阅读说明：本技术 通过使用非均相催化剂进行氧化酯化来生产甲基丙烯酸甲酯的方法 (Method for producing methyl methacrylate by oxidative esterification using heterogeneous catalysts ) 是由 D·A·克拉普特切特夫 K·W·林巴贺 D·A·希克曼 J·赫伦 K·W·奥尔森 D·W· 于 2018-06-25 设计创作，主要内容包括：一种从甲基丙烯醛和甲醇制备甲基丙烯酸甲酯的方法。所述方法包含使包含甲基丙烯醛、甲醇、氧气和碱的混合物与包含载体和贵金属的非均相催化剂的催化剂床在具有至少四个区的管式反应器中接触,其中包含催化剂床的反应区与不包含催化剂床的混合区交替。(A process for preparing methyl methacrylate from methacrolein and methanol. The process comprises contacting a mixture comprising methacrolein, methanol, oxygen, and a base with a catalyst bed comprising a heterogeneous catalyst of a support and a noble metal in a tubular reactor having at least four zones, wherein reaction zones comprising the catalyst bed alternate with mixing zones not comprising the catalyst bed.)

1. A process for preparing methyl methacrylate from methacrolein and methanol, the process comprising contacting a mixture comprising methacrolein, methanol, oxygen, and a base with a catalyst bed comprising a support and a heterogeneous catalyst of a noble metal in a tubular reactor having at least four zones, wherein reaction zones comprising catalyst beds alternate with mixing zones not comprising catalyst beds.

2. The process of claim 1, wherein each reaction zone comprises a catalyst bed comprising catalyst particles having an average diameter of from 200 microns to 10 millimeters.

3. The process of claim 2, wherein the superficial velocity of liquid through the reactor is from 2 to 30 mm/s.

4. The method of claim 3, wherein each mixing zone comprises heat exchange means and at least one of: (i) a static mixing device, (ii) a jet mixing device, and (iii) at least one impeller with a tip speed of 0.1 to 10 m/s.

5. The process of claim 4 wherein the temperature of the catalyst bed is from 40 to 120 ℃.

6. The process of claim 5, wherein the pH in the catalyst bed is from 4 to 8.

7. The process of claim 6, wherein the reactor has from four to ten zones.

8. The process of claim 7, wherein the base is sodium hydroxide or sodium methoxide.

9. The process of claim 8, wherein the catalyst particles have an average diameter of from 400 microns to 7 millimeters.

10. The method of claim 9, wherein the noble metal is selected from the group consisting of gold and palladium.

Background

The present invention relates to a method for preparing methyl methacrylate from methacrolein and methanol using a heterogeneous catalyst.

Methyl methacrylate is produced by oxidative esterification reactions where it is known that a decrease in the pH of the reaction mixture is detrimental. The prior art reports that addition of base to the reactor to raise the pH is performed to extend the catalyst life. A solution to this problem is to mix the base into a portion of the reaction mixture or reactants in a separate vessel, see, e.g., U.S. publication No. 2016/0251301. However, there is a need for more efficient processes that can provide improved selectivity.

Disclosure of Invention

The invention relates to a process for preparing methyl methacrylate from methacrolein and methanol, comprising contacting a mixture comprising methacrolein, methanol, oxygen and a base with a catalyst bed comprising a support and a heterogeneous catalyst of a noble metal in a tubular reactor comprising at least four zones, wherein reaction zones comprising catalyst beds alternate with mixing zones not comprising catalyst beds.

Detailed Description

Unless otherwise indicated, all percentage compositions are weight percent (wt%) and all temperatures are in units of ° c. Unless otherwise indicated, the average is an arithmetic average. The noble metal is any one of gold, platinum, iridium, osmium, silver, palladium, rhodium, and ruthenium. More than one noble metal may be present in the catalyst, in which case the limits apply to the total amount of all noble metals. The "catalyst center" is the centroid of the catalyst particle, i.e., the average position of all points in all coordinate directions. The diameter is any linear dimension through the center of the catalyst and the average diameter is the arithmetic mean of all possible diameters. The aspect ratio is the ratio of the longest to the shortest diameter. A "zone" is a portion of the length of a tubular reactor (i.e., a reactor having a substantially circular cross-section) in which the reaction mixture flows along the length of the reactor (central axis) perpendicular to the cross-section. The length of a zone is its dimension along the central axis of the reactor. The tubular reactor operates as a continuous reactor.

