogen and/or water. The distribution of these products is changed by adjusting the operating parameters of the hydrogenolysis reaction or the type of catalysts.

In particular, the first hydrogenation product is dimethyl succinate, and its final content in the reaction product is generally very low, below 10% or even below 2%. The higher the temperature and pressure and the lower the volumetric velocity, the lower the residual amount of DMS.

In the prior art, several inventions have been used to hydrogenate maleic anhydride esters to produce 1,4-butanediol or, in some cases, to produce dimethyl succinate.

US patent US 2110488 can be considered the first application concerning the hydrogenation of esters of aliphatic carboxylic acids to obtain aliphatic alcohols using a catalyst consisting of copper oxide and chromium oxide, in the temperature range from 250 to 300 ° C and in the pressure range from 150 to 300 atmospheres.

On the other hand, US Patent US 4,032,458 is the first application for the preparation of 1,4-butanediol, where a solution containing maleic acid is first esterified with a monohydric alcohol. The resulting dialkyl ester of maleic acid is hydrogenated in two reaction steps: in the first step, the double bond present in maleic acid is saturated, and in the second step, 1,4-butanediol and a monohydric alcohol are obtained. A copper-chromite catalyst is used at both stages of the reaction. Reaction temperatures range from 100 to 300°C, reaction pressures from 172 to 241 barg, and the catalyst is dissolved in suspension. In addition, since the catalyst of the first stage of the reaction is copper chromite, the selectivity towards DMS is low, which leads to the formation of a hydroxyl group.

WO 82/03854 is a general process for hydrogenating a carboxylic acid ester in the vapor phase using a catalyst comprising a mixture of copper oxide and zinc oxide. One of its several possible uses concerns the preparation of 1,4-butanediol from the ester of maleic, fumaric or succinic acid.

US Patent US 4,584,419 describes a method for hydrogenating a maleic anhydride ester to butanediol in the vapor phase in the presence of a copper chromite catalyst.

US Patent US 4,656,297 presents another process for producing butanediol by vapor phase hydrogenation of dimethyl succinate with a copper chromite catalyst by adding methanol to the feed ether.

US Pat. No. 4,751,334 also presents a vapor phase hydrogenation process with a Cu-Cr or Cu-Zn oxide catalyst to produce 1,4-butanediol from diethyl maleate sequentially in two or three adiabatic hydrogenolysis zones.

US patents US 5606099 and US 5872276 relate to a process for the hydrogenation of dialkyl maleate to dialkyl succinate in the liquid phase at a pressure of 50 to 400 bar, a temperature of 30°C to 160°C and a hydrogen/feed molar ratio of 20 to 40. The catalyst consists of compressed powders elements of the iron subgroup with the addition of elements of transition groups IV or V or elements of subgroup VI, having a certain compressive strength and internal surface area.

US Patents US 6,100,410, US 6,239,292, US 6,274,743, US 6,350,924, US 6,433,192 and US 6,433,193 describe various special integrations between a maleic anhydride plant and a BDO plant, where maleic anhydride absorbed in an organic solvent is first esterified and the resulting maleic anhydride The ether is driven away by the stream hydrogen and then subjected to hydrogenation in the vapor phase.

US patent US 6191322 presents a method for producing BDO by mixed-phase catalytic hydrogenation of gamma-butyrolactone or succinic anhydride esters or mixtures thereof in two catalytic layers sequentially with the introduction of cold hydrogen between reaction steps at a pressure of 75 to 90 bar and a temperature of 160°C to 250°C. The catalyst can be chosen between copper zinc oxide or copper chromite.

US Pat. No. 6,248,906 is a process in which a maleic anhydride ester is subjected to two successive hydrogenation steps, where primary hydrogenation produces essentially GBL and THF, and secondary hydrogenation converts a portion of the GBL to BDO. The first primary hydrogenation reaction zone completely hydrogenates DMM to DMS using a noble metal catalyst in the vapor phase at a moderate pressure of 3 to 40 bar and a feedstock with a high hydrogen/ether molar ratio.

US Patent US 6,989,455 relates to a two-step hydrogenation process to butanediol, specific to a feed stream containing maleic acid in aqueous solution. The product of the first stage of the reaction contains succinic acid. Due to the corrosive effect of maleic acid, the temperature of the first stage does not exceed 130°C. However, this method was not commercially used due to the poor stability of the catalysts used in an aqueous solution of maleic acid.

