METHOD

TECHNICAL FIELD

The present invention relates to a process for the production of a liquid methanol and water condensate from a gaseous C0 ₂ and H ₂ feed or from a gaseous C0 ₂ , CO and H ₂ feed, or to a process for the production of a liquid methanol condensate from a gaseous CO and H ₂ feed. The invention further relates to a reactor for the preparation of a liquid condensate product from a gaseous mixture, more particularly to a reactor suitable for use in said process for the production of a liquid methanol and water condensate or a liquid methanol condensate.

BACKGROUND OF THE INVENTION

In view of increasing production of renewable electric energy, a reliable and efficient storage method is desired, especially in view of the temporal mismatch between supply and demand of renewable electric energy. Methanol can be produced from the abundantly available component C0 ₂ and from H ₂ obtained by electrolysis of water using electricity produced from renewable energy. Methanol is therefore one of the most promising potential energy storage means, storing the electric energy into an easy transportable liquid fuel. Methanol and even mixtures of methanol and water can be reconverted into electricity in a direct methanol fuel cell. Moreover, methanol is an excellent gasoline substituent and can be converted to other fuel substituent such as dimetylether or (back) to synthesis gas.

Nowadays, methanol is produced catalytically on copper-containing catalysts from synthesis gas comprising CO, C0 ₂ and H ₂ by the following exothermic hydrogenation reactions (1) and (2):

CO + H ₂ <→ CH ₃ OH (1)

C0 ₂ + 3 H ₂ <→ CH ₃ OH + H ₂ 0 (2)

The synthesis gas is usually produced from fossil resources, especially from natural gas or coal. During methanol synthesis, the methanol forming reactions (1) and (2) proceed simultaneously with the water-gas shift equilibrium reaction (3):

C0 ₂ + H ₂ <→ CO + H ₂ 0 (3) The single-pass conversion of synthesis gas comprising CO, C0 ₂ and H ₂ , i.e. the conversion of synthesis gas without recycling unreacted gaseous components to the reaction zone, is limited by chemical equilibrium. The equilibrium methanol yield is a function of temperature, pressure and feed composition. Typical equilibrium yields for the conversion of stoichiometric amounts of CO and H ₂ are 50-80%, while typical equilibrium yields for the conversion of stoichiometric amounts of C0 ₂ and H ₂ are 20-40%. The limited equilibrium conversion requires considerable gas recycle streams, i.e. multiple-pass conversion, to realize efficient use of unconverted reactants. These recycle streams lead to large capital and operational costs.

Various concepts have been suggested in the prior art to bypass the equilibrium conversion and to limit the recycle streams of unconverted reactants. An example is in-situ methanol removal from the reaction mixture by absorbing it into a solvent such as n- dodecane, tetraethylene glycol dimethyl ether (TEGDME) or alcohols. By selective removal of methanol from the reaction mixture, the equilibrium conversion is no longer limiting and higher methanol yields may be obtained.

Krishnan et al. (Continuous operation of the Berty reactor for the solvent methanol process, Industrial & Engineering Chemistry Research, 30(7) (1991), p. 1413-1418) applied the solvent TEGDME to remove methanol from the gas phase. A Berty type reactor was used in combination with a packed Cu/ZnO/Al ₂ 0 ₃ catalyst. A gas feed consisting of H ₂ /CO/CH ₄ /C0 ₂ (70%/l 5%/5%/l 0%) was fed to a reactor operated at 220°C and 78.5 bara. A conversion of 91.7% was reached. A disadvantage of the TEGDME process is the lower efficiency factor of the catalyst because the TEGDME is present in the catalyst pores. A further disadvantage is the need to recover methanol from the TEGDME.

A further solution proposed in the prior art is condensation of methanol out of the gaseous reaction mixture in the reactor. Ben Amor et al. (Methanol synthesis in a multifunctional reactor, Chemical Engineering Science, 54 (1999), p. 1419-1423) and Haut et al. (Development and analysis of a multifunctional reactor for equilibrium reactions: benzene hydrogenation and methanol synthesis, Chemical Engineering and Processing, 43 (2004), p. 979-986) developed a reactor with an internal condenser in order to shift the equilibrium. The reactor comprised a single-pass radial catalyst bed. Around this radial catalyst bed a condenser spiral was positioned to condense the methanol. A syngas feed having a 2: 1 ratio of H ₂ :CO was fed to the inner part of the radial catalyst bed. The condenser temperature and catalyst temperature were held at 100°C and 250°C, respectively. The condensed products were separated in situ using a hydraulic seal while unconverted reactants could only reach the catalyst bed by back diffusion. A purge flow was necessary to prevent accumulation of inert gasses in the reactor. A conversion of 83% of the CO was reached.

Van Bennekom et al. (Methanol synthesis beyond chemical equilibrium, Chemical Engineering Science, 87 (2013), p. 204-208) describe a process for the preparation of methanol over a Cu/ZnO/Al ₂ 03 catalyst bed at a temperature of about 273 K and a pressure of 15-20 MPa which is sufficiently high to induce condensation of the methanol from the gaseous reaction mixture at the reaction temperature. Methanol was produced from two different syngas feeds. The first syngas feed composition contained 67-68% H ₂ , 24% CO, 3- 4% C0 ₂ and 5% CH ₄ and was converted to methanol at a temperature between 468 and 543 K and a pressure of 15-20 MPa. The second syngas feed composition contained 70% H ₂ , 5% CO, 20% C0 ₂ and 5% CH ₄ and was converted to methanol at a temperature between 484 and 543 K and a pressure of 20 MPa. A conversion of more than 99% was found for the syngas composition with 3-4% C0 ₂ . For the syngas containing 20% C0 ₂ the conversion was 92.5%.

During the reaction, the condensed product comprising methanol and water contacted the catalyst particles. A disadvantage of this process is that immersion of the catalyst particles in liquid condensate of methanol and water may impair the catalyst activity and stability and introduces an additional mass transfer resistance, affecting productivity.

