NO823211L - SYNTHESIS gas system. - Google Patents

Abstract

Process and apparatus for producing synthesis gas from hydrocarbonaceous gases, in which a hydrocarbonaceous gas is preheated, reacted with oxygen in the presence of steam and cooled using a rotating matrix comprising glassy ceramic fibers. After reaction with oxygen, the gases can be passed over a catalyst to further convert the said hydrocarbon-containing gases into synthesis gas. This method and apparatus are particularly useful in a process for producing methanol in which carbon dioxide in the synthesis gas is absorbed using crude methanol and recycled. by stripping the CO-rich crude methanol msd gas which has passed over the catalyst for the conversion of synthesis gas to methanol under conversion conditions in which it is thereby recycled and finally recovered as methanol yield carbon present in both carbon dioxide and unconverted methane present in said gases. The process requires neither expensive reforming furnaces nor expensive synthesis gas compressors which are conventionally required by steam reforming methods and use much less oxygen than in conventional partial oxidation processes. This process is characterized by further methanol production by means of partial oxidation processes and apparatus so that they comprise an even more efficient, integrated process for the production of methanol, in which all methane carbon used in the combined process is converted to methanol with the exception of about 7% required as fuel for energy for efficient separation of the required oxygen from air and compression thereof to the operating pressure.

Description

SYNTHESIS GAS SYSTEM

BACKGROUND OF THE INVENTION

Up f i nn e1ses area e t

The invention is within the area of production of synthesis gas and methanol.

Description of prior art

Hydrocarbon-containing materials, including petroleum oil and natural gas, coal etc., i.e. fossil energy, are becoming increasingly expensive and less available. As long as petroleum oil was abundantly available, transportation and other liquid fuels could be easily obtained using well-known and relatively cheap refining processes.

However, with increasing scarcity, such refined liquid fuels will no longer be sufficient and extensive research and development efforts have been carried out to supplement such fuels. It is e.g. well known that "synthesis gas", namely a mixture of mainly carbon monoxide and hydrogen, can be partially converted over suitable catalysts into a wide variety of increasingly valuable products, including hydrogen, carbon monoxide, ammonia and hydrocarbon and oxyhydrocarbon liquids. Possible liquid products include methanol, ethanol, higher alcohols, ethylene glycol and liquid hydrocarbons. For the most part, these can be used either directly or converted further into chemical products or into liquid fuels.

This synthesis gas can be derived and has been derived from any of the hydrocarbon-containing raw materials by reacting these with steam and/or oxygen at relatively high temperatures. The major disadvantage of such previous processes, especially those which convert one fuel into another fuel, is, however, that the thermal energy efficiency is relatively low, so that you are left with significantly less fuel heating value than you assumed. Processes used up to now thus result in a solution with increasing scarcity of and increasingly expensive fossil energy.

In a recent carefully considered comparison of the state of the art carried out by the experienced Lurgi Company (Gasonlin Production from Natural Gas or Coal. E. Supp, 1980, table II) the following thermal efficiencies (based on the lower heat values in starting materials and end products with electricity) were thus as a by-product calculated back to equivalent heat to an easily accessible basis) shown to be economically achievable.

Thermal efficiencies %

In case I, conventional steam reforming of natural gas is used to produce synthesis gas, which is then compressed and catalytically converted into methanol. To produce gasoline, the 1:methanol is catalytically converted using the methanol-to-gasoline (MTG) process of Mobil Corporation.

Case II is similar except that oxygen is added in another stage of (autothermal) reforming, Case III autothermally reforms coal with steam and oxygen and in Case IV the purge gas is catalytically methanated to form a synthetic natural gas. In all cases, fuel required for all steps in the process is supplied and included in the calculations.

Several important conclusions can be drawn from the values given above. First, the previous attempts to convert coal into more useful liquid fuels have a higher thermal efficiency than other methods: Fischer-Tropsch (about 40%) and (improved) coal hydrogenation (about 54%). Secondly, natural gas is much more efficient as a raw material than coal. On the basis of energy loss, comparing the highest values given, over 24% more energy is lost using coal than using natural gas. Interpretation of total energy considerations therefore favors the use of natural gas for this purpose, provided of course that there are other uses for coal where its relative thermal efficiency is higher than in the above-mentioned use (such a use is clearly direct development of electricity in steam-electric plants where coal should replace natural gas in that their thermal efficiencies are roughly the same).

In addition, the total capital investment costs for using coal to produce liquid fuels are about 3 times as high as for natural gas. Consequently, the price of natural gas would have to rise to approximately 6 times the price of coal before coal becomes economically competitive for these purposes, as coal wastes both scarce energy and scarce capital. Fortunately, it is believed that sufficient natural gas has been found and will continue to be found in the world so that it is almost as abundant as coal so that natural gas will be favored for these purposes from an economic standpoint for many decades. It is then seen that a surplus of relatively low-temperature by-product heat is produced and which can be converted into electricity in case I which relates to conventional steam reforming. However, this does not improve the overall thermal efficiency to any great extent. A much better practice is to go to case II. This would, however, require increased investment in an oxygen plant for air separation and compressor, which has rarely been economically justified or practised.

Finally, it can be seen that even with the various combinations, the thermal efficiencies are still quite low and represent a very significant energy loss.

SUMMARY OF THE INVENTION

In light of the foregoing, it is a general purpose of the present invention to provide a new method and apparatus for more efficient production of synthesis gas from natural and other similar gases in which the thermal efficiency is substantially increased.

When producing methanol from methane gas, e.g. using

the preferred integrated process as described below, it is shown that loss in carbonaceous feed and fuel conversion efficiency can be lowered by a factor of 4 from the equivalent for conventional steam reforming, i.e. from a loss of

about 28% of carbon in feed and fuel to as low as about 7% (corresponding then to an overall thermal efficiency as high as about 80% on a basis comparable to those given above).

Another purpose of the invention is to utilize the inherent pressure of most of the natural gas when it is produced from its source in an efficient manner.

Another object of the invention is to reform the feed under pressure sufficiently high that the synthesis gas does not need any compression for many of its applications, e.g. usually including the production of hydrogen, carbon monoxide, methanol and gasoline using the MTG process.

It will be understood by those skilled in the art that in conventional steam reforming, which is highly endothermic and in which the feed gas is heated inside catalyst-filled tubes in a large heated furnace, the allowable tube wall temperatures are severely limited, despite the use of very expensive materials for the purpose. And furthermore, it is well known that the steam reforming reaction:

because of its production of four molecules from two, is very strongly affected in a harmful direction by the pressure, i.e. the reaction will reverse in accordance with the square of the pressure.

Consequently, conventional steam reforming methods have been limited to a range of about 18 to 30 atmospheres of pressure and thus require compression of this large volume of low molecular weight gas in many stages with centrifugal compressors.

Another object of the invention is, in the desired production of only methanol, to eliminate excess production of hydrogen. In the case of catalytic methanol production using the reaction:

a H^/CO ratio of 2 is desired, while the above-mentioned steam reforming reaction produces a ratio of 3. This excess hydrogen requires a high degree of emission of gases that are discharged from the methanol converter and carries with it excessive losses of methanol vapor and other carbonaceous constituents as well as hydrogen. In addition, this excess hydrogen produces a very non-optimal converter gas composition, much higher than the ratio

3 to 1 in the feed gas, due to the removal in the ratio

2 to 1 in the form of methanol.