US Patent US 7,598,404 relates to a process designed to produce primarily tetrahydrofuran using increased heat of reaction. The feedstock is first hydrogenated in a pre-reactor zone containing a catalyst that promotes saturation of the carbon double bond, i.e. obtaining dimethyl succinate from dimethyl maleate. Thus, the heat of reaction can be used to evaporate part of the liquid feedstock. Accordingly, in the second reaction zone, where THF and (optionally) a diol and/or lactone are produced, the action of the reagent occurs strictly under conditions of complete evaporation.

Chinese patent CN 101747149 discloses a method for producing BDO by sequential two-stage hydrogenation using dialkyl ester of maleic acid, where the effluent from the first reaction stage is separated into two phases and the resulting liquid product is again brought into contact with hydrogen in a second hydrogenation section. The goal is to ensure high conversion and selectivity. However, 1,4-butanediol is the main product of both the first and second steps of the reaction.

EP 2,782,893, similar to US 7,598,404, relates to the first reaction step in which hydrogen saturates the double bond and preferentially forms dimethyl succinate, and where one of the objectives is improved utilization of the heat of reaction. Both stages of the reaction occur in a mixed vapor-liquid phase.

Patent CN 102070448 describes a process for the production of dimethyl succinate, which includes two reaction steps of esterification and hydrogenation. The method is dedicated to one product (DMS), and therefore, in combination with a special BDO installation for obtaining PBS, requires large investments.

Patent CN 102190582 relates to a method for simultaneously producing DMS at low pressure and

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BDO at medium pressure from DMM. However, two pressure levels are required and therefore the number of hydrogen compressors must be doubled accordingly.

Patent CN 104193619 describes a process for the production of dimethyl succinate by hydrogenation and purification using as feedstock an intermediate product, a high content dimethyl maleate stream and methanol obtained from the esterification unit of a butanediol plant. Instead of actually producing DMS and BDO together in one installation, the invention involves two hydrogenation plants, each of which has an independent reactor, compressor, heat exchangers and other equipment connected in parallel.

Most of the patents mentioned above and others not included in this application relate to the preparation of butanediol and/or tetrahydrofuran by hydrogenation of dialkyl maleate, primarily in the vapor phase or, in some cases, in the liquid or mixed phase. Reference to the formation of dimethyl succinate in a separate reaction step, as for example in US 4,032,458, US 6,248,906, US 7,598,404 and EP 2,782,893, is always incidental and does not refer to the preparation or co-production of dimethyl succinate. Only patents CN 102190582 and CN 104193619 represent the joint receipt of BDO and VHI. However, in both cases, two separate hydrogenation compressor units are required for the two products, which requires higher capital investment and production costs.

Several other patents, such as US 5606099, US 5872276 and CN 102070448, relate only to the preparation of dimethyl succinate by hydrogenation of dimethyl maleate.

The subject of the present invention is a method for the efficient and inexpensive co-production of 1,4-butanediol and dialkyl succinate from dialkyl maleates in a single hydrogenation unit using a less expensive mixed vapor-liquid phase process.

Brief description of the invention

The present invention relates to a method for the joint production of 1,4-butanediol and dialkyl succinate by catalytic hydrogenation of dialkyl maleates.

The method essentially consists of the following steps:

a) hydrogenating the dialkyl maleate stream in a first reaction step over a suitable catalyst to produce dialkyl succinate;

b) separating the effluent from the first reaction stage in a flash vessel to obtain a liquid stream rich in dialkyl succinate, and dividing said stream into three parts:

i) the first part corresponding to less than 50 mol. % of the total amount obtained, then subjected to phase separation at low pressure, separation of absorbed hydrogen and distillation to obtain dialkyl succinate of high purity ii) the second part is sent to the second stage of the reaction iii) the recirculation liquid, which is cooled by a pump in a heat exchanger, is sent back to the inlet first stage of reaction

c) further hydrogenating said second portion of the dialkyl succinate produced in the first reaction step in a second reaction step using another suitable catalyst to produce primarily 1,4-butanediol along with small amounts of gamma-butyrolactone and tetrahydrofuran as by-products; the effluent is subjected to cooling, phase separation and distillation to yield 1,4-butanediol, where the first reaction step occurs at an inlet temperature of 80°C to 130°C, a pressure of 30 to 80 bar and an hourly liquid space velocity of 0 .5 to 2.0;

the second stage of the reaction takes place at an inlet temperature from 160°C to 190°C, at a pressure from 30 to 80 bar and at an hourly liquid space velocity from 0.1 to 0.3;

the hydrogen/organic feed molar ratio ranges from 15 to 30 in the first reaction stage and from 30 to 60 in the second reaction stage;

the amount of dimethylsuccinate not converted in the second stage of the reaction and separated by distillation in the form of a mixture with a certain amount of gamma-butyrolactone is returned to the second stage of the reaction;

In both reaction stages, conditions such as hydrogen/organic feed ratio, pressure and temperature are such as to maintain the reactors in a mixed vapor-liquid phase.

