JPS6024774B2 - Methanol production method - Google Patents

Description

DETAILED DESCRIPTION OF THE INVENTION The present invention involves the production of methanol in a liquid phase methanol reactor by incorporating a methanol-forming catalyst into a non-trivalent liquid to produce a synthesis gas containing hydrogen and carbon monoxide. The present invention relates to a method for producing methanol which involves contacting with.

Referring to US Pat. No. 96, in the production of methanol from carbon monoxide and hydrogen, an inert organic liquid medium such as pseudocumene is saturated with carbon monoxide and hydrogen, Methanol is produced by contacting a medium with a methanol-forming catalyst, and both fixed and fluidized bed systems are described. The preferred catalyst particle size for fixed bed operations is said to be in the range of about 3200 to about 6400 microns, while particle sizes of about 200 to about 4800 microns are recommended for fluidized bed catalysts. There is. U.S. Pat. No. 4,031,123 uses a methanol-forming catalyst bed containing carbon monoxide, carbon dioxide, and hydrogen in a paraffinic or cycloparaffinic liquid, thereby suppressing the methanol concentration in the liquid during the reaction. A method for producing methanol is disclosed.

It is said that the catalyst bed may be fixed or slurried in the liquid, or may be fluidized by the liquid. Depending on the type of catalyst bed used, i.e. fixed bed, fluidized bed or slurried bed, and the flow rate of the liquid used, suitable average particle sizes are from about 190 to
In the range of approximately 6400 microns. Preferred particle sizes for fluidized bed operations are between 16 and 20 mesh, or about 850 to about 1000 microns. Sher
Win and Frank are involved in the production of methanol by three-phase reaction, Hydrocar Processing (1
(November 976), pages 122-124, describes a methanol synthesis process that uses recycled inert hydrocarbons to fluidize a heterogeneous catalyst that modulates the heat of the exothermic reaction.

They explain that catalytic activity increases with decreasing particle size over the range from 1000 to 3000 microns, but this is not directly proportional. Kolbel et al.
EmopeanSyms. Chem. React. E omitted. 3rd edition, Pergamon Prees, 0 spread “d(
(1965), page 115, report on the hydrogenation of carbon monoxide to methane, in which a Ni--Mg catalyst is suspended in a paraffinic hydrocarbon.

Thernst reaction to the hydrogenation of benzene to cyclohexane over Raney-nickel catalysts
The Petrole process uses the product cyclohexane as the circulating liquid to transport the catalyst from the bottom of the reactor and back to the top of the reactor through an external heat exchanger. DMa
CEP of n etc. 60 (1964), p. 43, and Cha et al., Oil and Gas Journal (June 10, 197). A similar reaction system is available at Ard's Advanced Chemistry.
aI Engineering Volume 7, Academy
cPress, New York. It is listed on page 71. Most commonly, in systems where the catalyst is mixed into a liquid, it is assumed that mixing is accomplished by a flywheel or rising gas bubbles while the catalyst remains trapped in the reactor. The aforementioned CSterhokoard;
G℃vin ship rao, Chemical E increase i
neering Cape Jom 1, 9 (1975), 22
9Bage; and Roy et al., Chemical Eng.
neerlngScieMe, 19 (196 years), 2
See page 16.

Traditionally, the preferred liquid phase methanol production process uses a fluidized bed catalyst in which the circulating inert liquid hydrocarbon and the syngas feedstock are introduced cocurrently from the bottom of the reactor, and the hydrocarbon liquid is introduced into the gaseous The catalyst was fluidized with the aid of feed. In such systems, the reaction castle, which expands with the volume of the fluidized catalyst bed, is physically limited, primarily governed by factors such as the size of the catalyst particles and the velocity of their liquid hydrocarbons. It must be said that the conventional method considered preferable has some undesirable aspects. For example, the overall control of the fluidized bed requires relatively uniformity of catalyst particle size. Otherwise, assuming the velocity of the liquid is sufficient to fluidize large particles, 4.
This is because small particles are carried out of the desired reaction chamber. Even if uniform, suitably sized catalyst particles are initially charged to the reactor, particle attrition is an unavoidable consequence of the constant movement of the particles in the fluidized bed, and under such circumstances The inherent difficulty in producing catalyst particles that do not cause wear in the process necessitates an undesirable distribution of particle size and the formation of catalyst fines. The contamination of catalyst particles into the gas or liquid exiting the reactor results in depletion of catalyst in the fluidized bed and also creates problems for downstream process equipment. Also, the rate of inert liquid hydrocarbon circulating through the reactor is limited by the need to maintain fluidized catalyst particles within the desired boundaries of the reaction castle. In this way, hydrocarbon liquids that act as heat absorbers for highly exothermic synthesis reactions typically have a significant temperature gradient over the length of the passageway in the reactor.
This significantly hinders control of the reaction temperature. A further consequence of the relatively low liquid velocities required by the fluidized bed process is the need to cool the hydrocarbon liquid outside the reactor, due to its poor heat transfer properties due to its low velocities. This is to make it impossible to cool the liquid in the reactor, for example, by using a cooling coil. One object of the present invention is to provide a liquid phase methanol synthesis process in which wear of catalyst particles is substantially eliminated to the extent that it is not a problem.