It is another object of the invention to provide and consume heat with maximum profitable efficiency so that no excess of low-level heat is lost.

It is another purpose of the invention to utilize autoter-

production of synthesis gas in which total oxygen as cre-

ves is conserved to only a fraction of the amount required in normal autothermal synthesis gas processes.

Another purpose of the invention is to optimize the composition of the feed synthesis gas to a methanol converter so that the rate of methanol formation over the catalyst is maximized and so that all gas that must be removed from there is used in a useful way by providing driving force for the process, including compression of air for the air separation plant and compression of the oxygen therefrom.

Another object of the invention is to provide a "support" fuse for further reduction of the content of harmful methanol catalyst poisoning, especially sulfur compounds including U^ S and COS so that the catalyst life is increased (especially for the extremely dense models so-called "low temperature" catalysts for methanol production).

Another object of the invention is to provide a methanol converter which will allow catalytic conversion at much larger capacities in a single container than hitherto practically possible with exothermic reactions, in which at least a significant proportion of the heat released by and during the course of the reaction is absorbed by means of a tubular stream of liquid which is thereby heated or brought to a boil.

Another purpose of the invention is to provide an integrated process and apparatus for the production of methanol from natural gas which is capable of very high capacities and yet is relatively compact in that it is suitable for the main and efficient production of liquid fuel from natural gas on adjacent locations, including barge-mounted and other off-shore units.

In accordance with the invention, synthesis gas is produced much more efficiently by preheating the hydrocarbon-containing feed to a temperature of more than 816°C, the hydrocarbon-containing gas reacts autothermally with the oxygen in the presence of steam and then the hydrocarbon gas is reformed. In direct contrast to the conventional steam reforming method, no intentional reforming is carried out below 816°C or during heat exchange heating. Consequently, the requirements for both steam, in comparison with conventional steam reforming, and oxygen, in comparison with so-called partial oxidation systems, are reduced to a minimum and it becomes practical to carry out the reforming reaction under such high pressures that subsequent necessary compression of the synthesis gas is reduced to a minimum, and in many practical applications of the synthesis gas the need is completely eliminated. This is an important economic advantage due to the fact that these compressors are large and very expensive and of course require large amounts of energy.

In accordance with the invention, the preheating step is preferably carried out during the reforming of the hydrocarbon-containing gas by heat exchange with the reformed product from the reforming reaction. This heat exchange is preferably carried out by using a rotating heat regenerator with radial flow where a sectioned matrix consists of ceramic fibers in glassy form, a construction of such being described in the parallel application 82. (US Serial No. 228,909) . In this way, the contact material in the heat exchange step is inert to the Boudouard reaction with carbon deposition, as more fully explained below in the detailed description. It is further preferred that oxygen is also preheated in part of the rotating matrix so that heat is not only conserved but also that any carbon that is deposited on the matrix will be gasified by the oxygen so that it is ensured that no clogging of carbon

can act.

Furthermore, in accordance with an important aspect of the invention, the steam required in the reforming step is generated by passing the hydrocarbon-containing gas feed, together with liquid water, in heat exchange conditions with a heat-producing catalytic reaction of the synthesis gas or a gas derived from the synthesis gas. This important step produces the steam in the reformer reactant gas at a higher total pressure than would otherwise be possible by developing steam separately and adding it to the reactant gas, again helping to eliminate synthesis gas compressors, particularly in the production of methanol from the synthesis gas.

In another important aspect of the invention, the step of generating heat in such a heat-producing catalytic reaction and recovering the heat in a useful manner while thereby controlling the temperature therein is advantageously carried out in a new catalyst-layer structure with radial flow with transverse heat-absorbing tubes, so that not only the average layer temperature is controlled but also the temperature profile along the radical flow path. This structure also entails very practical facilities for removing the old catalyst for replacement with a new one.

Another aspect of the invention relates to the carbon dioxide product from the reforming reaction. Thus, the water-gas exchange equilibrium prevailing under the reforming conditions requires CO2 to be present in accordance with the equilibrium for the reaction: CO + H20 = C02+ H2>For a given ratio H2/CO (about 2), C02 is required proportional to steam. Minimizing the steam requirement as with the present invention thus minimizes C02, but what is left still requires excess hydrogen for the production of methanol in accordance with the reaction: C02+ 3H2= CHeOH + H20 and produces unwanted water in the methanol.

In accordance with a preferred aspect of the invention when producing methanol from synthesis gas, carbon dioxide is therefore absorbed in the synthesis gas using a liquid absorbent and the carbon dioxide is then stripped from there and returned to the reforming zone, so that the net production of CO^, and water in methanol, becomes small or can be neglected. It is advantageous to use the crude methanol itself as a liquid absorbent.

Furthermore, since the synthesis gas comprises unconverted methane as well as carbon monoxide and hydrogen and since methane is inert to and builds up in the recirculation loop for the methanol conversion, another aspect of the invention is to use hot exhaust gases from the catalytic converter, especially with their methanol vapor content, as a stripping agent for to strip CO^ from the C02~-rich absorbent. This step not only provides stripping gas and heat, but also recycles unconverted methane to the reforming zone.

To further reduce unconverted methane build-up as an inert component in the recycle loop for the methanol converter and still preserve the hydrogen content of this gas so that its ratio to carbon monoxide can be optimized, to more than 2, while its production ratio is only about 2, depressurize in accordance with the invention raw methanol to a lower pressure than the methanol synthesis pressure so that the higher content of dissolved methane evaporates and then recycles at least part of this raw methanol and brings it further into contact with the converter gas.

In accordance with the invention, in a preferred embodiment for the production of methanol, the net production of crude methanol with an insignificant sulfur content is advantageously used to contact the synthesis gas so that sulfur-containing pollutants are removed from there and thereby protect the extremely sulfur-sensitive modern so-called "low temperature" catalysts for the conversion of synthesis gas to methanol. Furthermore, by taking the final net production of crude methanol directly from the hot CO^-poor methanol after it has been heated and stripped of CO^ and then stripping it with the (mainly sulfur-free) methanol converter exhaust gases that are flushed from the system in order to remove inert components including nitrogen, argon and unconverted methane, sulphur-containing as well as other pollutants are stripped from the product methanol so that it is prevented from being polluted without particular consumption of additional energy.

In another important aspect of the preferred embodiment of the invention, the above-mentioned reforming/catalytic conversion method is combined with the direct production of methanol from methane by partial oxidation according to the aforementioned concurrent application (USSN 228,909). This combined process has numerous advantages as more fully explained in the following detailed description, and the partial oxidation method cooperates in several advantageous ways with the aforementioned reforming/catalytic method. Thus, it provides heat to the method as well as methanol and thus helps to strip CO^ from the C02~-rich absorbent and to purify the methanol, as well as to produce an already partially converted off-gas from it.

As shown in the following, all parts of the invention work together to produce and consume synthesis gas with extremely high efficiency, as partially shown in the case of methanol production, where a total yield of more than 93% is achieved

in comparison with the yield of approximately 7 2% according to the state of the art. Similar advantageous conservation of energy will occur in other applications of synthesis gas, such as e.g. production of hydrogen, ammonia, etc.