The above-described mixed vapor-liquid phase operation represents one of the fundamental aspects of the method presented in this application. In fact, in most patents and most industrial plants for the production of BDO from maleic anhydride using dimethyl maleate as an intermediate, the hydrogenation step occurs in the vapor phase. Due to the very large amount of circulating hydrogen gas, both reactants and reaction products are in the vapor phase. The possible dimethyl succinate produced in the hypothetical first reaction step could be separated from the reaction gas only through a substantial cooling and condensation step followed by a heating step of the DMS fraction fed to the hydrogenolysis step to produce butanediol.

In conclusion, the process according to the present application, characterized by DMS separation obtained by simple phase separation, is only applicable to reactions with relatively low hydrogen circulation and mixed phase hydrogenation reactions.

In any case, said implementation in the mixed phase is not a limitation. On the contrary, as well described in EP 2782893, hydrogenation in the vapor-liquid phase compared to the vapor phase offers various advantages in terms of lower investment costs, lower energy consumption, higher reaction selectivity, better BDO product quality and less proneness of catalysts to fouling and deactivation problems .

Moreover, the advantages of the present invention are also particularly enhanced when the method is applied to a dialkyl maleate feedstock having high purity. In fact, if the very high selectivity of the first reaction step using a special catalyst designed to hydrogenate the double bond is matched to the high purity of the feedstock, the result is a product of exceptional purity. The resulting dialkyl succinate requires very simple purification or can even be used as is without any purification.

High purity dialkyl maleate can be prepared, for example, in accordance with US Pat. No. 6,392,088, in which the dialkyl ester is prepared by esterifying a dicarboxylic acid anhydride with methanol in a two-stage reaction-distillation system. The reaction product is then subjected to atmospheric pressure stripping to remove excess methanol and then vacuum distillation to separate the homogeneous catalyst and residual monoalkyl ether as a bottom product. Pure dimethyl ether is isolated as the overhead stream of a vacuum distillation column.

To maintain the high purity of the resulting dialkyl succinate, the return of unreacted DMS to hydrogenation, in contrast to the prior art, should be directed to the second stage, and not to the first. In fact, since DMS forms an azeotrope with GBL, the return stream inevitably mixes with some GBL, which, when returned to the first stage of the reaction, should degrade the quality of the resulting DMS.

The two-stage reaction system proposed in the present invention makes it possible to jointly produce DMS and BDO, a mixture of which is used as a raw material for the production of polybutylene succinate, in one industrial installation using one compressor for circulating hydrogen. In addition, the proposed method minimizes the amount of circulating hydrogen by using a lower molar ratio of hydrogen/organic feedstock in the first reaction stage, where the organic feedstock flow rate is higher. The molar ratio will be higher in the second stage, where the flow rate of the organic feedstock will be lower after selecting a portion of the product corresponding to the desired volume of DMS production, and where the reaction conditions and catalyst require more hydrogen.

Examples of suitable catalysts for use in the first stage include low palladium catalysts supported on carbon or alumina, or other noble metal, ruthenium or platinum supported on carbon, alumina or silica. The noble metal content is from 0.1 to 2.0 wt% relative to the total mass of the catalyst.

Examples of suitable catalysts for the second stage hydrogenolysis reaction include copper-containing catalysts such as copper-chromite catalysts or copper-zinc oxide catalysts.

Typically, the hydrogen used in the process is technical pure hydrogen with a low content of inert and other hydrocarbon components. Its amount is adjusted so that the stoichiometric excess is small but sufficient to ensure completion of the reaction.

Both reaction zones operate at a pressure of from 30 (3 MPa) to 80 bar (8 MPa), preferably from 50 to 75 bar. Too low a pressure increases the yield of GBL and does not allow maintaining the reaction components in the liquid phase. Pressures above 80 bar do not provide a clear performance advantage, but only increase capital and operating costs.

The reaction temperature is maintained at the inlet to the first reactor from 80°C to 130°C and at the inlet to the second reactor from 160°C to 190°C. Due to the exothermic nature of the hydrogenation reactions, the outlet temperature of both stages is higher than the inlet temperature. The return of the cooled product in the first stage and the introduction of small amounts of cold hydrogen recycle gas in the second stage allow the reaction outlet temperature to be controlled.