It is an object of the invention to provide a simplified process for the preparation of methanol from a gaseous feed comprising H ₂ , CO and/or C0 ₂ enabling full conversion of the gaseous feed into a liquid product with high selectivity towards methanol.

SUMMARY OF THE INVENTION

The inventors have developed a process for the production of a liquid methanol condensate from a gaseous CO and H ₂ feed or for the production of a liquid methanol and water condensate from:

• a gaseous C0 ₂ and H ₂ feed; or

• a gaseous C0 ₂ , CO and H ₂ feed,

by producing this liquid condensate in a reactor having one or more reaction zones and one or more condensation zones, wherein the one or more condensation zones are operated at a lower temperature than the one or more reaction zones, characterized in that there is convective mass exchange between the one or more reaction zones and the one or more condensation zones.

It was surprisingly found that by removing a liquid methanol and water condensate or a liquid methanol condensate from the gaseous reaction products in a condensation zone positioned inside the reactor and by allowing for convective mass exchange between the one or more reaction zones and the one or more condensation zones, the gaseous feed can be fully converted into a liquid product comprising methanol with a high selectivity of the reactants towards methanol. It was further surprisingly found that this process can be operated in a continuous or semi-continuous way without the need to purge part of the gaseous mixture inside the reactor to the environment, i.e. the process can be operated in a gas-in liquid-out mode (GILO), ensuring near complete carbon efficiency with only minor losses due to reactants dissolved in the liquid methanol and water condensate or in the liquid methanol condensate. Even more surprisingly, it was found that this full conversion could be realized under natural convection driven mass exchange between the one or more reaction zones and the one or more condensation zones.

The present inventors further provide a reactor for the preparation of a liquid condensate product from a gaseous mixture. This reactor can conveniently be operated in a GILO mode and is particularly suitable for use in the process for the production of a liquid methanol and water condensate from a gaseous C0 ₂ and H ₂ feed, or a gaseous C0 ₂ , CO and H ₂ feed, and is also particularly suitable for use in the process for the production of a liquid methanol condensate from a gaseous CO and H ₂ feed.

BRIEF DESCRIPTION OF THE FIGURES

In Figure 1 is schematically shown a reactor according to the invention comprising a means for forced circulation, a single annular reaction zone surrounding a single tube and a single condensation zone.

In Figure 2 is schematically shown a reactor according to the invention comprising a single annular reaction zone surrounding a single tube and a single condensation zone, said reactor being suitable for convective mass exchange by natural convection.

In Figure 3 is schematically shown a multi-tubular reactor according to the invention comprising more than one reaction zone and more than one condensation zone, said reactor being suitable for convective mass exchange by natural convection.

In Figure 4 is schematically shown a reactor according to the invention having a single reaction zone and a single tube positioned next to each other, said reactor being suitable for convective mass exchange by natural convection. In Figure 5 is schematically shown a multi-tubular reactor according to the invention comprising more than one reaction zone and more than one condensation zone, said reactor being suitable for convective mass exchange by natural convection. Figure 6 schematically shows a top view of the multi-tubular reactor shown in Figure 5, wherein the cooling means are not shown.

DETAILED DESCRIPTION

Accordingly, in a first aspect of the invention a process is provided for the production of a liquid methanol condensate from a gaseous CO and H ₂ feed or for the production of a liquid methanol and water condensate from:

• a gaseous C0 ₂ and H ₂ feed; or

• a gaseous C0 ₂ , CO and H ₂ feed,

by producing a liquid condensate in a reactor having one or more reaction zones and one or more condensation zones, wherein the one or more condensation zones are operated at a lower temperature than the one or more reaction zones, characterized in that there is convective mass exchange between the one or more reaction zones and the one or more condensation zones.

The process is performed in a "gas-in, liquid-out" reactor, meaning that the reactants fed to the reactor are gaseous and that the product which is removed from the reactor is liquid.

The term 'liquid condensate' as used herein either refers to a liquid methanol condensate when it is produced from a gaseous CO and H ₂ feed or refers to a liquid methanol and water condensate when it is produced from

• a gaseous C0 ₂ and H ₂ feed; or

• a gaseous C0 ₂ , CO and H ₂ feed.

The liquid condensate is produced from a gaseous C0 ₂ and H ₂ feed, a gaseous C0 ₂ , CO and H ₂ feed, or a gaseous CO and H ₂ feed. Since the single pass equilibrium conversion of a stoichiometric C0 ₂ and H ₂ feed is much lower than the single pass equilibrium conversion of a stoichiometric CO and H ₂ feed, the benefits of the present invention are most pronounced if the liquid condensate is produced from a feed rich in C0 ₂ . Hence, in a preferred embodiment, the liquid condensate is produced from a gaseous C0 ₂ and H ₂ feed.

In a preferred embodiment, the process as described herein before comprises the steps of: a) feeding a gaseous mixture comprising:

• C0 ₂ and H ₂ ;or • C0 ₂ , CO and H ₂ ; or

• CO and H ₂ ;

to a reactor;

b) directing the gaseous mixture over a catalyst bed for methanol production in one or more reaction zones of said reactor to produce a gaseous mixture comprising methanol or methanol and water;

c) directing the gaseous mixture obtained in step b) to one or more condensation zones in said reactor wherein a liquid methanol and water condensate or a liquid methanol condensate is separated from the remaining gaseous mixture and collected, wherein the liquid condensate does not contact the catalyst bed;

d) recycling the remaining gaseous mixture obtained in step c), preferably together with fresh feed gas as defined in step a), to the catalyst bed in the one or more reaction zones of step b); and

e) removal of the collected liquid methanol and water condensate or the liquid methanol condensate obtained in step c) from the one or more condensation zones of the reactor. The process can be operated batch-wise, continuously or in a semi-continuous way. When the process is operated in a continuous way, the gaseous mixture of reactants is continuously fed to the reactor and the liquid condensate product is continuously removed from the reactor during operation. If the process is operated in a semi-continuous way, the gaseous mixture of reactants is continuously fed to the reactor whereas the liquid condensate product is removed from the reactor at regular time intervals during the process. When the process is operated batch-wise, removal of the liquid product from the reactor during operation is not necessary but is nevertheless possible.