Further objects and advantages of the invention will be apparent from the subsequent part of the description in which the details of the method and the apparatus in preferred embodiments are described with reference to the attached drawings in which:

BRIEF DESCRIPTION OF THE DRAWINGS

Fig. 1 is a flowchart of the preferred embodiment of the method for producing synthesis gas and methanol in accordance with the invention in combination with the method and the apparatus described and indicated in the concurrent application no.

(USSN 228.909) .

Fig. 2 is a schematic horizontal section taken along the line 2-2

in fig. 1 and shows the upper cross-section of the preferred embodiment of the reforming container designated 31 in fig. 1. Fig. 3 is an enlarged, vertical partial section taken along the line 3-3 in fig. 2, and Fig. 4 is an enlarged vertical axial section of the preferred embodiment of the methanol converter designated 28 in Fig. 1.

DETAILED DESCRIPTION

Methanol process

With reference to fig. 1, the gas supply in line 12, preferably natural gas with a high methane content, under a pressure preferably in the range from about 30 to about 150 atmospheres absolute (ata) and more preferably in the range from about 50 to 80 ata, is introduced into the recycled reaction exhaust gas in the line 13, so that it thereby cools it further, and the mixture is washed with the help of washing water that enters the container 10 from the line 14 to primarily remove methanol produced by partial oxidation of methane as described in detail in the previously mentioned application (USSN 228,909). Schematically shown in the vessel 10 are two sections of parallel countercurrent scrubbing which would be practical for large capacity units (eg over 1500 tonnes per day total methanol capacity). These sections can be bubble sheet sections or filled sections as is well known in the art.

The scrubber exhaust gases 15 from the container 10 are circulated using

of the recirculation fan 16 and the main part is preferably mixed with compressed oxygen in the line 11 so that the mixture comprises between approximately 0.5 mol% and 5 mol% oxygen, preferably 1-3%

(produced by an air separation plant including an air com-

pressor and oxygen compressor which are preferably driven by steam and/or gas turbines which in turn receive their steam and fuel at least mainly from the process and which are not shown here) immediately before it enters the heat exchanger reactor 17, shown schematically in fig. 1. As described in the aforementioned second application, the preferred heat exchanger reactor 17 is a rotary radial flow regenerative heat exchanger rotating at between 2 and 20 RPM, preferably 5-12 RPM, with the reaction space towards its center.

In the sectioned matrix 19 in the reactor 17, the gas mixture is heated from approximately 121°C to an average temperature

preferably between about 482°C and about 649°C. Because this matrix comprises a greatly expanded surface, preferably several layers of (inert) ceramic or glass cloth, the partial oxidation reaction therein is inhibited and is minimal. When the reactant gas comes out of the matrix 19, however, it is at a sufficient temperature so that it will quickly self-accelerate with regard to reaction rate and relatively quickly (0.1 - 2 seconds) consume essentially all the oxygen. This exothermicly reacted mixture has gained a substantially increased temperature, roughly in relation to its original oxygen content, and is now hotter than the matrix 19 which preheated it. It passes back through the matrix 19, giving off sensible heat down to a temperature of between about 163°C and about 260°C, preferably between about 193°C and about 232°C.

It then emits useful heat, e.g. for home cooking purposes, as schematically shown in fig. 1, where it emits heat in the heat exchange loop 101 to the fractionation stripper 20 and in the heat exchange loop 102 to the CC^ stripping section 21 in the stripping absorber 22. Such applications with regard to the current in the line 13 can take place either in series or in parallel, but will usually most advantageously be in series with the current in the landing 13 so that first high-temperature heaters are emitted to the bottom of the section 21. After the current in the line 13 finally is cooled by such utilization, significant amounts of water are condensed and some methanol products from the partial oxidation reaction so that by their latent heat they increase the sensible heat that is recovered. As will be further seen in the following, practically all released heat from the partial oxidation reaction is thereby usefully recovered. Furthermore, it can be seen that each heat increment is recovered at practically as high a temperature as possible so that it contributes significantly to the overall high thermal efficiency of the process.

Under these conditions of partial oxidation reaction with a high ratio of methane to oxygen and relatively high pressure, the overall main reaction is:

As described in the previously mentioned application, methanol is produced under preferred conditions with high conversion selectivity from methane, together with small amounts of formaldehyde, carbon monoxide and water vapor and very small amounts of carbon dioxide, hydrogen, formic acid, hydrogen peroxide, dimethyl ether, dimethyl peroxide, etc. The presence of such smaller components in the re-circulation ion gas in line 18 will significantly promote the initiation of the free radical reactions involved in the partial oxidation, as they are all (with the exception of carbon dioxide and water) more reactive in this environment than the methane itself.

In the previously mentioned washing operation of the above-mentioned smaller components, only formic acid, hydrogen peroxide and formaldehyde will dissolve to a particular extent in the washing water in the line 14 and, as they are less volatile than methanol in the presence of water, they will remain in the washing water and will finally contribute to the formation of synthesis gas, as described in more detail below. Smaller products that are not washed out will mostly be recycled with the net off-gas flow in line 18 from container 10

with its content of minor constituents also providing their contribution to synthesis gas formation.

Methanol of substantial purity is thus not only produced as an end product from this partial oxidation plant. Although methanol can very advantageously be produced in the rest of the process without using this partial oxidation plant, its inclusion in combination with the rest of the process has numerous advantages and this is related to the invention.

In terms of the plant's heat output, it works in a way like an oven but better than this. It thus develops heat directly in the stream that is treated rather than in another stream (combustion stream) which is finally discarded so that potential heat and product are thereby lost. Second, even latent heat of condensation is easily recovered, as referred to above, which apparently can never be easily or practically accomplished with the combustion products from a furnace. Third and most important, it is a "furnace" in which methanol is produced as a recovered product rather than carbon dioxide which can very rarely be economically recovered or utilized profitably from combustion gases.

Fourthly, it provides an equilibrium in the overall methane/oxygen ratio in the whole process so that the overall reaction in the whole process is mainly CH^+ 1/ 20^ = CH^OH so that it provides the highest possible efficient utilization of both feed carbon - and hydrogen atoms. The reason why it provides this equilibrium is that the relevant partial oxidation reaction itself consumes more oxygen than the indicated half mole per moles of reacted methane while the subsequent production of synthesis gas, as taught by the invention requires less.

Fifthly, to the extent that the minor products produced by the partial oxidation reaction are carbon monoxide and/or hydrogen, direct reforming has already been carried out so that the burden for subsequent reforming requirements is thereby eased. In addition, any formaldehyde etc. which is formed and which finally passes to reforming either in aqueous solution or in the gaseous feed thereto, where it is very easily reformed without the essential need for oxygen, steam or heat input for an endothermic reaction, since the reaction: CH2° = co + H2 is almost thermo- neutral.

Finally, this "furnace" with internal heat development will preheat the exhaust gases 23 from there which are the supply to the reforming section. The preheating that can easily be added will be more than sufficient for the next process step and allows some of the heat content therein to be recovered from there in the form of high-pressure steam (preferably 63.3 kg/cm 2 or more) so that heat is not only delivered where it is required for the process itself but that heat is also provided for the production of motive power required therein.