One of the fundamental aspects of the process according to the present invention is the use of different molar ratios of hydrogen/organic feedstock in the two reaction steps. In the first reaction stage, where the flow rate of organic feedstock is higher, the ratio is between 15 and 30. In the second stage, where reaction conditions and catalyst require more hydrogen, the molar ratio is between 30 and 60.

The contact time with the catalyst, expressed by the liquid space velocity, ranges from 0.5 to 2.0 h-1 in the first reaction stage and from 0.1 to 0.3 h-1 in the second reaction stage.

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The method according to the present invention makes it possible to modify the relative molar ratio of dimethyl succinate and butanediol obtained in a very simple way: simply by changing the flow rate of DMS extracted at the outlet of the first reaction stage. Therefore, in accordance with this application, the DMS/BDO ratio can be adjusted from 0.05 in the case of relatively low production of FBS and relatively high production of BDO used for other BDO derivatives or, in the case of pure product for sale, to about 0. 4-0.5, when all products are intended to obtain PBS. When the ultimate goal is to produce polyester copolymers such as PBST or PBSA, or mixtures of PBS with other polyesters using VDO, for example, a mixture of PBS and PBAT, this method allows the DMS/BDO molar ratio to be adjusted within the range of 0.05 to 0.40 , depending on the characteristics of the final copolymer or polymer mixture.

According to well-known information described in the technical literature and old applications, this method also makes it possible to co-produce, in addition to DMS and BDO, also various amounts of tetrahydrofuran (THF) and gamma-butyrolactone (GBL).

For example, according to the method described in EP 2782893, the THF/BDO ratio can be changed by removing moisture from the return gas. THF production is increased even by dehydration of butanediol at moderate temperatures from 120°C to 250°C, which can occur when in contact with an acid catalyst such as an aluminosilicate or a polymer resin with a sulfonic acid group.

In addition, the (THF+GBL)/BDO ratio can be increased by decreasing the pressure of the hydrogenation reaction.

Moreover, the yield of gamma-butyrolactone can also be varied from zero to about 5-15 percent by adjusting its recycle rate, from full recycle to zero recycle, from the purification section to the hydrogenation section.

Brief description of drawings

Fig. 1 is a simplified flow diagram of a hydrogenation system, as described in paragraphs (a), (b) and (c) of the Summary of the Invention, for the co-production of dialkyl succinate and 1,4-butanediol and (optionally) gamma-butyrolactone and/or tetrahydrofuran by catalytic hydrogenation of dialkyl maleates.

In fig. 2 shows the state of the art in three different configurations:

fig. 2A is a flow diagram showing the production of PBS from n-butane using maleic anhydride and two different hydrogenation units, one to produce BDO and one to produce succinic acid;

fig. 2B is a flow chart showing the production of PBS using n-butane as raw material to produce succinic acid and natural gas and/or coal to produce butanediol using the Reppe process;

fig. 2C is a flow diagram for the production of PBS using renewable feedstocks, generally defined as bio-based materials.

Fig. 3 is a flow diagram of a process according to the present invention with the co-production of butanediol and dimethyl succinate used as raw materials for the production of polybutylene succinate.

Detailed Description of the Invention

The present invention will now be described with reference to the drawings.

Fig. 1 illustrates a plant according to the present invention for co-producing DMS, BDO and optionally GBL and THF by catalytically reacting a dialkyl maleate, typically dimethyl maleate, with hydrogen in a mixed vapor-liquid phase.

This figure shows only the main unit operations relevant to the process of the present application, without including, for the sake of clarity, auxiliary equipment such as pumps or heat exchangers that are typically included in an industrial plant according to chemical engineering principles.

Dimethyl maleate, obtained by catalytic esterification of maleic anhydride, is supplied as stream 1 in the liquid phase. Stream 1 is mixed with stream 12, which is a mixture of fresh hydrogen stream 2 with return gas stream 11.

Thus, the resulting stream 3 is a mixed vapor-liquid fluid. After joining the cooled recycle liquid stream 7 at the outlet of the first hydrogenation reactor, the resulting mixed stream 4 enters the first reaction stage 21. This is an adiabatic downflow fixed bed reactor containing a specific catalyst having high selectivity for carbon double bond saturation, contained in the dialkyl maleate.

In the catalyst bed of reactor 21, dimethyl maleate reacts with dimethyl succinate with very high conversion and selectivity. This reaction is highly exothermic and therefore the temperature of effluent stream 5 at the outlet is slightly higher than the temperature of stream 4 at the inlet. In a preferred embodiment of this process, the outlet temperature of the first reactor stage corresponds to the inlet temperature of the second reaction stage and therefore no heating or cooling is required between the two reaction stages. However, for the purpose of control over all possible operating conditions of the installation, the presence of a heat exchanger, even if it does not operate under normal operating conditions, can contribute to better temperature control.