In order to reach a high conversion of the reactants, the molar ratio (H ₂ -C0 ₂ )/(C0 ₂ +CO) of the gaseous mixture fed to the reactor must be close to 2. In case the molar ratio (H ₂ - C0 ₂ )/(C0 ₂ +CO) of the gaseous mixture fed to the reactor is lower than 2, such as lower than 1.95, the gaseous reactants CO and C0 ₂ will accumulate in the reactor. In case the molar ratio (H ₂ -C0 ₂ )/(C0 ₂ +CO) of the gaseous mixture fed to the reactor is higher than 2, such as higher than 2.05, the gaseous reactant H ₂ will accumulate in the reactor. In case non-stoichiometric reactants accumulate in the reactor, the process when operated in a continuous way or in a semi-continuous way may require purging of gaseous components to the environment or corrective additional dosing of H ₂ or C0 ₂ at regular time intervals to avoid a build-up of gaseous reaction products in the reactor. Preferably, the molar ratio of (H ₂ -C0 ₂ )/(C0 ₂ +CO) of the gaseous mixture fed to the reactor is between 1.75 and 2.25, more preferably between 1.85 and 2.15, even more preferably between 1.95 and 2.05, most preferably between 1.95 and 2.0.

If the gaseous mixture fed to the reactor comprises components other than CO, C0 ₂ , and H ₂ taking part in the reaction, selectivity towards methanol may be reduced resulting in the formation of compounds other than methanol. If these other compounds are not condensable in the one or more condensation zones, the process when operated in a continuous way or in a semi-continuous way may require purging of gas to the environment at regular intervals to avoid a build-up of gaseous reaction products in the reactor. The term 'condensable' as used herein in the context of 'condensable compounds' refers to a gaseous compounds that can be condensed from the gaseous mixture leaving the reaction zone at the temperature and pressure in the condensation zone. Similarly, if the gaseous mixture fed to the reactor comprises components other than CO, C0 ₂ , and H ₂ which are inert in the reaction and are furthermore not condensable in the one or more condensation zones, the process when operated in a continuous way or in a semi-continuous way may require purging of gas to the environment at regular intervals to avoid a build-up of inert gaseous components in the reactor. Hence, in an embodiment of the invention, the process when operated in a continuous way or in a semi- continuous way comprises purging part of the gaseous mixture in the reactor to the environment, preferably at regular time intervals.

Purging of gaseous components to the environment results in loss of valuable reactants and a corresponding decrease in the degree of conversion of the reactants fed to the reactor. High selectivity towards methanol and a high degree of conversion of the reactants are preferred. Hence, in a preferred embodiment, the gaseous mixture fed to the reactor comprises less than 10 mol%, more preferably less than 5 mol%, even more preferably less than 2 mol % of gaseous components other than CO, C0 ₂ , and H ₂ . Examples of components other than CO, C0 ₂ , and H ₂ that may be comprised in the gaseous mixture fed to the reactor are CH ₄ and N _2.

As explained herein before, the benefits of the present invention are most pronounced for low, single pass equilibrium conversions, hence when the liquid condensate is produced from a feed rich in C0 ₂ . Hence, in a preferred embodiment, at least 90 mol% of the gaseous mixture fed to the reactor consists of C0 ₂ and H ₂ , more preferably at least 95 mol%, even more preferably at least 98 mol%, still more preferably at least 99 mol%.

C0 ₂ is abundantly available in the atmosphere and, in a more concentrated form, in natural gas fields, in flue gases and in biologically produced gases. H ₂ can be obtained by electrolysis of water. In a preferred embodiment, the C0 ₂ is obtained by gas separation from atmospheric air or from a C0 ₂ -rich gas and the H ₂ is produced by electrolysis of water using electricity obtained from renewable energy or by photo-electrolysis of water. In this way, abundantly available reactants are used and renewable electricity, preferably an excess of renewable electricity due to temporal overproduction, can be efficiently stored in the chemical bonds of methanol.

The inventors found that the process using a gaseous feed having a molar ratio (H ₂ - C0 ₂ )/(C0 ₂ +CO) of close to 2 and comprising less than 2 mol % of gaseous components other than CO, C0 ₂ , and H ₂ can be operated in a continuous or semi-continuous way without a need to purge part of the gaseous mixture inside the reactor to the environment, i.e. the process can be operated in a gas-in liquid-out mode (GILO), ensuring near complete carbon efficiency.

The liquid condensate may comprise up to 5 wt% of other compounds than methanol and water, preferably less than 1 wt%, even more preferably less than 0.1 wt%. Examples of these other compounds are alcohols, such as ethanol, propanol and higher alcohols, small acids and esters and gaseous reactants dissolved in the condensate. In a preferred embodiment, the liquid condensate comprising methanol and water substantially consists of methanol and water. In case the gaseous mixture fed to the reactor consists of CO and H ₂ , i.e. no C0 ₂ is present, no water will be formed and the liquid condensate will be a liquid methanol condensate.

The gaseous mixture leaving the one or more condensation zones, after separating the liquid condensate in step c), is recycled to the catalyst bed for methanol production in the one or more reaction zones of step b). If the process is operated in a continuous or semi- continuous way, this gaseous mixture is preferably mixed up with fresh gaseous feed mixture as defined in step a).

The catalyst bed comprises a catalyst for synthesis of methanol from hydrogen and carbon oxide gases such as CO and C0 ₂ . The process as defined in the foregoing is not limited by the type of methanol catalyst. Examples of methanol catalysts that can be used in the present process include Cu, Cu/ZnO, Cu/ZnO/Al ₂ 0 ₃ , Cu/Ti0 ₂ , Cu/ZnO/Ti0 ₂ , Cu/Zr0 ₂ , Cu/ZnO/Al ₂ 0 ₃ /Zr0 ₂ , Cu/ZnO/Cr ₂ 0 ₃ /Al ₂ 0 ₃ /Ga ₂ 0 ₃ and Raney-Cu based catalysts.