Preferably, the first zone is a mixing zone into which fresh reactants are fed. Preferably, the base is also fed to the first mixing zone, either together with the reactants or separately. In a preferred embodiment of the invention, the base, the oxygen or both are fed to at least one subsequent mixing zone as well as to the first mixing zone. Preferably, a portion of the reaction mixture is recycled from the subsequent mixing zone to the first mixing zone. Preferably, the reactor has at least four, preferably at least three, preferably at least two; preferably not more than ten, preferably not more than five mixing zones. Preferably, the ratio of the average length of the reaction zone to the average length of the mixing zone is from 1000:1 to 1:5, preferably from 500:1 to 1:2, preferably from 100:1 to 1: 1. Preferably, the ratio of the average length of all zones to the reactor diameter is from 1000:1 to 1:10, preferably from 500:1 to 1:5, preferably from 100:1 to 1: 2. The length of the zones need not be the same. Preferably, the reactor is substantially vertical, with the reaction mixture and gases flowing upward.

Preferably, the reactor comprises at least one cooling zone wherein heat is removed from the reaction mixture passing through said zone. The mixing zone may also be a cooling zone. Preferably, cooling is accomplished by contacting the reaction mixture with a heat exchanger, which may comprise coils, fins, or other typical heat exchange surfaces.

Preferred bases include alkali metal hydroxides and C₁-C₄Alkoxides, preferably sodium hydroxide and potassium hydroxide and sodium methoxide or potassium methoxide or sodium ethoxide or potassium ethoxide, preferably sodium hydroxide or sodium methoxide. Preferably in the form of a solution, preferably in methanol, ethanol or water; preferably methanol or water. Preferably, the alkoxide is added to methanol or ethanol. Preferably, the concentration of base in the solution is from 50 to 1 wt%, preferably from 45 to 2 wt%, preferably from 40 to 5 wt%.

Preferably, the reaction mixture in the mixing zone is mixed using static mixing means, mechanical agitation or jet mixing. Preferably, one or more impellers are used to accomplish the mechanical agitation. Preferably, the tip speed of the impeller is 0.1 to 10 m/s; preferably 1 to 5 m/s. Preferably, the mixing zone contains heat exchange means for cooling or heating purposes.

Preferably, the superficial velocity of the liquid through the catalyst bed is from 1 to 100 mm/s; preferably at least 2mm/s, preferably at least 3mm/s, preferably at least 5 mm/s; preferably not more than 30mm/s, preferably not more than 25mm/s, preferably not more than 20 mm/s. Preferably, each reactor diameter, mixing zone has at least one impeller. Preferably, the linear tip speed of the impeller is 0.1 to 10 m/s; preferably at least 0.2m/s, preferably at least 0.5m/s, preferably at least 1m/s, preferably at least 2 m/s; preferably not more than 8m/s, preferably not more than 6 m/s. Preferably, the specific energy consumption is from 0 to 5W/kg; preferably at least 0.5W/kg, preferably at least 1.0W/kg; preferably not more than 4W/kg, preferably not more than 3.5W/kg. Preferably, the H/T of the reactor is at least 1.2, preferably at least 1.3, preferably at least 1.4; preferably not more than 5, preferably not more than 4, preferably not more than 3.

Preferably, the oxygen concentration at the reactor outlet is from 0.5 to 7.5 mol%; preferably at least 1 mol%; preferably not more than 6 mol%.