The effluent stream 5 from the first stage of the reactor at a controlled temperature is a mixture of liquid and vapor, where the liquid phase is mainly dimethyl succinate and the vapor phase is mainly hydrogen.

The effluent from the first reaction stage enters a flash tank to produce a liquid stream rich in dialkyl succinate, which is divided into three parts. The first part of stream 8, corresponding to 5 to 50 mole percent of the total production volume, is crude DMS produced in this installation. It is then subjected to low pressure phase separation to separate the absorbed hydrogen and distillation to obtain a high purity dialkyl succinate (not shown in the figure).

The second part of the liquid, that is, stream 15, is directed to the second reaction stage, for example, through a downcomer or downcomer mounted on a tubular tray above the catalyst bed of the second reaction stage.

The third part, recirculating liquid steam 7, after cooling, is sent back to the input of the first reaction stage along with stream 3.

In fig. 1, the evaporation of the flow leaving the first reactor occurs in the enlarged upper part of the second reaction stage, equipped with a tubular tray with a tubular or drain type overflow device. Thanks to the increased cross-section of the reactor, the vapor phase, mainly hydrogen, and the liquid phase, mainly dimethyl succinate, are separated. The vapor phase, ie stream 16, passes through one or more tray tubes. The liquid phase accumulates on the plate, its level is controlled by an overflow device. The first portion of the liquid (stream 8) and the third (stream 7) described above are drawn from the side nozzle, both streams are regulated by flow rate control. The liquid in the excess flow 15, together with the vapor-phase flow 16, enters the catalyst layer of the second-stage reactor 22.

An enlarged reactor top, as shown in Figure 1, is the best and most economical solution for phase separation. Of course, to obtain the same results, the method can use various technical solutions, for example, the use of an evaporation tank separated from the reactor.

Another solution is to use two reactors located one above the other, with intermediate phase separation.

Liquid stream 14, containing a small amount of unreacted DMS recycled from the purification section of the plant, usually mixed with some GBL, also enters the catalyst bed in reactor 22.

An efficient liquid distributor is used at the top of the catalyst bed in the reactor to prevent channeling effects.

Reactor 22 is also an adiabatic downflow fixed bed reactor containing a catalyst other than the first stage catalyst, and more specifically, a catalyst suitable for the hydrogenolysis of dimethyl succinate to 1,4-butanediol. Depending on pressure and temperature conditions, different amounts of gamma-butyrolactone and/or tetrahydrofuran can be obtained together with BDO.

The product stream 9 from the reactor is first cooled and then enters the separator 23, where the crude liquid product 13, containing the main BDO with a lower percentage of gammabutyrolactone, tetrahydrofuran, methanol, water, unreacted dimethyl succinate and impurities, is separated from the vapor-phase stream 10, consisting of mainly from hydrogen with smaller amounts of methanol, tetrahydrofuran and water. The gas stream 10 is then returned through the compressor 20, the pressure drop of which corresponds to the total pressure drop in the reactors, heat exchangers, separators, pipelines and control valves of the plant.

Stream 11 leaving the compressor, together with fresh hydrogen stream 2, constitutes the total reaction gas 12 used in the hydrogenation reactors.

Examples

The method according to the present application will now be described by way of examples with reference to FIGS. 2A, 2B and 2C, which represent the state of the art in three different configurations, and FIG. 3 is a block diagram of the configuration of installations in accordance with the steps and operating conditions described in the present invention.

For clarity, FIG. 2A, 2B, 2C and 3 use the same nomenclature, namely, numbers that are multiples of ten indicate production units:

plant for the production of maleic anhydride by oxidation of n-butane, plant for the production of succinic acid by hydrogenation of maleic anhydride, 30A is a plant for the production of butanediol by esterification of MA with methanol to obtain dimethyl maleate and by hydrogenolysis of DMM to BDO in accordance with the state of the art,

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30B is a plant for the production of butanediol by esterification of succinic acid with methanol to produce dimethyl succinate, as well as by hydrogenolysis of DMS to BDO, is a plant for the production of polybutylene succinate as a result of the catalytic polycondensation of butanediol with succinic acid or with dimethyl succinate, is a plant for the production of methanol from natural gas or other raw materials, is a plant for the production of formaldehyde from methanol, is a plant for the production of acetylene from natural gas or coal, is a plant for the production of butanediol using the Reppe method, is a plant for the production of succinic acid from renewable raw materials using biotechnological way,

100 is a plant for the production of butanediol by esterification of MA with methanol to obtain dimethyl maleate and by co-production of dimethyl succinate and butanediol by two-step hydrogenation in accordance with the method described in this application.