Separating the liquid condensate from the gaseous mixture entering the one or more condensation zones can for example be accomplished by letting the methanol or methanol and water condense on a condenser coil. The liquid condensate can be collected in a receptacle that is positioned in such a way that the product can be easily removed during the process and further in such a way that the liquid condensate cannot contact the catalyst bed, since direct contact of the catalyst bed with the liquid condensate may negatively impact the catalyst activity or even deactivate the catalyst.

As explained herein before, a liquid condensate is produced in one or more condensation zones positioned inside the reactor. In order to initiate condensation, the temperature within the one or more condensation zones inside the reactor should be below the dew point temperature of the gas mixture entering the one or more condensation zones.

The temperature within the one or more condensation zones varies as a function of position inside said zones. This variation depends on the distance between the one or more reaction zones and the one or more condensation zones, on the position of the catalyst bed inside each of the one or more reaction zones and on the position of the cooling means inside the one or more condensation zones. The temperature within the one or more condensation zones as used herein is defined as the temperature at the surface of the cooling means.

The temperature of the catalyst varies throughout the catalyst bed. This variation depends on the distance between the one or more reaction zones and the one or more condensation zones, on the position of the catalyst bed inside each of the one or more reaction zones and on the position of the heating means inside the one or more reaction zones. Jcataiyst is defined as the minimum temperature of the catalyst throughout the catalyst bed.

In a preferred embodiment the one or more reaction zones are operated at a temperature ^catalyst between 175 and 300 °C, the reactor is operated at a pressure between 0.1 and 10 MPa, preferably between 1 and 6 MPa, and the temperature in the one or more condensation zones is below the dew point temperature Jdew of the gas mixture entering the one or more condensation zones.

In another preferred embodiment, the one or more reaction zones are operated at a temperature Jcataiyst between 150 and 300 °C, the reactor is operated at a pressure p equal to or lower than pressure p _maK = 4.9343· 10 ^"6 _cata i _yst ³ - 1.5642- 10 ^"3 _cata i _yst ² + 0.22454· r _cata i _yst - 11.451, wherein temperature Jcataiyst is expressed in degrees centigrade and pressure p in MPa, and wherein the temperature within the one or more condensation zones is below Jdew = + 1.1 l73-(p/p _maK ) + 0.2808), wherein temperature Jdew is expressed in degrees centigrade. This embodiment is particularly preferred when the gaseous mixture fed to the reactor comprises a mixture of carbon oxides with considerable amounts of C0 ₂ and CO, such as for example gaseous mixtures comprising CO and C0 ₂ in molar ratios between 2: 1 and 1 :2, more preferably between 1.5: 1 and 1 : 1.5.

In still another preferred embodiment, the one or more reaction zones are operated at a temperature Jcataiyst between 150 and 300 °C, the reactor is operated at a pressure p equal to or lower than _Pmax = 5.9209· 10 ^"6 _cata i _yst ³ - 2.2725· 10 ^"3 _cata i _yst ² + 0.36418 _cata i _yst - 20.472, wherein temperature Jcataiyst is expressed in degrees centigrade and pressure p in MPa, and wherein the temperature within the one or more condensation zones is below Jdew = 7cataiyst"(- 0.409 -(p/pmaxf + 1.1173-(p//? _max ) + 0.2808), wherein temperature Jdew is expressed in degrees centigrade. This embodiment is particularly preferred when the carbon oxides in the gaseous mixture fed to the reactor mainly consist of C0 ₂ , such as for example at least 80 mol% of the carbon oxides, more preferably at least 90 mol%, even more preferably at least 95 mol%.

In a further preferred embodiment, the one or more reaction zones are operated at a temperature Jcataiyst between 150 and 300 °C, the reactor is operated at a pressure p equal to or lower than _Paa3L = 5.8538· 10 ^"6 _cata i _yst ³ - 2.2271 · 10 ^"3 ·Γ _Μ , ₃ ι _γ5 , ² + 0.33028 _cata i _yst - 17.122, wherein temperature Jcataiyst is expressed in degrees centigrade and pressure p in MPa, and wherein the temperature within the one or more condensation zones is below Jdew = ?cataiyst"(- 0.409 -(p/pmaxf + 1.1173-(p//? _max ) + 0.2808), wherein temperature Jdew is expressed in degrees centigrade. This embodiment is particularly preferred when the carbon oxides in the gaseous mixture fed to the reactor mainly consist of CO, such as for example at least 80 mol% of the carbon oxides, more preferably at least 90 mol%, even more preferably at least 95 mol%.

As explained herein before, there is convective mass exchange between the one or more reaction zones and the one or more condensation zones. The term 'convective mass exchange' as used herein encompasses simultaneous convective mass transfer in two directions, i.e. from the one or more reaction zones to the one or more condensation zones and vice versa. Hence, gaseous components are circulated through the reactor. In an embodiment, this convective mass exchange is realized using a means for forced circulation such as a blower, pump or compressor.

The one or more condensation zones are operated at a lower temperature than the one or more reaction zones. The present inventors found that if the one or more condensation zones are positioned in the horizontal plane above the one or more reaction zones, that is above the one or more reaction zones but not necessarily on top of the one or more reaction zones, convective mass transfer from the one or more reaction zones to the one or more condensation zones takes place via natural convection due to density differences induced by the temperature difference between the one or more condensation zones and the one or more reaction zones. If the reactor comprises one or more well-positioned separate tubes for convective mass transfer from the one or more condensation zones to the one or more reaction zones, also natural convection in this direction takes place. Convective mass exchange by natural convection is advantageous since no means for forced circulation is needed. This results in reduced capital and operational costs. Hence, in a preferred embodiment, the convective mass exchange is realized by natural convection due to the temperature difference between the one or more condensation zones and the one or more reaction zones.