Yet another object is to provide a process in which a wide range of fluid velocities are available through the reactor. Yet another object of the present invention is to provide a liquid methanol reaction process in which the temperature distribution within the reactor can be largely controlled. The process of the present invention utilizes relatively fine catalyst particles mixed into an inert hydrocarbon liquid as opposed to conventional fluidized catalyst beds in the liquid phase production of methanol.

Advantages of the process of the invention include reduced catalyst costs, the ability to use high temperatures for reactor operation, improved temperature distribution in the reactor, and use of inexpensive reactors. Catalyst costs are reduced due to the combined effects of various aspects of the novel catalyst incorporation method of the present invention compared to fluidized bed methods.

Preparation of catalyst particles for fluidized bed liquid phase methanol production systems typically involves obtaining a catalyst powder and using a suitable binder to form the catalyst into uniformly sized particles that are resistant to abrasion. Requires pelletizing. In the process of the invention, the catalyst can be used in powder form and therefore does not require complicated preparation operations. As currently understood, the control mechanism in liquid-phase methanol synthesis, whether fluidized bed or catalyst-loaded, involves mass transfer of reactants through a liquid film surrounding catalyst particles. It is. Reducing the catalyst particle size results in an increase in the effective area of the catalyst, thereby reducing resistance to mass transfer. Small catalyst particle size results in increased catalyst productivity per unit weight, reducing the amount of catalyst charged to the reactor to reach a given production level. Also, the small size of the catalyst particles is less sensitive to attrition and therefore produces less catalyst debris and reduces the downstream problems caused by it in the reactor. Finally, the frequency of catalyst replacement in the reactor is reduced because there is no need to maintain a uniform size catalyst pellet in the reactor. If necessary, there is no need to consider the problem of wear and tear on the chick soot bellets as a reason for replacing the catalyst. When using the catalyst incorporation method, high temperature reactor conditions are possible and high reactivity is achieved without adversely affecting methanol yield. The increased reactor operating temperature allows the exotherm of the reaction to be recovered as high pressure steam. Improved temperature distribution within the reactor represents another potential advantage of the catalyst incorporation method.

For example, countercurrent reactor configurations not easily employed in fluidized bed processes can be used, with product gas exiting the reactor at the cooled end of the liquid feed to create a temperature profile that improves Cq conversion. This results in a 20-30 qo advantage in thermodynamically limited CO conversion. Also, high fluid velocities are possible in the process of the present invention, resulting in improved gas-liquid transfer characteristics, resulting in improved temperature distribution within the reactor. Increasing the fluid velocity also increases the reaction rate because it improves the gas-liquid material transfer properties.
The previously mentioned advantages of the catalyst incorporation process combine to allow for the construction of reactors that are less expensive and of greater height-to-diameter for a given production level. In the method of the present invention, a powdered catalyst, preferably having a size of less than about 125 microns, more preferably from about 10 to about 125 microns, is actively suspended in an inert hydrocarbon liquid; The incorporated catalyst is circulated through the reactor. Thus, in a fluidized bed catalyst system, one must carefully determine the appropriate fluid velocity through the reactor and/or size the appropriate catalyst pellets to establish the desired boundaries of the fluidized catalyst bed. is required, whereas when using the liquid-entrained catalyst system of the present invention, the liquid flow rate through the reactor can only be determined by other considerations. The ability to increase the rate of inert hydrocarbon liquids and entrained catalyst through the reactor has several advantages not available to practitioners of the preferred fluidized bed liquid phase methanol production process. Provides selectivity. Without creating a significant gradient in liquid temperature from inlet to outlet, for example, the liquid can be circulated through the reactor at a rate fast enough that the liquid temperature gradient is significantly reduced. In this way a relatively narrow optimum reactor temperature range, which is highly favorable for equilibrium conditions, is maintained and highly exothermic reactions can proceed under near-isothermal conditions. Catalyst liquid-entrained systems do not require heat exchange outside the reactor system, as the high liquid velocity through the reactor significantly increases the potential for good heat transfer between that liquid and heat exchange equipment within the reactor. First, it becomes possible to remove excess heat, for example through cooling coils within the reactor.