Exhaust gas in line 2 3 is preferably obtained from the delivery side (in line 18) by the recirculation fan 16 without admixture of oxygen. As detailed in the aforementioned co-pending application, this gas passes via a line (not shown)

to one or more separate sections of the chamber device 24, preferably passes inwards through an insulated section of the matrix 19 into a separate collecting inner chamber device (not shown) which communicates with the outlet from the container shown as line 23. The matrix 19 can and preferably preheats as well this exhaust gas to a temperature higher than the average preheating given off by the matrix 19 to the total gas that is heated thereby, by arranging the said separate section or the said sections of the chamber device 24 at one or more respective points with respect to the direction of rotation of the matrix so that this has just been heated by relatively hot, flowing gas immediately before it is brought into contact with the aforementioned exhaust gas which is to be preheated.

This discharge in the line 23 then preferably passes through the boiler 25 and emits heat there and then passes at least the main part of it through the line 26 to the water side 27 of the methanol converter container 28. The construction and operation of this converter 28 is described in detail in the the following with reference to fig. 4.

The gas from the line 26 mixed with hot process water causes some of this water to evaporate into the gas and the mixture passes up through pipes, is detached from the water after the passage and the gases are now united after being saturated with water (and any other components in the water) with any desired flow through the valve 29, and this gaseous mixture in the line 30 becomes the main supply to a reforming container 31. By regulating the flow through the valve 29, any desired water vapor content in the gases in the line 30 can be achieved up to saturation of the total gas therein at the temperature of the water at the detachment point.

If additional water content is required in the total gas supplied to the reforming container 31, recirculation gas in the line 32 can also be introduced to the water side 27 in the container 28 and also absorb its contribution of water vapour. Such a modification, although feasible and within the scope of the invention, will not usually be necessary in this preferred embodiment for the production of synthesis gas and methanol. The reason why a relatively very low ratio between water and carbon can be allowed in this embodiment compared to conventional steam reforming is further explained in detail in the following. It is sufficient at this point to note that this lower process steam requirement is a significant contributing factor to the very high feed and fuel efficiency (over approximately 93%) of this preferred embodiment for the production of synthesis gas and methanol. In conventional steam reforming, where steam must be produced and mixed with process gases beyond the amount or pressure of steam that can be developed from the heat release from the methanol production reaction, thus generally requires compression of the relatively large volume of synthesis gas and unnecessarily wasted condensation of water from the reformed gases at lower temperatures than for its development, due to the increased volume and the mixing itself.

In the container 31, preferred gas from the line 30 comprising mainly methane and steam, recirculation gas from the line 32, mainly consisting of carbon dioxide and oxygen from the line 33 are each preheated separately and are then combined during which a series of reactions take place which are initially highly exothermic, as the oxygen actually "flows" in an excess of hydrocarbons, as well as steam and Co2, and the steam and Co2 will then first rapidly and then more slowly combine endothermicly with remaining hydrocarbons, e.g. according to the following reactions:

The detailed construction and operation of the container

31 is described below with reference to the figures

2 and 3. As can be seen from the following, on an aggregate basis inside the vessel 31, little if any of the second highly endothermic above-mentioned reaction (depending on C0 2 ~ production in the partial oxidation ion in the reactor 17) actually takes place this reaction can even be slightly reversed). As can be seen from what follows, the largest part of at least the CO^ required in the relevant complex equilibria, such as e.g. the equilibrium for the water-gas reaction: CO + H20 = C02 + E^, thus supplied to the reforming reaction environment not by production therein but by means of recirculation. Here, too, this recycling operation is a contributing factor to the high yield, as in the production of synthesis gas, C02 must be considered a pollutant rather than a desired component. E.g. in the production of methanol from synthesis gas, the production of methanol by the reaction CO + 3H2 = CH^OH + H20 is demanding with regard to hydrogen and produces unwanted water in the methanol product. In addition, high C02 contents in the gas above the methanol conversion catalyst will likely lead to inhibition of the desired reaction: CO + 2H2 = CH3OH. Furthermore, a significant loss of C02 from the overall process means that e.g. in the purge gas from the methanol converter, mean a loss of both carbon and oxygen from the process.

After the above-mentioned reforming reactions have taken place homogeneously in the container 31 to a significant extent, the temperature, due to the endothermic nature of these reactions, has fallen to the point (below approximately 1260°C) at which the rates for these homogeneous reactions become relatively low, and then the reforming reactions are completed over a so-called secondary reforming catalyst, as these catalysts are well known in the field and can be easily obtained commercially (for operation, it is said up to a maximum temperature of 1300°C). The exhaust gases leave this secondary catalyst preferably at a temperature of 1038-1149°C, pass back into heat exchange with the incoming gases, give off sensible heat thereto and exit at a temperature of preferably 288 to 482°C, more preferably 299-343°C . At these relatively low temperatures, the synthesis gases have little or no noticeable tendency (even over metals) to return to equilibrium states corresponding to this temperature, as these equilibria would tend to deposit carbon by the aforementioned Boudouard reaction: 2CO = C02+ C , or reform methane by the opposite of the above-mentioned reforming reactions.

It should be pointed out that this process for the production of synthesis gas is not only very efficient, but in a unique way it solves the difficult problems associated with it.

Firstly, as mentioned above, when hot synthesis gas is cooled, the carbon monoxide in it will have a strong tendency to deposit carbon according to the Boudouard reaction: 2CO = CC>2+ C. This carbon formation makes it difficult for the existing (autothermal) processes that use oxygen to form synthesis gas at a relatively high temperature. The invention avoids these problems in three ways: 1) Below about 927°C, in which the preferred embodiment equilibrium consistent with the Boudouard reaction will become increasingly stronger as the temperature is lowered tends to

to form carbon, the synthesis takes place in the middle of the glassy ceramic fiber heat exchange matrix which is then cooled very quickly. Now, the Boudouard reaction is strongly catalyzed by metal surfaces, especially Group VIII metals, transition-

the metals, especially including the common high-temperature structural metals iron and nickel. Heat cannot thus be recovered from the hot synthesis gas using heat transfer through metal surfaces. On the other hand, non-reducible and non-oxidizable ceramic oxides are almost non-catalytic towards this reaction but they are of course poor heat conductors and consequently cannot be particularly effective for the transfer of heat by means of recuperator devices but only by the regenerative measures in accordance with the invention. 2) The conventional steam reforming process requires external heating through metal tube walls and only avoids carbon deposition during synthesis gas formation by using relatively large amounts of steam, and it can recover heat from the hot synthesis gas only by maintaining the heat recovery metal tube walls below about 343°C ( using boiling water), where the Boudouard reaction is now limited to a very low reaction rate even in the presence of metals. The process according to the present invention, in contrast, does not form carbon monoxide during the heating phase and consequently the high steam demand is avoided, and is able to recover heat at much higher temperatures and emits the synthesis gas itself, for further cooling in the presence of metal walls, below a temperature of 482°C, where the reaction is relatively slow, preferably below about 34 3°C, where the Boudouard reaction is no longer a particular problem. 3) The range of cooling from approximately 927°C to approximately 482°C, or preferably approximately 927°C to 343°C is then the critical temperature range for carbon deposition and it must be pointed out that only a very small deposition rate e.g. which could be expected from a "non-catalytic" surface within a relatively short time, possibly within a few minutes, clog a regenerative heat exchanger. To overcome this problem with safety, several of the elements of the preferred embodiment cooperate. Firstly, the matrix of the preferred regenerator consists not only of "inert materials", but also of fibrous ones and will consequently filter out such carbon as is formed near the point of formation so that possible downstream problems are then avoided. Next, there will only be a deposition amount corresponding to less than about 12 seconds of matrix contact time with the synthesis gas by the matrix preferably rotating at about 5 OPM or more and by it preferably being regenerated at each revolution. Finally, in the preferred embodiment, not only the hydrocarbon-containing feed gases and steam (and CC^) will be preheated but also the oxygen of the same matrix. Any deposited carbon will thus be quickly and completely burned by this mainly pure compressed oxygen. No clogging problems will thus occur and all problems associated with the instability of the synthesis gas and the effective recovery of heat from it will be solved without the use of larger quantities of steam and then only the quantity required based on equilibrium considerations at the formation temperature, which will always be above about 816°C and preferably above about 1038°C where the synthesis gas has its highest stability.