In a second aspect of the invention, a reactor 1 for the preparation of a liquid condensate product from a gaseous mixture is provided, said reactor 1 comprising:

a) a housing 2;

b) an inlet 3 for feeding a gaseous mixture to the housing 2;

c) one or more reaction zones 4 inside the housing 2 covering only part of the cross-sectional area of the plane 5 perpendicular to the vertical axis 6 of the housing 2 and one or more tubes 7 covering at least part of the cross-sectional area of the plane 5 not covered by the one or more reaction zones 4, wherein the one or more reaction zones 4 and the one or more tubes 7 are separated by walls 8 parallel to the vertical axis 6 of the housing 2, wherein the walls 8 do not reach to the top 9 of the housing 2 and do not reach to the bottom 10 of the housing 2;

d) one or more condensation zones 11 inside the housing 2 which are connected to the one or more reaction zones 4 and to the one or more tubes 7 for convective mass exchange between the one or more reaction zones 4 and the one or more tubes 7 via the one or more condensation zones 11 ;

e) a recycling zone 12 inside the housing 2 which is connected to the one or more reaction zones 4 and to the one or more tubes 7 for convective mass exchange between the one or more tubes 7 and the one or more reaction zones 4 via the recycling zone 12;

f) a catalyst bed 13 positioned in each of the one or more reaction zones 4;

g) a cooling means 14 positioned in each of the one or more condensation zones 11 for producing a liquid condensate product;

h) one or more receptacles 15 for collecting a liquid condensate product, said one or more receptacles 15 being positioned in the housing 2 such that the liquid condensate product formed by condensation in the one or more condensation zones 11 cannot contact the catalyst bed 13;

i) an outlet 16 for removal of liquid product from the one or more receptacles 15 inside the housing 2; and

j) optionally a purge 17 for removal of gaseous components from the housing 2. The reactor can advantageously be applied in the process for the production of a liquid methanol and water condensate or a liquid methanol condensate as defined herein before.

The housing 2 of the reactor 1 preferably is bar-shaped. The shape of the top 9 and the bottom 10 planes of the housing 2 can be independently chosen from circular, rectangular, triangular, or multi-angular shapes. In a preferred embodiment, the housing 2 has a shape selected from a cylindrical or a rectangular bar. The diameter or the length and width of the top 9 and the bottom 10 planes of the housing 2 can be chosen independently.

The reactor 1 according to the invention is not limited as regards the aspect ratios of height versus the diameter of the bottom plane, height versus the length of the bottom plane or height versus width of the bottom plane. Similarly, the reactor 1 according to the invention is not limited as regards the aspect ratios of height versus the diameter of the top plane, height versus the length of the top plane or height versus width of the top plane.

The housing 2 of the reactor 1 has an imaginary vertical axis 6 as indicated by the dashed line in Figures 1-5. The housing 2 of the reactor 1 further has an imaginary plane 5 perpendicular to the imaginary vertical axis 6 of the housing 2, as indicated by the dashed line in Figures 1-5.

The housing 2 comprises one or more reaction zones 4 covering only part of the cross- sectional area of the plane 5 perpendicular to the vertical axis 6 of the housing 2 and one or more tubes 7 covering at least part of the cross-sectional area not covered by the one or more reaction zones 4. In a preferred embodiment, the one or more reaction zones 4 and the one or more tubes 7 cover the full cross-sectional area of the plane 5 perpendicular to the vertical axis 6 of the housing 2.

The one or more reaction zones 4 and the one or more tubes 7 are separated by walls 8 parallel to the vertical axis 6 of the housing 2. The reactor 1 is not limited as regards the arrangement of the one or more reaction zones 4 and the one or more tubes 7 in the housing 2. Non-limiting examples of reaction zone and tube arrangements include a single annular reaction zone surrounding a single tube 7, a single annular tube 7 surrounding a single reaction zone 4 and a multi-tubular system comprising multiple reaction zones 4 placed in a tube 7. The term 'annular' as used herein does not only refer to a circular annular zone or tube but can for example also refer to a rectangular annular zone or tube.

Each of the one or more reaction zones 4 comprises a catalyst bed 13. Any type of catalyst bed may be used, such as for example a fixed bed consisting of catalyst particles or a monolithic, metallic, carbon or inorganic structure supporting catalytically active components. In a preferred embodiment, the catalyst bed 13 does not completely cover the cross-sectional area of the reaction zone 4 perpendicular to the vertical axis 6 of the housing 2. In this way, gaseous components can freely flow through the reaction zone 4 and can contact the catalyst bed 13. In a more preferred embodiment, the catalyst bed 13 is a porous catalyst bed which completely covers the cross-sectional area of the corresponding reaction zone 4 perpendicular to the vertical axis 6. This means that the porous catalyst bed acts like a plug closing the reaction zone 4. Gaseous components can contact and cross the catalyst bed through the pores. In a preferred embodiment, the pressure drop across the porous catalyst bed is less than 0.05 MPa, more preferably less than 0.001 MPa.

Preferably, every reaction zone 4 has a means 18 for regulating the temperature inside the reaction zone 4, more particularly the temperature of the catalyst bed 13. If the reactor 1 comprises a single tube 7 which is surrounded by an single annular reaction zone 4, the means 18 can be positioned outside the housing 2, for example as a heating jacket surrounding part of the housing 2. In other cases, a means 18 can be placed in every reaction zone 4, for example close to the catalyst bed 13 or even inside the catalyst bed 13.

Walls 8 do not reach to the top 9 of the housing 2 and further do not reach to the bottom

10 of the housing 2. The space between the bottom 10 of the housing 2 and the lowest point of the walls 8 and the space between the top 9 of the housing 2 and the highest point of the walls 8, respectively constitute the recycling zone 12 and the one or more condensation zones 11, or vice versa. In order to realize convective mass exchange by natural convection, it is required that the space between the top 9 of the housing 2 and the highest point of the walls 8 constitutes the one or more condensation zones 11. Hence, in a preferred embodiment, the one or more condensation zones 11 are positioned above the one or more reaction zones 4 and the recycling zone 12 is positioned below the one or more reaction zones 4.