Preferably, the support is particles of an oxide material; preferably gamma-, -or theta-alumina, silica, magnesia, titania, zirconia, hafnia, vanadia, niobia, tantala, ceria, yttria, lanthana or a combination thereof. Preferably, in a bagIn the part of the noble metal-containing catalyst, the surface area of the support is greater than 10m²G, preferably greater than 30m²G, preferably greater than 50m²G, preferably greater than 100m²G, preferably greater than 120m²(ii) in terms of/g. In the portion of the catalyst containing little or no noble metal, the surface area of the support may be less than 50m²G, preferably less than 20m²/g。

Preferably, the aspect ratio of the catalyst particles is not more than 10:1, preferably not more than 5:1, preferably not more than 3:1, preferably not more than 2:1, preferably not more than 1.5:1, preferably not more than 1.1: 1. Preferred shapes for the catalyst particles include spherical, cylindrical, rectangular solid, annular, multilobal (e.g., cloverleaf cross-section), shapes with multiple pores, and "carriage wheels"; preferably spherical. Irregular shapes may also be used.

Preferably, at least 90 wt.% of the noble metal is in the outer 70% of the catalyst volume (i.e. the volume of the average catalyst particle), preferably the outer 60%, preferably the outer 50%, preferably the outer 40%, preferably the outer 35%, preferably the outer 30%, preferably the outer 25% of the catalyst volume. Preferably, the outer volume of any particle shape is calculated for a volume having a constant distance from its inner surface to its outer surface (the surface of the particle), measured along a line perpendicular to the outer surface. For example, for a spherical particle, the outer x% of the volume is the spherical shell, the outer surface is the surface of the particle, and the volume is x% of the entire sphere volume. Preferably, at least 95 wt.%, preferably at least 97 wt.%, preferably at least 99 wt.% of the noble metal is in the outer volume of the catalyst. Preferably, at least 90 wt.% (preferably at least 95 wt.%, preferably at least 97 wt.%, preferably at least 99 wt.%) of the noble metal is within a distance of no more than 30%, preferably no more than 25%, preferably no more than 20%, preferably no more than 15%, preferably no more than 10%, preferably no more than 8% of the catalyst diameter from the surface. The distance to the surface is measured along a line perpendicular to the surface.

Preferably, the noble metal is gold or palladium, preferably gold.

Preferably, the catalyst particles have an average diameter of at least 200 microns, preferably at least 300 microns, preferably at least 400 microns, preferably at least 500 microns, preferably at least 600 microns, preferably at least 700 microns, preferably at least 800 microns; preferably no more than 30mm, preferably no more than 20mm, preferably no more than 10mm, preferably no more than 5mm, preferably no more than 4mm, preferably no more than 3 mm. There was no significant difference between the average diameter of the support and the average diameter of the final catalyst particles.

Preferably, the catalyst is produced by precipitating the noble metal from an aqueous solution of a metal salt in the presence of a support. Preferred noble metal salts include tetrachloroauric acid, sodium thiosulfate, sodium aurothiomalate, gold hydroxide, palladium nitrate, palladium chloride and palladium acetate. In a preferred embodiment, the catalyst is produced by an incipient wetness technique, wherein an aqueous solution of a suitable noble metal precursor salt is added to the porous inorganic oxide, such that the pores are filled with the solution and then the water is removed by drying. The resulting material is then converted to the final catalyst by calcination, reduction or other treatment known to those skilled in the art to decompose the noble metal salt to the metal or metal oxide. Preferably, C contains at least one hydroxy or carboxylic acid substituent₂-C₁₈The thiol is present in solution. Preferably, C contains at least one hydroxy or carboxylic acid substituent₂-C₁₈The thiols have 2 to 12, preferably 2 to 8, preferably 3 to 6 carbon atoms. Preferably, the thiol compound comprises no more than 4, preferably no more than 3, preferably no more than 2 total hydroxyl and carboxylic acid groups. Preferably, the thiol compound has no more than 2, preferably no more than 1 thiol group. If the thiol compound contains a carboxylic acid substituent, it may be present in the acid form, the conjugate base form, or a mixture thereof. The thiol component may also be present in its thiol (acid) form or in its conjugate base (thiolate) form. Particularly preferred thiol compounds include thiomalic acid, 3-mercaptopropionic acid, thioglycolic acid, 2-mercaptoethanol, and 1-thioglycerol, including conjugate bases thereof.