Then the exhaust gases in the line 34 are preferably cooled by heat exchange in several stages with water that boils at one or more pressures suitable for the process and are also cooled by preheating such water streams suitable for the process. All these steps are symbolized by the heat exchanger 35 and by the heat exchange loop 94 which boils mainly water at the bottom of the fractionating stripper 20 using heat from these exhaust gases. During much of the latter part of this cooling, water condenses from the gases and separates out at one or more suitable points. This separation is symbolized by the exhaust gases in the line 36 passing through the separator 37 where the condensate is removed via the line 38. After final cooling with cooling water or other coolant in the heat exchanger 39, preferably below approximately 4 3°C where the now relatively dry gases in the line 40 enter in the separator-absorber container 41 • In the bottom section of this container 41, the gases are brought in countercurrent contact with an absorbent for C02 in an amount dg in sufficient steps in the absorber so that most of the CC<2 is preferentially absorbed in the aforementioned absorbent. In the case of this preferred embodiment for the production of methanol, it is preferred that the absorbent is the raw methanol itself, so that, in addition to other advantages, possible contamination of the product is avoided. the methanol with absorbent. This raw methanol produced in the methanol co-converter is also almost completely sulfur-free and will also tend to absorb any sulfur compounds including r^S and COS

both of which are more soluble in methanol than CO 2 , so that there will be a tendency to minimize the deactivation of these highly sulfur-sensitive modern so-called "low temperature" methanol conversion catalysts.

After absorption, the prepared synthesis gas passes through

in this preferred embodiment up through the riser 42 and joins the converter recirculation gases from line 43. These gases then pass through the circulation fan 44. Phineas Davies, et al in US Patent 3,326,956 teaches that a small amount of C02 are desirable in gases where methanol converter catalysts. In part, because water and C02 can be interconverted using the water-gas equilibrium reaction given above,

and that the "low temperature" methanol converter catalysts are active as a catalyst for this reaction, a small, controlled amount of (relatively pure) water condensate is shown in fig. 1 sprayed into line 45 via line 46. It is found that such a method is at least as effective as the use of C0 2 in preventing harmful reduction of the catalyst activity by not introducing larger amounts of C0 2 . If desired, it is of course within the scope of the invention either to reduce the efficiency of the absorption step or to use a small controlled bypass (not shown) of exhaust gases around the absorber. Such a by-pass before cooling the exhaust gases for direct introduction to the inlet of the methanol converter via line 4 7 and the valve 48 will thus not only supply C02 but also

water to the converter.

After preheating in the heat exchanger 49, the supply to the methanol converter 28 in the line 50 enters the converter, is partially converted to methanol as more fully described with reference to fig. 4 and leaves the converter 28 via line 51. Most of these gases are cooled in the heat exchanger 49, while preferably a small stream of these converter exhaust gases is used directly as a relatively hot stripping gas when introduced into the bottom part of the CC^ stripping section 21 in stripping -absorber 22 via line 52. Such direct use of the converter exhaust gases as stripping gas is very advantageous as explained in more detail below.

After cooling in the heat exchanger 49, the converter exhaust gases in the line 53 are further cooled and methanol (and comparatively smaller amounts of contaminants such as water, dimethyl ether and ethanol) are condensed in the condenser cooler 54, pass into the separator 55 via the line 43 and liquid from the line 43 is separated from the converter recirculation gases in the separator 55.

It is now seen that with the present invention, in contrast to conventional synthesis gas processes, no large and expensive synthesis gas compressors are required (the compressors that are necessary are simple fans to compensate for pressure drops in the recirculation streams). The reforming reaction, despite the harmful effect of the pressure on it, can be carried out in the same pressure range as the conversion of synthesis gas, e.g. to methanol. This is partly due to the autothermal nature of the reforming reaction, where it is not necessary to transfer heat through temperature-limited metal tube walls from a combustion zone to an endothermic reforming zone. Other fully or partially autothermal reforming processes

are known, but these have not even made the high efficiency levels of the present invention approachable. These reactions can of course, as in all reforming reactions,

tions, are completely carried out, as some of the hydrocarbons to be reformed, mainly methane, remain. In the present invention, under preferred conditions, this unreformed methane in the relatively dry synthesis gas in line 40 is kept below approximately 3 mol percent of the dry synthesis gas product from the reforming step and preferably below approximately 2 mol percent. In the context of conversion reactions, such as e.g. methanol formation, where liquid is formed and removed and the gases are recycled, this methane content, as well as the content of other inert constituents, will "build up" and if it is not removed, it will build up to a degree to detrimentally affect the rate of conversion by drastic and reduce the partial pressure of the reactants CO and . These latter partial pressures must remain at significant values because the equilibrium for the reaction CO + 2E^= CH-jOH varies in favor of methanol formation directly as the square of the hydrogen partial pressure and in relation to the carbon monoxide partial pressure and the conversion rates vary accordingly. Means for controlling the amount of internal components without excessive yield loss are thus a very important part of the present invention.

Such internal components as argon, which enters the system together with the oxygen, and nitrogen, which can enter both together with the oxygen and in the natural gas, must of course be removed by means of some kind of flushing devices from the methanol-converter-recirculation gas to outside the system. On the other hand, in the case of ammonia production from hydrogen produced from the synthesis gas by means of well-known steps, nitrogen is of course not an inert component.

In the case of methanol synthesis from the synthesis gas, another complication is that it is generally found that while a ratio of two to one between H2 and CO is required for methanol formation, a ratio between H2 and CO over the conversion catalyst that is quite substantially higher is desirable.

If the carbon monoxide concentration is too high in relation to hydrogen, this will tend to inhibit the rate of methanol formation, apparently by preferential absorption of CO on the catalyst due to the exclusion of sufficient hydrogen absorption. At the same time, as previously mentioned, the present process produces an equilibrium with a hydrogen/carbon ratio just suitable for methanol production, i.e. mainly 2 to 1.