Numbers 1 to 5 represent the feedstock, fossil or renewable, used to produce PBS, namely stream 1 is n-butane used to produce maleic anhydride by gas phase catalytic oxidation, stream 2 is hydrogen used for hydrogenation/ hydrogenolysis of maleic anhydride to succinic acid and butanediol, Stream 3 is natural gas used to produce methanol, Stream 4 is natural gas or coal used to produce acetylene, Stream 5 is a renewable feedstock to produce biosuccinic acid.

Other numbers represent intermediates, final product or recycle streams, namely stream 1 is n-butane used to produce maleic anhydride by gas phase catalytic oxidation, stream 2 is hydrogen used to hydrogenate/hydrogenolyze maleic anhydride to succinic acid and butanediol, stream 3 is natural gas used to produce methanol, stream 4 is natural gas or coal used to produce acetylene, stream 5 is a renewable feedstock to produce biosuccinic acid,

Other numbers indicate intermediate products, final product or recycle streams, as follows:

stream 11 is maleic anhydride derived from n-butane, stream 12 is maleic anhydride used to produce succinic acid, stream 13 is maleic anhydride used to produce butanediol, stream 21 is succinic acid derived from maleic anhydride, Stream 31 is butanediol derived from maleic anhydride, stream 32 is butanediol derived from bio-succinic acid, stream 41 in Figures 2 and 3 is the end product polybutylene succinate, stream 42 is methanol recycled from the PBS plant to a plant for producing BDO in accordance with the method of the present invention, stream 51 is methanol produced from natural gas, stream 61 is formaldehyde produced from methanol, stream 71 is acetylene produced from coal or natural gas, stream 81 is is butanediol obtained from formaldehyde and acetylene by the Reppe process, stream 91 represents the total amount of bio-succinic acid produced from renewable resources, stream 92 represents the fraction of bio-succinic acid used as raw material for the production of PBS, stream 93 represents is the bio-succinic acid fraction used as a feedstock for the production of butanediol, stream 101 is dimethyl succinate co-produced in accordance with the process of the present invention, stream 102 is butanediol co-produced in accordance with the process of the present invention.

In fig. 2A are block diagrams showing the operating sequence of a plant for producing PBS according to the first embodiment of the prior art, where the feedstocks are n-butane (stream 1) and hydrogen (streams 2a and 2b). N-butane is supplied to unit 10, where raw maleic anhydride is produced by catalytic oxidation with air or oxygen, which is then

- 9 045095 is regenerated with an organic solvent and purified to obtain maleic anhydride in a liquid state in the form of stream 11. Stream 11 is divided into two parts. Stream 12 enters unit 20, where it is hydrogenated to succinic anhydride and hydrated to succinic acid (stream 21). The second fraction of maleic anhydride (stream 13) enters the butanediol unit 30A, where it is transformed, through esterification, hydrogenation and distillation, to high purity butanediol (stream 31). Streams 21 and 31 are the raw material for the production of polybutylene succinate (stream 41) at polycondensation unit 40.

In fig. 2B shows a second embodiment in accordance with the prior art, where the production of succinic acid (stream 21) follows the same route as in FIG. 2A, and BDO is produced on installation 80 according to the Reppe method. The Reppe method uses formaldehyde (stream 61) and acetylene (stream 71) as raw materials. In turn, formaldehyde is produced from methanol (stream 51) in unit 60, and methanol from natural gas (stream 3) in unit 50. Acetylene (stream 71) is produced from coal (stream 3) or natural gas in unit 50. Total for To finally obtain the PBS, seven different installations are connected in series or in parallel.

Fig. 2C represents a third option, a more popular and environmentally friendly method for producing PBS: corn, sugarcane molasses or other biomass (stream 5) is transformed through fermentation and several purification steps in unit 90 to produce high-purity solid succinic acid (stream 91). Part of the YAK is supplied in the form of stream 93 to the BDO 30B installation, where, through esterification, hydrogenation and purification, it is converted into a high-purity BDO stream 32. The BDO, together with the remaining part of the YAK stream 92, enters the polycondensation unit 40, where PBS is obtained by esterification and polycondensation (stream 41). The scheme seems to be very simple: PBS is produced in only three installations. Unfortunately, the complexity and low selectivity of the biofermentation method determine the very high cost of such production of succinic acid, which today limits the commercial success of this route.