In one embodiment of the invention, the catalyst is produced by precipitation, wherein a porous inorganic oxide is immersed in an aqueous solution containing a suitable salt of a noble metal precursor, and then the salt is allowed to interact with the surface of the inorganic oxide by adjusting the pH of the solution. The resulting treated solid is then recovered (e.g., by filtration) and then converted to the final catalyst by calcination, reduction, or other pretreatment known to those skilled in the art to decompose the noble metal salt to the metal or metal oxide.

The present invention is applicable to a process for producing Methyl Methacrylate (MMA) comprising treating methacrolein with methanol in an Oxidative Esterification Reactor (OER) containing a catalyst bed. The catalyst bed contains catalyst particles. OER further contains a liquid phase comprising methacrolein, methanol and MMA and a gas phase comprising oxygen. The liquid phase may further comprise by-products such as Methacrolein Dimethyl Acetal (MDA) and Methyl Isobutyrate (MIB). Preferably, the temperature of the liquid phase is 40 to 120 ℃; preferably at least 50 ℃, preferably at least 60 ℃; preferably not more than 110 c, preferably not more than 100 c. Preferably, the catalyst bed has a pressure of 0 to 2000psig (101kPa to 14 MPa); preferably not more than 2000kPa, preferably not more than 1500 kPa. Preferably, the pH in the catalyst bed is from 4 to 10; preferably at least 5, preferably at least 5.5; preferably not more than 9, preferably not more than 8, preferably not more than 7.5. Preferably, the catalyst bed is in a tubular continuous reactor.

OER typically produces MMA, along with methacrylic acid and unreacted methanol. Preferably, methanol and methacrolein are fed to the reactor at a molar ratio of methanol to methacrolein of 1:10 to 100:1, preferably 1:2 to 20:1, preferably 1:1 to 10: 1. Preferably, the catalyst bed further comprises an inert material. Preferred inert materials include, for example, alumina, clay, glass, silicon carbide, and quartz. Preferably, the average diameter of the inert material before and/or after the catalyst bed is equal to or greater than the average diameter of the catalyst, preferably from 1 to 30 mm; preferably at least 2 mm; preferably no more than 30mm, preferably no more than 10mm, preferably no more than 7 mm. Preferably, the reaction product is fed to a methanol recovery distillation column which provides an overhead stream rich in methanol and methacrolein; preferably, this stream is recycled back to the OER. The bottoms stream from the methanol recovery distillation column comprises MMA, MDA, methacrylic acid, salt and water. In one embodiment of the invention, the MDA is hydrolyzed in a medium comprising MMA, MDA, methacrylic acid, salt and water. MDA can be hydrolyzed in the bottoms stream from the methanol recovery distillation column; the stream comprises MMA, MDA, methacrylic acid, salt and water. In another embodiment, the MDA is hydrolyzed in an organic phase separated from the methanol recovery bottoms stream. It may be necessary to add water to the organic phase to ensure that there is sufficient water for MDA hydrolysis; these amounts can be easily determined from the composition of the organic phase. The product of the MDA hydrolysis reactor is phase separated and the organic phase is passed through one or more distillation columns to produce MMA product and light and/or heavy byproducts. In another embodiment, the hydrolysis may be performed within the distillation column itself.

Preferably, the oxygen concentration at the reactor outlet is at least 1 mol%, preferably at least 2 mol%, preferably at least 3 mol%; preferably not more than 7 mol%, preferably not more than 6.5 mol%, preferably not more than 6 mol%.

In a preferred embodiment of the invention, the pH at the reactor outlet is from 3 to 7.5; preferably at least 3.5, preferably at least 4, preferably at least 4.5, preferably at least 4.8, preferably at least 5; preferably not more than 7.3, preferably not more than 7.0, preferably not more than 6.7, preferably not more than 6.4. Preferably, no base is added to the reactor or to the liquid stream entering the reactor. Preferably, the reactor is not connected to an external mixing tank where the base is introduced. The pH in the reactor may be higher, possibly above 7 near the inlet and drop below 6 near the outlet.

A preferred embodiment of a fixed bed reactor for oxidative esterification is a trickle bed reactor containing a fixed bed of catalyst and having both gas and liquid feeds passing through the reactor in a downward direction in the trickle flow, the gas phase is a continuous liquid phase therefore, the area at the top of the reactor above the fixed bed will be filled with a gas phase mixture of nitrogen, carbon dioxide, oxygen, and volatile liquid components at their respective vapor pressures at typical operating temperatures and pressures (50-90 ℃ and 60-300psig (510-2200kPa), if the gas feed is air, this vapor mixture is within a flammable envelope.