Further objects of the invention are therefore simultaneous

to remove internal components from the methanol conversion circuit to a sufficient extent so that their concentration above the catalyst is not too great (i.e. methane, argon, nitrogen and CO2 total content less than 20% when the gases leave the catalyst), further to remove a quantity of combustible gases therefrom together with the inert constituents to outside the process and which is not greater than the amount that can be effectively used by efficient devices to generate the compressed oxygen supply (in conjunction with the steam available for power generation from the process), further allowing a near stoichiometric establishing the ratio of hydrogen to carbon monoxide, allowing a more than stoichiometric ratio over the methanol conversion catalyst, and yet at the same time not losing excessive relative amounts of hydrogen due to its higher than stoichiometric content in the converter circuit.

The simultaneous solution for these requirements, in accordance with aspects according to the invention, is achieved by several cooperating means. Firstly, the CO2 in the synthesis gas is absorbed and recycled, as mentioned. This keeps the C02 proportion of the inert components low in the conversion circuit. It also results in the production of very little water in the methanol product so that subsequent distillation requirements are thereby lowered, reducing parasitic heat losses in the subsequent use of raw methanol in which it evaporates, such as e.g. the previously mentioned MTG process and considerably increases the dissolution capacity of the crude methanol used for absorption of various unwanted gaseous components.

Next, the methane portion, which is much more difficult to absorb from the synthesis gas, similarly to CO, where crude methanol is used as absorbent, as economically prohibitive amounts will be required, is controlled by the inert components by a combination of devices. By using the converter exhaust gas in line 52 as a stripping medium for stripping recycled CO2, methane in this stream is returned to the reforming reactions in container 31 where it is usefully reformed. In the preferred embodiment, this return is over. 50% of the quantity that enters the conversion circuit.

Methane and other gases are absorbed in the raw methanol product which is separated from the gas in the converter circuit in the separator 55. In the preferred embodiment, this absorption is promoted by the addition of recycling methanol to the converter exhaust gases when these are cooled and condensed in the condenser cooler 54. In fig. . 1, this is schematically shown as the line 56. This method, over the large surface in the condenser 54, substantially saturates the methanol with dissolved gases at relatively little additional capital and operating costs. The methanol from the separator is then introduced via the line 57 to the bottom evaporation section 58 in the stripper absorber 22. In the preferred embodiment, this evaporation section is operated at a pressure just high enough so that the gaseous outflow from the stripper absorber 22 at the line 59 serves as fuel gas for a modern, relatively efficient gas turbine, i.e. approximately 17 ata. In the pre-evaporation section 58, a significant proportion of the dissolved gases nitrogen, argon, hydrogen, carbon monoxide and methane (but little CO^ or heavier components) will evaporate by depressurization and finally be used as fuel gas, as already mentioned. The total amount of combustible gases provided in this way to fuel gas is controlled by the amount of recycle methanol added to this circuit at line 56. However, since a large proportion of these combustible gases are CO and H_ desirable for methanol conversion rather than as fuel gas (i.e. that one have made one "process investment" with them) it may be desirable, depending primarily on the nitrogen and argon contents of the conversion circuit gas, to depressurize less fuel gas at this point than is required for fuel. In these cases, it would be advantageous to supplement this evaporation gas directly by separately introducing natural gas, as shown by line 60. This natural gas can be introduced either above the liquid in! .section 58, shown by line 61 and valve 62, or if the argon and/or nitrogen contents would like to build up to too high values in the conversion circuit, and natural gas is fed into this liquid, schematically shown by line 63 and valve 64. This latter method will have the effect that nitrogen and argon are reduced in the conversion circuit and the methane content is raised, although this is partially removed by the equilibrium with gas in the cooler 54 and return to the reforming vessel 31 (via line 52) and consumption there.

The aforementioned removal blow-through operation will be most advantageous under conditions of relatively low working pressure and relatively high levels of inert constituents argon/nitrogen.

It is clear and of course that these measures for

removing inert constituents gives working flexibility to the process, especially in the case of process difficulties, which, due to the very high yield intended, is itself an intimately connected process with regard to several of its aspects. The advantageous reason for adding natural gas fuel through line 60 rather than directly to the gas turbine will be apparent from what follows with reference to the operation of the stipper absorber 22.

It is now seen that the devices are used to control the content v of inert components in the converter circuit, due to the lower solubility of hydrogen in methanol than the other gases present. present, acts to conserve hydrogen in the system so that it can build up to a ratio to carbon monoxide that is greater than the production ratio in the reforming container 31. In the preferred embodiment, the ratio E^/CO in the reformed exhaust gas in the cladding 34 is thus approximately 1.95, while the ratio at the output of the methanol converter in line 51 is approximately 2.68, and while at the same time the ratio in the fuel gas is significantly less than 2.

As shown in fig. 1, it is preferred that the source of the methanol used as feed to this pressure relief evaporation/bubbling is methanol pumped back up to the conversion pressure from the pressure relief evaporation section 58 via line 65 and pump 66. Only the amount of net produced crude methanol enters the CC ^-absorber section 103 in separator-absorber vessel 41 via line 67. Other arrangements are possible but this would tend to isolate the pressure relief evaporation/bubbling-through methanol with its relatively high methane content from the methanol system used to absorb CO^. In this way, little if any methane is fed back into the synthesis preparation gas at equilibrium with the CC 2 -poor methanol which enters the container 41 via the line 68'. It is desirable to reduce f^S and COS levels as much as possible with these currents. It is therefore preferred for line 67, rather than this leading to mixing with the CC^ -poor methanol in line 68, that it is fed directly to the top of the absorption section, while the methanol in line 68 is introduced for absorption several steps below the stage for line 67. in this way f^S and COS are finally absorbed by essentially S-free methanol as produced in the methanol converter. By minimizing sulfur in the gases of the methanol converter (usually to well below 0.1 parts per million) this will of course extend the catalyst life.

C02~rich methanol leaves container 41 via line 6 9

and at least a significant portion of it is preheated in the heat exchanger 70. It may be desirable (not shown) to leave a relatively small portion cooler than the main portion

and introduce it at a point higher in the stripper-absorber 22 than the main flow, to act as a deflector,

so that methanol is thereby condensed from the stripping gas in section 21. Such a method is well known. In addition, it is preferred that the stripping gas from section 21 is washed with process water condensate, e.g. from the line 38, in the washing part 71 to remove methanol from recirculation gas in the line 32, which is then circulated by means of the fan 72. In order to reduce energy requirements for pumping and compression to a minimum, it is preferred that the CC^ stripper 21 is operated in the same pressure range as the methanol converter 28, as determined by the pressure drop in the line 52. Alternatively, the fan 72 can be inserted in other places, e.g. in line 52 where the molar current is smaller than in line 32

but the temperature is higher.

The flow in line 73 from the bottom of washing section 71, comprising methanol and water, is depressurized in absorber section 74, together with the wash-bottom fractions in line 75. Stripping gas from there, together with gas from riser 76, is washed with process water condensate, as from line 38, for the recovery of methanol from there and exits via line 59 as fuel gas. The bottom fractions from the absorber section 74 pass via the line 77 and the valve 78 to the fractionation stripper 20 where the methanol content is stripped out, as described below.