Of course, the block diagrams in Figs. 2A, 2B and 2C do not represent a complete prior art scenario for obtaining PBS. For example, in the combination of FIG. 2B and FIG. 2C BDO is obtained using the Reppe method, and YAK is obtained using biofermentation to produce PBS partly from fossil and partly from renewable resources.

In another example, the biotechnology illustrated in FIG. 2C, modified because BDO can be obtained directly by fermentation of biomass without intermediate production of succinic acid. In this installation option, 90 and 30V are connected in parallel rather than in series.

Whatever the combination of biotechnological and conventional petrochemical methods, the result is a variable, but always high cost of the final PBS product.

Finally, in FIG. 3 shows a block diagram of a plant configuration including the co-production of butanediol and dimethyl succinate, in accordance with the steps and operating conditions of the present invention. As in FIG. 2A, unit 10 is a unit for producing maleic anhydride by catalytic oxidation of a 1 n-butane stream. The maleic anhydride stream 11 enters the BDO unit 100, which is described in detail in FIG. 1. Two by-products from unit 100, DMS stream 101 and BDO stream 102, together enter unit 40 where polybutylene succinate is produced (stream 41). The methanol (stream 42) produced by the transesterification reaction is returned back to the BDO 100 unit and used to esterify maleic anhydride into dimethyl maleate.

By the way, it should be noted that producing PBS using dimethyl succinate instead of succinic acid has two advantages.

First, while succinic acid is a solid, dimethyl succinate is a liquid under normal environmental conditions and is therefore easy to handle and mixes well with another reagent, butanediol, which is much simpler and more effective than the solid. substance.

The second advantage is associated with the expected approximately half smaller amount of tetrahydrofuran formed during the synthesis of PBS compared to the use of succinic acid. Tetrahydrofuran is a by-product of butanediol dehydration, so the use of DMS helps reduce the specific consumption of BDO and, in cascade, hydrogen, maleic anhydride and n-butane.

Example 1 according to the present invention.

In the installation configuration according to FIG. 3, n-butane at 100 kg/h (stream 1) with a purity of 97% is oxidized with air in installation 10 for the production of maleic anhydride to obtain 100 kg/h of purified maleic anhydride (stream 11) with a purity of 99.9%. Maleic anhydride enters the BDO synthesis unit 100 along with 6.5 kg/h of hydrogen (stream 2) and 29.1 kg/h of recycled methanol (stream 42) from the PBS unit 40. All gamma-butyrolactone obtained during hydrogenation, and also unreacted DMS is separated in the distillation section of the plant and then recycled to the second hydrogenation reactor. The products coming from the BDO unit to the PBS 40 polycondensation unit are 66.6 kg/h of DMS (stream 101) and 45.4 kg/h of BDO (stream 102). The mole fraction of DMS in the total amount of raw materials supplied to the PBS installation is about 0.47.

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The yield of polybutylene succinate (stream 41) is about 77.4 kg/h.

Thus, in this example, a flow of n-butane at a rate of 100 kg/h results in an output of 77.4 kg/h of polymer.

Of course, the material balance in this example is simplified because it does not reflect the oxygen consumed, unreacted n-butane, and the water, carbon monoxide, carbon dioxide, and organic acids produced by Unit 10, THF, and the small amount of byproducts of Unit 100, THF, produced at unit 40, and other minor losses.

Example 2 according to the present invention.

In the configuration of the method according to the present invention shown in FIG. 1:

a flow of dimethyl maleate at a rate of 2 kmol/h with a purity of 99% (stream 1) and hydrogen at a rate of 6.4 kmol/h with a purity of 99.99% (stream 2) is added to hydrogen-enriched recirculating stream 11 containing 45 kmol/h of hydrogen and 10 kmol/h dimethylsuccinate (stream 7) returned from the outlet of the first reactor.

The resulting combined inlet temperature of the first stage of reaction 21 (stream 4) is about 120°C. The first reactor contains one catalytic layer consisting of ruthenium supported on alumina.

The output from the first reaction stage (stream 5) is evaporated into a vapor phase separation vessel, stream 16.

The resulting liquid is divided into three parts as follows: The 7:10 kmol/h flow is cooled and returned to the inlet of the first stage.

Stream 8: 0.9 kmol/h is isolated as crude DMS product.

The remaining liquid portion, stream 15, enters the second reactor 22 along with hydrogen gas (stream 16) separated in the flash tank.

The second reactor contains a copper chromite catalyst in one layer. Inlet temperature 170°C.

The pressure in both reaction stages is 70 barg, with the pressure in the second stage being slightly lower due to the pressure difference.