The relevant fuel mixture, temperature and pressure needs to be understood that L OC. decreases as L OC increases with temperature and pressure, and given that methanol gives L OC lower than the other two important fuels (methacrolein and methyl methacrylate), a conservative design selects a feed oxygen to inert gas ratio that ensures a composition lower than L OC at the highest expected operating temperature and pressure.

Examples of the invention

Example # 1: multi-zone reactor

A series of operations were carried out in which 20 wt% methacrolein, 200ppm inhibitor and the remainder methanol were mixed and fed to a catalytic zone consisting of an 3/8"(9.5mm) stainless steel tubular reactor containing a short silicon nitride front-end and 10g catalyst, followed by a mixing zone consisting of a 150ml liquid volume stirred vessel with a pitched blade turbine, followed by a second catalytic zone consisting of a 3/8" stainless steel tubular reactor containing a short silicon carbide front-end and 10g catalyst. The catalyst consisted of 1.5 wt% Au on a Norpro 1mm diameter high surface area alumina spherical support. Air is fed to a first catalyst zone sufficient to have approximately 5% oxygen in the outlet gas and a gas containing 8% oxygen in nitrogen is fed to a second zone sufficient to have 4% to 5% oxygen in the outlet gas. The reactor was operated at 60 ℃ and 160psig (1200 kPa). The pH at the outlet of catalyst zone 1 was approximately 6.3. The product of the reactor was sent to a liquid-vapor separator and the vapor was sent to a condenser with liquid reflux. The results are described in the table below. In some cases a base consisting of sodium methoxide/methanol is added to the mixing zone. The mixing zone was stirred at 600RPM in some cases and not stirred in other cases. The product distribution of MMA is the% MMA in the product derived from the methacrolein reactant. The product distribution of michael adducts is the% adducts in the product derived from methacrolein reactant. The space-time yield is in mol MMA/Kg catalyst in hours.

Comparative example # 2: recirculation reactor

An operation was carried out in which 20 wt% methacrolein, 200ppm inhibitor and the remainder methanol were fed to an 3/8 "stainless steel tubular reactor containing a short silicon carbide front stage followed by 10g catalyst. The catalyst consisted of 1.5 wt% Au on a Norpro 1mm diameter high surface area alumina spherical support. A gas containing 8 mol% of oxygen in nitrogen was fed to the reactor at 300sccm, and the oxygen concentration in the exhaust gas was 4 mol% to 5 mol%. The reactor was operated at 60 ℃ and 160 psig. The product of the reactor was sent to a liquid-vapor separator and the vapor was sent to a condenser with liquid reflux. A portion of the product stream from this separator is recycled to the reactor inlet and combined with the feed entering the reactor. The results are described in the table below. The product distribution is the% MMA in the product from the methacrolein reactant. A base consisting of 0.15 wt% sodium methoxide in methanol was added to a liquid-vapor separator containing a pitched-blade impeller for mixing purposes.

Operation of

Feeding of the feedstock

Recycle of

Alkali

Effluent liquid

Product distribution

Product distribution

Conversion rate

STY

(g/hr)

(g/hr)

(g/hr)

(pH)

(MMA％)

(adduct%)

(％)

(m/Kghr)

4

20

180

20

6.8

93.7

1.1

60

2.7

Conclusion

The data obtained in the multi-zone reactor indicate that mixing in the mixing zone is an important parameter for reducing michael adduct formation and increasing selectivity to MMA, which is generally measured here by product distribution, for the addition of base to the reactor. The michael adduct formation approximately doubles when improper mixing is utilized in the mixing zone relative to more typical and appropriate mixing at 600 RPM.

Even with the addition of very dilute base, a comparison of the multizone reactor with a recirculating reactor with 90% recirculation shows that the multizone performance is similar or superior (with respect to MMA (selectivity), conversion and space-time yield) to a recirculating reactor with a good product distribution.