The CG^-poor methanol from the bottom of the CC^-drop section

21 passes via line 79 and pump 80 through heat exchanger 70 and water cooler 81. It is preferred that the net methanol preparation remaining in this system is removed from the hot CC^-poor methanol in line 79 through line 82 and valve 83 into the methanol stripping section 84. This hot methanol (usually between about 160 and about 204°C) contains relatively little dissolved gases and is at such a high temperature that stripping impurities from there, such as e.g. dimethyl ether and low boiling impurities are relieved using the gases from the pressure relief evaporation section 58. In this way the second (low boiling impurities) methanol purification tower, which is usually required, is rendered unnecessary along with its energy requirements. The stripping in question here thus uses the "free" availability of the cleaning gas (which is almost free of dimethyl ether etc.). The heat content of the methanol heats the said stripping gas and the resulting sensible heat thereof is of course recovered by the combustion thereof, while the remaining heat content of the liquid methanol leaving the stripping section 84 via the line 85 and the valve 86 is useful in its distillation in the fractionating stripper 20 .

In the upper section of fractionating stripper 20, methanol product is withdrawn as top product via line 86, condenser 88, separator 89, pump 90 and line 91 while methanol reflux is returned to the tower via line 92. Dissolved gases, with relatively very little quantity due to the mentioned stripping operation with the methanol and the low solubility of these substances in the aqueous mixture in the absorber section 74, from which the bottom product enters the fractionation stripper 20 via the line 77, comes out of the system via the line 93. These emission gases can be used as fuel for a burner operated near atmospheric pressure which partially reheats the combustion gases from the gas turbine (not shown) before; these gases give off sensible heat to an excess heat boiler (not shown) or they can be recomplied to gas turbine fuel gas pressure and introduced to this fuel gas -system at an appropriate point.

In the lower stripping section of the fractionation stripper 20, methanol is stripped primarily from water that has been used to wash methanol from the various gaseous streams in the process. It is an important element of the invention to realize that this stripping is achieved at least to a significant extent "for free" in the sense that in order to achieve a given production of dry product methanol, it entails a certain boiling, and that an addition up to a certain substantially of the amount of methanol partly as an aqueous mixture with a relatively low concentration does not require any major boiling, i.e. the stripping steam is already available. This is a further reason why: (1) the addition of the direct partial oxidation system, which recovers methanol by aqueous washing,

is so advantageous, and (2) the use of methanol as an absorbent, which is partially recovered by water washing, is so advantageous.

Ethanol, which can be produced as a minor contaminant both by the partial oxidation (primarily from ethane) and in the catalytic conversion systems, is preferably removed from the product methanol from the fractionating stripper 20 s on the vapor side outlet 95, at an intermediate point in the tower where the liquid phase is primarily aqueous and the methanol content is relatively low. It can be treated in the same way as the exhaust gases from line 93 so that the heat value is utilised.

Formaldehyde and other contaminants, which in the aqueous (or methanolic) liquid phase are less volatile than methanol, will tend to accumulate in the bottom fractions from the fractionating stripper 20 and a part will be recycled to the partial oxidation washing section in the container 10 via the pump 96

and the line 14. In the washing section of the container 10

these pollutants will establish themselves in equilibrium with the gas, and function as "recirculation initiators" for the partial oxidation reaction, as previously explained. Excess wash water in this system is passed via line 97 through lines not shown, through fcr heater(s), symbolized by heat exchanger 35, and enters the water side 98 of methanol converter 28 as shown via line 99. This process water will be something on the sour side because of the presence of a possible relatively small amount

of organic acids such as formic acid, and it is necessary that this circuit be of a corrosion-resistant alloy. As explained in detail below, with reference to fig. 4, this water and its volatile organic contaminants will evaporate into the gases which are thereby saturated and the organic contaminants therein will be destroyed at high temperature in the reforming container 31 as previously mentioned. The container 31 thus acts as a cleaning device and the water condensed from

the gases from there will not contain any of the dissew pollutants. Surplus water from the total process can thus be taken from there as relatively very clean water, with the exception of dissolved gases such as e.g.

CO2 that can be recovered in an aeration device (not shown).

j

REGENERATIVE HEAT EXCHANGER

With reference to fig. 2, this shows a schematic cross-section of the preferred rotary heat exchanger 109

in the container 31, in which the inputs to the container 31 comprise the hydrocarbon-steam mixture in the line 30 in fig. 1, recirculation gas and oxygen, which are already preheated, while the product gas from the container 31 is cooled. This unit preferably comprises a rotating matrix 110 (preferably comprising heat-resistant ceramic fibers prepared and sectioned in accordance with the previously mentioned concurrent application), oxygen inlet 111, mixture inlets 112 and 113, recirculation gas inlets 114 and 115, and steam inlets 116, 117, 118- , 119 and 120. It is thus preferred (apart from the oxygen) that the unit is roughly symmetrical around the center line to the extent that the slightly higher pressures in the inlet gases compared to the outlet gases are balanced on each side. Although separate sealing shoes 121-127 may be used, it is preferred to combine each pair of inlets together into a single unit with flow-through channels and openings therein for the pre-

different currents, and correspondingly with shoes 128 to 132.

Gas outlets 133 and 134 from the container are also shown

31, preheated oxygen collection channel 135 and preheated hydrocarbon-containing collector chamber 136, together with sealing shoes 137-141. The matrix 110 is sectioned into an appropriate number of compartments to prevent cross-contamination of gases within the matrix. It should be particularly noted that these compartments, as each rotates to the relevant position, must be almost completely purged of combustible gases before oxygen is introduced therein and vice versa, using steam streams as shown. It is also important that these steam streams are kept at a higher pressure than the other streams so that leakage past the shoes mainly occurs only as steam and not as any of the other gases. In the case of the inner shoes, this additional steam pressure can be supplied from the compartment from the outside, with suitable width and positioning of the inner shoes 137-141, used to arrange leakage of steam rather than other gases therefrom, or additional steam connections to these shoes and distribution tracks therein ( not shown) can be used.

The chamber device 136 communicates by means of channel devices to the inlet of the reforming catalyst layer 100 shown in fig. 1. The outlet from this layer communicates by means of channel devices to chamber devices 142 and 143, and the reformed gas therein then passes out through the matrix 110, is collected in the chamber device 144

g 145, and leaves the container through nozzle outlets 133 and 134.

Fig. 3 illustrates a preferred variant of the heat exchanger 109. The cross-section is taken along the line 3-3, through the plane of the oxygen inlet nozzle 111. It comprises rotating cylindrical shell 146 with holes 147, the support grain 148 (attached to a driven axis, which does not is shown, at the center line), strut 149 which supports the plate 150, the matrix 110,

the ceramic core 151, plate 152 and insulation 164.

The enclosing container 15 3 (which may be insulated

internally and/or externally, not shown), has a large nozzle 154 to which cover 155 is attached, and which carries steam inlet nozzle 156 and oxygen nozzle 111 (as well as nozzles for the other inlet gases not seen in this view). The sealing shoe 157 for all inlet gases (supported by means of devices not shown) is made self-positioning by means of the gas contained in cylindrical bellows 15 8 and 159 which exert pressure on the shoe 157, this pressure being counteracted by injection of an inert current with higher pressure, preferably steam, with a relatively small and definite stroke at several points between the shoe 157 and the rotating shell 146 (illustrated by means of the tube 166 and the mouth 167) so that thereby it maintains a relatively fixed separation of the order of tenths of a millimeter between them.