The effluent stream from reactor 22 (stream 9) after cooling to approximately 40-50° C., not shown in FIG. 1, enters the separation tank 23, where the raw liquid BDO (stream 13) is separated from the return gas 10.

The crude BDO is then sent to distillation columns (not shown) to separate THF, methanol, water, and the GBL/DMS azeotrope and returned as stream 14 to the second reactor 22. Finally, the butanediol is separated from the light and heavy by-products. The hydrogen-rich gas stream 10 enters the compressor 20. The compressed gas 11 combines with fresh hydrogen 2, and the resulting stream 12, after heating (not shown), represents all the hydrogen gas entering the first stage reactor 21.

In the attached table. 1 and 2 summarize the main results of the operation, such as conversion, selectivity and operating parameters.

Table 1

First reactor - hydrogenation of DMM

Fresh raw material flow rate (FMF) 288.3 kg/h

Hydrogen flow rate

Fresh 6.4 kmol/h

Return 45 kmol/h

Total 51.4 kmol/h

Inlet/outlet temperature 120/170°C

Pressure 70 barg

Hourly liquid volumetric velocity 2.0 h-1

Molar ratio H2/DMM 25.7

DMM conversion 100%

VHI selectivity is above 99%

Unrefined DMS product

Flow rate 131.5 kg/h

Purity 98.5 wt.%

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Claims (6)

table 2 Second reactor - hydrogenolysis of DMS Fresh raw material flow rate (FMR) 162 kg/h Hydrogen flow rate 49.4 kmol/h Inlet/outlet temperature 170/180°C Pressure 69.8 barg Hourly liquid volumetric velocity 0.2 h-1 Molar ratio H2/DMS 44.9 VHI conversion 97.5% Selectivity for BDO 93% (applies to work with full return GBL) Unrefined BDO product Flow rate 170 kg/h Purity 53 wt.% (relative to total product, including methanol and water) CLAIM 1. A method for the joint production of dialkyl succinate and 1,4-butanediol from dialkyl maleates and hydrogen, including the following steps: a) hydrogenating the dialkyl maleate stream in a first reaction step over a suitable catalyst to produce dialkyl succinate; b) further hydrogenating a portion of the dialkyl succinate obtained in the first reaction step in a second reaction step using another suitable catalyst to produce mainly 1,4-butanediol along with small amounts of gamma-butyrolactone and tetrahydrofuran as by-products; the effluent is subjected to cooling, phase separation and distillation to obtain 1,4-butanediol, where the first reaction step is carried out at an inlet temperature of 80 to 130°C, a pressure of 30 to 80 bar (3 to 8 MPa) and hourly liquid volumetric velocity from 0.5 to 2.0. the second stage of the reaction is carried out at an inlet temperature of 160 to 190°C, at a pressure of 30 to 80 bar (3 to 8 MPa) and at an hourly liquid space velocity of 0.1 to 0.3. in both reaction stages, conditions such as hydrogen/organic feed ratio, pressure and temperature are such as to maintain the reactors in a mixed vapor-liquid phase, characterized in that the effluent from the first reaction stage is separated in a flash tank to produce a liquid stream rich in dialkyl succinate, and divide the specified stream into three parts: i) the first part, corresponding to less than 50 mol.% of the total production, is further subjected to low pressure phase separation to separate the absorbed hydrogen and distillation to obtain high purity dialkyl succinate; ii) the second part is sent to the second reaction stage; iii) the recirculation liquid, which is cooled by a pump in the heat exchanger, is sent back to the inlet of the first reaction stage; the molar ratio of hydrogen/organic raw material is from 15 to 30 at the first stage of the reaction and from 30 to 60 at the second stage of the reaction and dimethyl succinate, which is not converted at the second stage of the reaction and is separated in the distillation section in a mixture with a certain amount of gamma-butyrolactone, is returned to the second reaction stage. 2. The method according to claim 1, characterized in that the operation is performed continuously. 3. The method according to claim 1, characterized in that the catalyst used in the first stage of the reaction contains, relative to the total weight of the catalyst, from 0.1 to 2.0 wt.% of an active metal with a low content selected from the group consisting of palladium, ruthenium, platinum or mixtures thereof deposited on a support material containing carbon, alumina, silicon oxide. 4. The method according to claim 1, characterized in that the catalyst used for hydrogenolysis in the second stage is a copper-containing catalyst selected from a copper-chromite and a copper-zinc oxide catalyst. 5. The method according to claim 1, characterized in that the dialkyl maleate is dimethyl maleate and the dialkyl succinate is dimethyl succinate. 6. The method according to claim 5, characterized in that the dimethyl maleate has a very high purity, at least 97 wt.%. -