Also, the catalyst is preferably 5 to 4
The volumetric circulation rate of the process liquid is lower in the entrained catalyst system because it contains 0% by weight, providing added heat capacity compared to the fluidized bed system. The liquid hydrocarbon used must be capable of dissolving at least small amounts of hydrogen, carbon monoxide, and methanol; stable and substantially inert; and, importantly, reactive at the temperatures and pressures used. It must be able to maintain its liquid state within the container. Naturally, the catalyst must not dissolve or react with the liquid. In general, the vapor pressure of the liquid is 35 k9' absolute at a temperature of 250 °C (50 msec).
ia) should not be exceeded. Such a liquid is preferably an organic compound. Examples of organic compounds that can be used are aromatics, such as alkylated naphthalenes with 10 to 14 carbon atoms, alkylated biphenyls with 12 to 14 carbon atoms, and 7 to 12 carbon atoms and 1 to 14 carbon atoms. Polyalkylbenzenes with up to 5 alkyl substituents (e.g. pseudocumene, xylene, and jetylbenzene); saturated alcohols with 5 to 20 carbon atoms (e.g. cyclohexanol and n-octyl alcohol); 5 to 15 carbons saturated esters with atoms (e.g. n-amyl acetate and ethyl n-valerate); saturated paraffins (including cycloparaffins) with 6 to 30 carbon atoms (e.g. hexane, dimethylbentane, and hexadecane); and the above There are mixtures of substances, but paraffins and aromatics are preferred.

The reaction temperature ranges broadly from 100 to 50,000 ℃, preferably from 200 to 400 qO; and most preferably from 215 to 2,780 ℃.

Pressures are used absolute, from 14 k9/kg to 703 k9/kg, preferably from 35 k9' to 246 kg/kg, and most preferably from 35 k9' to 105 kg/kg. The ratio of hydrogen to carbon monoxide in the feed gas is preferably from 0.6 to 10 moles hydrogen per mole carbon monoxide. Other gases may be present in the synthesis gas, such as carbon dioxide or methane. The flow rate of the reactants ranges broadly from 0.1 to 10 k9 of combined gas per hour of catalyst k9, preferably from 0.3 k9 to 5
Up to k9. Liquid flow through the reactor must be sufficient to prevent excessive temperature rise, typically between 200 and 20
000 grams, preferably from 500 to 100 grams
up to 00 grams.

The catalyst used is any methanol-forming catalyst that is active over a specified temperature range, ie from 100 to 50 psi.

Methanol-forming catalysts are described in detail in the following references: French Patent No. 1489 Spot No. 2;
Catalyst (Tokyo) 1966 8, 279-83: USSR Patent No. 21956; No. 264355; No. 269924 and No. 271497: German Patent No. 1300917:
Khim. Prom. Ukrl969(6), 7-10
:Ko Iso GakugakuZassil969 72(11
), 2360-3; German Patent No. 2016596: 1
No. 930702: No. 2026182: No. 2026165; No. 2056612 No. 216 Mother No. 79 and 21 String 074
No. KhimIM. (Sofia) 1971, 43 (1
0), 440-3. The active elements of the meta/ru-forming catalyst used include copper, zinc, aluminum, magnesium,
Includes chromium, molybdenum, uranium, tungsten, vanadium, and rare metals. US Patent 33269
Particularly useful are low temperature methanol catalysts, such as those described in US Pat. The amount of catalyst incorporated into the inert liquid may vary as desired, but an amount of catalyst from about 5 to about 4 percent by weight in the inert liquid is usually preferred. One embodiment of a three-phase liquid-entrained catalytic reactor system is shown in FIG.