The one or more cooling means 14 are positioned in the one or more condensation zones 11. Preferably, every cooling means 14 is also partly positioned in a tube 7. A receptacle 15 is positioned under every cooling means 14 such that any liquid condensate product formed in the one or more condensation zones 11 is collected in or on a receptacle 15. The receptacles 15 are placed in such a position that the liquid condensate product cannot contact the catalyst bed 13. Examples of receptacles 15 include a simple collector bin or a plane having an opening above tube 7, said opening having raised edges and being covered by a cap 22 to prevent leakage of liquid condensate product into the tube 7. Examples of caps 22 are "bubble caps" or "chimney-type devices" which are well known to the person skilled in the art of process technology and are here used to allow the gas to flow downward while maintaining the liquid on the plane receptacle. If the reactor 1 comprises more than one reaction zone 4 and/or one or more tubes 7, the plane acting as receptacle can be a single plane covering all tubes 7 and having an opening above every tube 7 or multiple openings per tube 7. Such a plane may be an inclined plane such that liquid condensate product collected onto the plane runs to one side of the reactor 1 where it can be removed via outlet 16. The rate of liquid condensate product removal must be such that no liquid is entering the one or more tubes 7.

Liquid condensate product can be removed from the reactor 1, both when operated batch- wise, continuously or semi-continuously. In an embodiment, the liquid can be removed from the housing by pumping the liquid condensate product through liquid outlet 16. If the reactor is operated at an overpressure compared to ambient pressure, liquid condensate product can be removed via liquid outlet 16 by opening a valve in liquid outlet 16. The overpressure in the housing will then force the liquid product through liquid outlet 16.

As explained in the foregoing, the one or more reaction zones 4 and the one or more tubes 7 are separated by walls 8. Irrespective of the flow pattern in the reactor, the gaseous components in the one or more reaction zones 4 always flow in the opposition direction as the gaseous components in the one or more tubes 7. The one or more condensation zones 11 are operated at a lower temperature than the one or more reaction zones 4. This means that hot gas leaving the one or more reaction zones 4 is cooled in the one or more condensation zones 11 to produce a liquid condensate product and that the remaining cooled gaseous components are recycled to the one or more reaction zones 4 where they must be heated again to the reaction temperature. To reduce energy consumption required by heating and cooling the gaseous components, heat can be exchanged between the gas streams leaving the one or more reaction zones 4 and the gas streams flowing through the one or more tubes 7. Hence, in a preferred embodiment, a part 21 of the wall 8 has heat exchanging properties. In the context of the present invention, the wall 8 has heat exchanging properties if the thermal conductivity is at least 10 W-m^-K ^"1 . Part 21 may have a surface geometry that enhances heat exchanging properties. Non-limiting examples of such a surface geometry are surfaces in the form of a corrugated fin, surfaces being covered with pins, and surfaces having grooves or dimples.

If the one or more condensation zones 11 are positioned above the one or more reaction zones 4 and the recycling zone 12 is positioned below the one or more reaction zones 4, the part 21 of the wall 8 that has heat exchanging properties is preferably positioned just above catalyst bed 13. If the one or more condensation zones 11 are positioned below the one or more reaction zones 4 and the recycling zone 12 is positioned above the one or more reaction zones 4, the part 21 of the wall 8 that has heat exchanging properties is preferably positioned just below catalyst bed 13. In case the reactor is operated in a continuous or in a semi-continuous way and the gaseous mixture fed to the housing 2 via inlet 3 comprises inert gaseous components and/or gaseous components that do not react to products that are condensable in the condensation zone 11, gaseous components may accumulate in the reactor. In such a case it may be necessary to purge part of the gaseous components inside the housing 2 to the environment at regular time intervals or continuously. Hence in an embodiment, the reactor 1 comprises a purge 17. Preferably, the purge is positioned such that the gaseous components to be purged comprise the least amount of reactants and condensable product. Hence, in a preferred embodiment the purge 17 is positioned downstream of the condensation zone 11, i.e. after separating the liquid condensate product, and both upstream of the catalyst bed and upstream of the point where remaining gaseous reactants are mixed with fresh gaseous feed.

All zones in the reactor 1, i.e. the one or more reaction zones 4, the one or more condensation zones 1 1, the one or more tubes 7 and the recycling zone 12, are interconnected such that convective mass exchange can take place between the zones.

In an embodiment, the housing 2 comprises one or more means for forced circulation 20 of the gaseous mixture over the catalyst bed 13 via the one or more condensation zones 11, the one or more tubes 7 and the recycling zone 12 or vice versa. The one or more means for forced circulation 20 are preferably positioned in the recycling zone 12. In another preferred embodiment, the one or more means for forced circulation 20 are at least partly positioned in the recycling zone 12 and at least partly in the one or more tubes 7. In still another preferred embodiment, the one or more means for forced circulation 20 are positioned in the one or more tubes 7. Examples of means for forced circulation 20 include fans, blowers and compressors. In a preferred embodiment, the one or more means for forced circulation 20 have an inlet connected to the one or more tubes 7 and an outlet connected to the one or more reaction zones 4. In this configuration, the one or more means for forced circulation 20 allow for circulation of a gaseous mixture from the one or more reaction zones 4, to the one or more condensation zones 11 and via the one or more tubes 7 and the recycling zone 12 back to the one or more reaction zones 4. In another preferred embodiment, the one or more means for forced circulation 20 have an outlet connected to the one or more tubes 7 and an inlet connected to the one or more reaction zones 4. In this configuration, the one or more means for forced circulation 20 allow for circulation of a gaseous mixture from the one or more reaction zones 4, to the recycling zone 12, via the one or more tubes 7 to the one or more condensation zones 11 and back to the one or more reaction zones 4. In another preferred embodiment, the reactor comprises a heating means 19 for preheating the gaseous mixture that enters the housing 2 via inlet 3. Examples of heating means 19 that can be used to pre-heat the gaseous mixture are well-known in the art such as an oil or water heater or an electric heating element. Pre-heating may be advantageously used to decrease the time required for the start-up phase of a batch-wise operated reactor. Pre-heating of the gaseous mixture entering the housing may also advantageously be used in a continuously or semi-continuously operated reactor to decrease costs related to means 18 for regulating the temperature inside the one or more reaction zones 4.