A synthesis gas feed consisting of hydrogen and carbon monoxide is preheated in heat exchanger 1 by the reactor product gas and combined with the recycle gas to the bottom of reactor 2 for distributing gas bubbles throughout the reactor. Fed through a series of standard orifices. The liquid and catalyst mixture preferably consists of about 5 to 4% by weight methanol catalyst powder in paraffin oil and is passed from an intermediate drum 12 through a heat exchanger 4 by a circulating oil pump 13 to a vapor separation castle at the top of the reactor 2. It is sent to the part just below 3. The mixture of liquid and catalyst is approximately 24
The methanol flows into the reactor at a temperature of 1 and 2, and as it moves downward through the reactor, it absorbs the heat released by the methanol reaction and its temperature rises. The countercurrently flowing syngas and product gases are gradually cooled as they rise to the top of the reactor.
Due to the countercurrent flow of gas and catalyst-entrained liquid, the thermodynamic effect is that the gas exiting the reactor is cooler than when using a fluidized bed catalyst, and this lower temperature favors the equilibrium of the methanol reaction. has advantages. Due to the lower temperature, the vapor pressure of the circulating oil is also lower than in a fluidized bed catalyst system. This reduces the load on the condensate circulation system and thus improves the overall thermal efficiency of the process. The liquid and entrained catalyst exit the bottom of the reactor 2 and flow into an intermediate tank 12 which is arranged to prevent separation of liquid and catalyst.

High pressure steam can be generated from the boiler water in a heat exchanger 4 before circulating the liquid and catalyst through the reactor. Circulation through the heat exchange group should be carefully controlled to avoid solids build-up, breakage, abrasion, etc. The reactor product gas is first cooled by preheating the synthesis gas feed V in heat exchanger 1 and then in heat exchanger 5.
It is then cooled by heating the circulating gas in the heat exchanger 6 and then the boiler feed water in the heat exchanger 6, thus undergoing final cooling in the heat exchanger 7 by air or cooling water.

All of the methanol and evaporated process liquid is condensed and separated in gas-liquid separator 8. The produced methanol stream is suitable for direct fuel use or may be sent to a distillation system (not shown) to produce chemical grade products. The unconverted gas is sent back to the reactor via a circulation compressor 9 and a heat exchanger 5. In order to reduce or eliminate entrainment of catalyst fines from the reactor by the product gas stream, the reactor is designed to include a separation castle at its apex.

For example, the velocity of product gas flow at the top of the reactor can be reduced by providing an expansion castle. The larger cross-sectional area of such a zone will reduce the flow rate of gas leaving the castle, thus reducing the possibility that catalyst particles will be carried away from the liquid by the product gas. If the reactor is cylindrical and vertically arranged, the separation castle should have an inverted truncated cone-shaped expansion zone, and the smaller diameter of the trapezoid should be the same as the diameter of the reactor, and the liquid in the reactor should be The trapezoidal trapezoid is made to face the surface of the reactor so that the generated gas is discharged from the larger trapezoidal diameter to the outside of the reactor. Tests were conducted using a commercially available distilled methanol catalyst to compare the reaction rates that a catalyst-incorporated system could reasonably expect compared to using a fluidized bed-based system.

Comparisons of these different catalyst formats were approximated by using catalysts with particle sizes that corresponded to the formats used under each relevant condition. In all cases, the catalyst was dry powdered by standard techniques and slurried at 16% by weight in mineral oil before being introduced into a broken stainless steel reactor. Heat the reactor to 225 - 23000 ℃, 50% day 2, 25% CO, 10% CO2 by volume
and 15% CH4 was introduced into the reactor below the lily pad at a pressure of 35k9/kg. Three catalyst particle size distributions: {1} 37-74 micron fraction (200 x 400 mesh) sieved from catalyst powder; ■ 149 sieved from crushed catalyst tuple
Preliminary tests were conducted on one 177 micron section (80 x 100 mesh) and two total 0 micron catalyst legs (0.2 spots x 0.2 total tablets).

The data in Table 1 was obtained by aging the catalyst, and the results are shown in FIG. In the figure,
CO conversion is plotted against weight hourly space velocity (WHSV). Table 1 *Surface area, hourly space velocity (SHSV) expressed in cc of gas in STP/hr/geocatalyst surface area, based on the total external surface area of the catalyst in the reactor.

** The catalyst taken out as 0 microns (0.238 x 0.23 & netablett) is 26% and 20 x
8% were 80 mesh, and 66% were finer than 80 mesh (177 microns).