DETAILED DESCRIPTION OF THE DRAWINGS

The reactor of the invention will be further illustrated by means of the following, non- limiting drawings.

In Figure 1 is schematically shown a reactor 1 according to the invention comprising a means for forced circulation, a single condensation zone 11 and a single reaction zone 4 surrounding a single tube 7. Gaseous reactants are fed to the reactor via inlet 3. Inlet 3 comprises a heating means 19 to adjust the temperature of the gaseous reactants to the required temperature. The heating means 19 is positioned outside the housing 2. The housing 2 comprises a tube 7 which is surrounded by an annular reaction zone 4. The wall 8 separates the tube 7 from the annular reaction zone 4. Wall 8 runs parallel to the vertical axis 6 of the housing 2 and does not reach to the top 9 of the housing 2 and further does not reach to the bottom 10 of the housing 2, allowing for convective mass exchange between the tube 7 and the annular reaction zone 4 both near the top 9 of the housing 2 and near the bottom 10 of the housing 2. A catalyst bed 13 is positioned in the annular reaction zone 4. The reactor 1 in Figure 1 comprises an external means 18, for example a heating jacket, for regulating the temperature inside the annular reaction zone 4. The space between the top 9 of the housing 2 and the wall 8 constitutes the condensation zone 11. This condensation zone 11 comprises a cooling means 14 such as a condenser coil for condensation of a liquid condensate product. As is shown in Figure 1, the cooling means 14 is partly positioned in the tube 7. A receptacle 15 is positioned under the cooling means 14 such that any liquid condensate product formed in the condensation zone 11 is gathered in the receptacle 15. Hence, the liquid condensate product cannot contact the catalyst bed 13. The space between the bottom 10 of the housing 2 and the wall 8 constitutes the recycling zone 12. This recycling zone 12 comprises a means for forced circulation 20 of gaseous components through the housing 2, for example a fan. The means for forced circulation 20 forces circulation of gaseous components via tube 7, via the condensation zone 11, the reaction zone 4 and the recycling zone 12 back to the tube 7. By choosing a different fan configuration, fan position and/or a different position of the inlet 3, the flow pattern of gaseous components can also be directed into the opposite direction, i.e. via reaction zone 4, condensation zone 11, tube 7, recycling zone 12 back to the reaction zone 4. The liquid condensate product can be removed from the receptacle 15 via liquid outlet 16. The reactor comprises a purge 17 for purging gas to the environment. As explained in the foregoing, such a purge 17 is an entirely optional feature of the reactor 1.

Figure 2 shows a reactor 1 according to the invention comprising a single condensation zone 11, a single annular reaction zone 4 surrounding a single tube 7, said reactor 1 not having a means for forced circulation. The reactor of Figure 2 is similar to the reactor of Figure 1. However, the inlet 3 for feeding gaseous reactants to the housing 2 and the optional purge 17 are positioned differently. Furthermore, the means 18 for regulating the temperature inside the annular reaction zone 4, more particularly the temperature at the catalyst bed 13, is positioned inside the housing, right under the catalyst bed 13 in the reaction zone 4. This configuration of the reactor 1 allows for gas circulation via natural convection, i.e. from the reaction zone 4, via the condensation zone 11, the tube 7 and the recycling zone 12 back to the reaction zone 4.

In Figure 3 is schematically shown a multi-tubular reactor according to the invention comprising more than one reaction zone 4, more than one condensation zone 11 and more than one cooling means 14, said reactor not having a means for forced circulation. The reactor shown in Figure 3 differs from the reactor in Figure 2 in that it has more than one reaction zones 4. Every reaction zone 4 has a means 18 for regulating the temperature inside the reaction zone 4, more particularly the temperature at the catalyst bed 13, positioned inside the housing, right under the catalyst bed 13 in the reaction zone 4. The part of the cross-sectional area of the plane 5 not covered by the reaction zones 4 constitutes the tube 7. This configuration of the reactor 1 allows for gas circulation via natural convection.

Figure 4 shows a simple configuration of a reactor 1 according to the invention comprising a single reaction zone 4, a single condensation zone 11 and a single tube 7, said reactor 1 not having a means for forced circulation. The single reaction zone 4 is not annular and does not surround the tube 7. The reaction zone 4 and the tube 7 are positioned next to each other, separated by wall 8. Part 21 of the wall 8 has heat exchanging properties. The means 18 for regulating the temperature inside the reaction zone 4 is placed inside the catalyst bed 13. This configuration of the reactor 1 allows for gas circulation via natural convection.

Figure 5 shows a reactor 1 which is a special embodiment of the reactor shown in Figure 3. The part of the cross-sectional area of the plane 5 not covered by the reaction zones 4 constitutes the tube 7. The receptacle 15 is configured as an inclined plane positioned in tube 7 with openings having raised edges. Every opening is covered by a cap 22 to prevent leakage of liquid condensate into the tube 7. Liquid condensate product gathered onto the inclined plane of the receptacle 15 runs to one side of the housing 2 where it can be removed via liquid outlet 16.

Figure 6 shows a top view of the inclined plane formed by the receptacle 15 as shown in

Figure 5. The viewing direction is indicated by VI in Figure 5. In order to make Figure 6 more clear the top 9 of the housing 2, the condensation zones 11 and the cooling means 14 are not shown.

The invention and some of its embodiments can also be worded as follows. The present invention comprises a 'gas In, liquid out' reactor concept in which C0 ₂ is reacted with a co- substrate, preferably H ₂ , to form a liquid chemical, preferably methanol, whereby the thermodynamic equilibrium constraint is surpassed by creating an internal condensation zone (at lower temperature) in the reactor, next to a reaction zone (at higher temperature).

Flow circulation between the zones inside the reactor can be enforced mechanically, but according to a preferred embodiment the hot reaction zone and cold condensation zone are placed such, that this is supported, or replaced, by natural convection, induced by the temperature difference.