The calculated equivalent average particle diameter is 725 microns.

As is evident from Figure 2, the highly preferred entrained catalyst flowlet size of 37-74 microns provides significantly greater CO conversion at commercially practicable flow rates than is obtainable with larger particles. There is.

In fact, substantially calculated equilibrium conversion points are achieved in low flow rate reactions using catalyst particle sizes from 37 to 74 microns. FIG. 3 is a cross-plot of the relationship between CO conversion rate and WHSV shown in FIG.

However, in FIG. 3, the correlation between the Cq conversion of the catalyst mixed method and the Cq conversion of the fluidized bed catalyst method is shown outside of the conversion parameters. At low space velocities there is only a small difference between the two reaction systems, but at higher space velocities than are commercially possible, the catalyst-loaded method achieves a CO exchange rate that is four to five times higher than that of the fluidized bed catalyst. It can be seen that only 20-25% of the amount of catalyst is required to give a production corresponding to that obtained by fluidized bed catalyst operation.

[Brief explanation of the drawing]

FIG. 1 is a diagram illustrating one embodiment of a three-phase liquid-entrained catalytic reactor system, and FIG.
3 is a diagram showing the relationship between O conversion rate and WHSV, and FIG. 3 is a diagram showing a cross plot of the relationship between CO exchange rate and WHSV shown in FIG. 2. FIG. Below ``-1- F''. White. F thousand

Claims (1)

Claims: 1. A method for producing methanol in which a synthesis gas containing hydrogen and carbon monoxide is contacted with a catalyst in a liquid phase methanol reactor in the presence of an inert liquid, wherein the catalyst has a size of about 125 microns or less. Alkylated naphthalenes with 10 to 14 carbon atoms; alkylated biphenyls with 12 to 14 carbon atoms; 7 to 12 carbon atoms and 1 to 5 alkyl substituents in particle form; polyalkylbenzenes having 5 to 20 carbon atoms; saturated esters having 5 to 15 carbon atoms; and 6 to 30
A process for the production of methanol, characterized in that it is mixed into an inert liquid consisting of at least one of the group consisting of: saturated paraffins having up to 3 carbon atoms. 2. The method of claim 1, wherein the catalyst particle size is greater than about 10 microns. 3. The method for producing methanol according to claim 1, wherein the temperature of the reactor is from 100°C to 500°C. 4 The pressure of the reactor is absolute pressure 14.1 kg/cm^2
The method for producing methanol according to claim 3, wherein the methanol production method is 703 kg/cm^2. 5. The method for producing methanol according to claim 1, characterized in that the synthesis gas and the catalyst mixed in the inert liquid are introduced countercurrently into a reactor. 6. The method for producing methanol according to claim 4, wherein the reactor temperature is 215°C to 275°C and the reactor pressure is 35 to 105 kg/cm^2 absolute. 7. The method for producing methanol according to claim 1, characterized in that the amount of catalyst mixed into the inert liquid is 5 to 40% by weight based on the total weight of the inert liquid and catalyst. (a) about 5 to about 40% by weight of methanol-forming catalyst particles, based on the total weight of the inert liquid and catalyst particles;
It is mixed into an inert liquid in an amount of up to 10 particles.
to approximately 125 microns;
(b) The synthesis gas is mixed with the catalyst in the reactor. 2
At a temperature of 15°C to 275°C and an absolute pressure of 35 to 105k
g/cm^2 of pressure, and (c) separating methanol from the catalyst, inert liquid and unreacted synthesis gas, according to claim 1. methanol production method. 9. The method for producing methanol according to claim 8, wherein the synthesis gas further contains carbon dioxide and methane. 10. The method for producing methanol according to claim 8, characterized in that the synthesis gas and the liquid containing the mixed catalyst are introduced countercurrently into the reactor. 11. Claims characterized in that the product gas coming out of the reactor is sent through an open area arranged between the upper liquid level of the liquid in the reactor and the reactor outlet of the product gas. 8. The method for producing methanol according to 8. 12. The method for producing methanol according to claim 8, wherein the reactor is provided with an internal cooling device. 13 The catalyst is incorporated into the inert liquid in the form of particles having a size of about 37 to about 74 microns and the reactor temperature is 215
Claim 1, characterized in that the contact is carried out at ~275°C and an absolute pressure of 35-105 kg/cm^2 in the reactor pressure.
The method for producing methanol described in .