In a preferred configuration, the reactants are fed in stoichiometric amounts to the reactor and fully converted into the liquid phase product. The reactor then operates in essentially a 'gas in - liquid out' (GILO) mode of operation, ensuring near complete carbon efficiency with only minor losses due to dissolved reactants in the product phase.

In a further preferred configuration, the reactor with auxiliary cooling and heating equipment is designed in such a way that it can handle intermittent supply of electricity and reactants.

In an embodiment, the invention relates to a method for the synthesis of a product from one or more gaseous reactions, using in-situ condensation, and removal of all products via the liquid phase, i.e. the 'gas-in liquid-out' concept. The in-situ condensation is preferably used to circumvent reaction equilibrium constraints. Preferably, the in-situ condensation is realized by creating zones of different temperature, e.g. for reaction and for condensation, inside the reactor embodiment and mass exchange between reaction and condensation zone(s) is realized by convection. This convection preferably is realized by natural convection induced by the temperature difference between the zones. In a preferred embodiment, the method for the synthesis of a product from one or more gaseous reactions is a method for the production of methanol from C0 ₂ , as dominant carbon source (>90%), and H ₂ .

In another embodiment, the invention relates to a method for methanol production starting from C0 ₂ and H ₂ as input, producing a liquid condensate consisting of methanol and water, in a 'gas-in liquid-out' reactor with minimally one hot reaction zone and minimally one colder condensation zone with convective mass exchange between the zones. In a preferred embodiment, said mass exchange is realized by natural convection on basis of the induced temperature differences.

The following examples are meant to further illustrate the invention and some of its preferred embodiments without intending to limit its scope.

EXAMPLES

Temperature and pressure do not only determine the reaction rate for methanol production, but also the position of reaction equilibria and the dew point of the gas phase mixture (relevant to condensation). As a high operating pressure favors methanol production, but will simultaneously increase capital cost and gas compression cost, the optimum pressure is most likely in the range between 0.1 and 10 MPa, preferably between 1 and 6 MPa. A Proof of concept run is designed at 5 MPa total pressure. The reaction temperature is most likely in the range of 175-300°C, preferably between 190 and 250°C. When operating the reactor at a single temperature (around 190-240°C), the required reactor pressure for condensation is very high (above 150 bar). By creating a colder condensation zone inside the embodiment, the concept of in-situ condensation can be realized at more favorable pressure levels (50 bar or less), while maintaining a high reaction rate in the warmer reaction zone.

Example 1

Prior to use, a GILO reactor according to the invention was loaded with 79 g of a commercial methanol catalyst (JM CP-488) and activated. The reactor was run at 215°C, at the top of the catalyst bed, and at approximately 5 MPa of total pressure. During the run, a stoichiometric H ₂ :C0 ₂ (3 : 1 mol/mol) gaseous feed was fed to the GILO reactor, which was kept at 4.85 MPa overpressure for most of the run (up to 410 minutes run time).

The required gas feed rate to maintain the reactor pressure was monitored as mbar/s of pressure drop in the feed vessel. This gas feed rate is a measure for the rate of methanol production. At fixed intervals, approximately every 0.25 MPa of pressure drop in the supply vessel, liquid condensate was tapped from the reactor. The liquid condensate was found to be approximately 50/50 in molar ratio of water/methanol via FTIR analysis. Internal gas circulation was enforced by a mechanical fan. Under these operating conditions a liquid (methanol/water) condensate was produced under steady state conditions for at least one hour (between 330-410 minutes of run time). At 412 minutes of run time, the mechanical 'forced' -circulation of the gas inside the reactor was switched off. From 412 to 438 minutes of run time, feed gas consumption continued indicating methanol production under natural convection circulation and a condensate was tapped from the reactor at approximately 440 minutes of run time. At 443 minutes of run time, the reactor was cooled and the reaction was stopped.

Example 2

In this example, the test equipment consisted of a 0.7 liter multifunctional reactor with two temperature zones and a tube-in-a-tube design and an internal fan to provide forced convective gas circulation inside the test equipment. Hence, unreacted gaseous components are recycled to the catalyst bed. In the electrical heated bottom zone, a conventional Cu/ZnO/Al ₂ 03 methanol catalyst was placed in the annular spacing. In the water cooled top zone inside the tube a collecting cup was placed to collect the produced condensate. The reactor pressure was maintained at a constant level by a pressure reducer between the reactor and a 3.8 liter supply vessel containing feed gas at a higher pressure.

The reactor was continuously fed with a premixed mixture of 25% C0 ₂ in 75% H ₂ from the feed supply vessel to maintain the reactor pressure at 5 MPa. The reaction zone was operated at 195°C and the cold zone at 90°C. The feed rate required to maintain the reactor pressure was about 36 NmL/min. Produced condensate was collected and tapped from the reactor for every 4 mL of produced condensate. The reactor was operated for 2 hours without a need to purge gaseous components in between. The produced condensate approached a 50/50 molar mixture of methanol/water with less than 0.2 mol% of impurities, said impurities mainly consisting of ethanol and formic acid. Using this test equipment under said conditions, including a means for forced circulation, about 6 kg MeOH/kg _cat /day could be produced.

Example 3

The test equipment of example 2 was used, now without making use of an internal fan to provide convective gas circulation. The reactor was operated at a pressure of 5 MPa, the hot zone at a temperature of 195°C and the cold zone at a temperature of 90°C. The feed rate of a premixed mixture of 25% C0 ₂ in 75% H ₂ required to maintain the reactor pressure was now about 30 NmL/min. The reactor was operated for 2 hours without a need to purge gaseous components in between. The produced condensate approached a 50/50 molar mixture of methanol/water with less than 0.2 mol% of impurities, said impurities mainly consisting of ethanol and formic acid. Using this test equipment at said conditions, but now without a means for forced circulation, about 5 kg MeOH/kg _cat /day could be produced. Since this production rate is about the same as in Example 2, convective gas circulation inside the test equipment driven by natural convection is shown to be effective.