During operation in accordance with the diagram shown in fig. 3, oxygen passes through the nozzle 111, through the space surrounded by the cylindrical waves 159, the hole 160 in the shoe 157 and through hole 147, is distributed across the outer surface of the matrix 110 by means of the chamber device 161 and passes through the matrix 110, is collected by the chamber device 162 and passes through the renoicrr hole 163 in the core 151 and is collected by means of the fixed oxygen collector 135 which is terminated by means of a cylindrical section containing the hole 156. Outside this hole, oxygen will encounter a (combustible) mixture comprising: a hydrocarbon gas and steam. and will therefore become an intense flame directed into the open space below. In and after this flame, reforming of the hydrocarbon-containing gases will take place as described with reference to fig. 1.

It is pointed out again that at all points of possible leakage along this oxygen path devices must be arranged so that steam leaks into the path rather than oxygen or that other flammable gases can have some opportunity to mix with oxygen. It is also necessary, in the case of some failure of the oxygen flow, that the oxygen path is completely flushed out, so that flammable gases do not return to this path. This is easily achieved by connecting a high-pressure steam line to the oxygen line

(using a self-acting pressure reducing valve)

immediately after a check valve at the outlet from the oxygen compressor. The pressure reducing valve can be set to just let out at least the bulk of the steam flowing at the working conditions, i.e. preferably set at something below the working oxygen pressure at the point of connection (of course the control connections for the process must be such that no operations can start before the steam has at least the pressure required at this point).

CATALYTIC REACTOR WITH CONTROLLABLE TEMPERATURE

With reference to fig. 4 shows a side sectional view of the preferred embodiment of the temperature controlled catalytic reactor. For the conversion of synthesis gas and other gases by catalytic means in which the reaction catalyzed is exothermic, this reactor has many advantages. First, it provides a large-scale means of maintaining exothermic catalytic reactions at essentially isothermal conditions. With such reactions this is important since (1) the rates of these reactions increase sharply with increased temperature, while equilibrium towards desired products becomes less favorable with increased temperature (2) with too high a temperature active catalysts will tend to "sinter" and become de-activated, and (3) recovered heat is higher the higher the temperature at which it is recovered (eg, that the higher the vapor pressure and energy available therefrom, the higher the temperature).

Second, it provides an advantageous means on a large scale to "program" and control an optimal temperature profile. This is important in accordance with (1) above, when the reaction over the catalyst is far from equilibrium (i.e. close to the gas supply inlet) where the optimal temperature, by assessing the reaction rate against equilibrium considerations, is relatively high,

while near the outlet, where the product concentration is close to a maximum, the optimum temperature is relatively low.

Third, it provides the means to

stabilize exothermic reactions. This is important since (1) if the inlet temperatures are not high enough,

or if the content of inert constituents is too high,

the reaction will not "flare up", i.e. self-accelerate on the catalyst surface to the point where it is self-sustaining with respect to temperature, and (2) the valuable products are often an unstable intermediate relative to possible exothermic reactions over a catalyst . E.g. in the desired production of methanol, this product is an intermediate both in the sense of hydrogenation of carbon monoxide and release of heat to the production instead of methane, in accordance with the reaction CO + 2H2 = CH4 + H20 + 49 Kcal. If the catalyst is overheated and i. enters a dangerous temperature range, the latter reaction will self-accelerate, the temperature rises rapidly and then not only is the catalyst destroyed, but in extreme cases the reactor itself can be irreparably damaged.

Other examples of such possible harmful behavior include the oxidation of benzene or butane to the desired maleic anhydride when the oxidation, and heat release, may tend to go much further, to the undesired CO 2 and H 2 O.

Fourth, it provides the means for very large capacity catalytic converters. This is important as many of the industrially important reactions, including many reactions that use synthesis gas as a starting product, such as e.g. production of methanol and ammonia, is carried out under relatively high pressure, and it is economically and apparatus-wise important that all necessary features can be enclosed in a container with a relatively small diameter in relation to the production capacity. Otherwise, the production capacities will be limited to a

respectively low level due to the thickness of the apparatus walls which are greater than what can be produced by the available technique for pressure vessel production. In the case of both methanol production and ammonia production, this has proven to be a problem that limits capacities for plants on the very large scales,

and expensive and cumbersome devices have become necessary to circumvent this diameter size problem (eg the new long horizontal Pullmann-Kellogg ammonia converter).

Finally, it provides means for easy rapid catalyst replacement. This is important because that limited catalyst life is usually 2 to 5 years.

These and other purposes are achieved as described below. The reactor preferably comprises the container 200, gas inlet/outlet nozzles 201 and 202, preferably conical tube plates 203

and 204, pipes 205, 206 and 207, catalyst chamber 208 surrounded radially by cylindrical catalyst screens 209 and 210, catalyst filling nozzle 211, nozzle cover 218 and main pipe 212, as well as inner catalyst screen cover 213, and lifting handle 214, liquid chamber 215, preferably water, and liquid/ gas separator compartment 216, preferably with guide plates 217. I

in addition, the reactor comprises cooling liquid inlet 219, which preferably distributes said liquid around the circumference by means of the guide plate 222, downflow channels represented by 220, preferably arranged as shown" ie in communication with the lower part of room 216, respectively. in the surrounding chamber space 221. Downflow channels 220 can also be placed radially inside the screen 209.

Gas to be converted catalytically can pass along

its path through the container in any direction, but it is generally preferred that the gas enters through nozzle 201, passes circumferentially around the reactor in chamber 221, enters screen 209, passes radially around the tubes and through the catalyst layer in chamber 208 and screen 210 and out through the nozzle 202. Spent catalyst is usually emptied by simply removing an elbow (not shown) connected to and below the nozzle 202, removing the nozzle cover 218 by using the lift handle 214 and lifting the inner screen 210. The catalyst will flow out easily, aided by the inclination of the preferred conical tube plate 204, (possibly using a vibrator).

For filling the catalyst, the screen 210 is dropped back and centered in place, preferably by means of a circumferential notch 227, and the catalyst is poured in through the nozzle 211. Relatively little breakage of the catalyst will usually occur during the filling operation, as first the free fall of the catalyst is stopped by the screen cover 213 and then due to that the pipes will tend to stabilize a very high angle of descent during the main part of the catalyst filling operation so that the free fall distances are further reduced. Catalyst fill above screen 210 acts as a reservoir to replace gradual subsidence due to compaction, etc., so that any tendency for relatively untreated gases to pass beyond the catalyst is prevented.

Another important advantage of having conical tube sheets, apart from their much greater mechanical strength against pressure differences compared to flat tube sheets, is that the radial gas flow is thereby kept at a more uniform velocity, the extra height towards the center line serving to compensate for reduced circumferential flow cross section.

This feature, together with the fact that the flow path is relatively short, means that the pressure drop through the layer is minimal.

A very important aspect of the invention is that the heat release which takes place inside the catalyst is transferred to and absorbed by liquid which is kept at an almost uniform temperature, preferably by its evaporation. Further,

in the case of methanol production, using modern active so-called "low-temperature catalysts", steam produced thereby can be used directly in the reforming to form the synthesis gas without the need to then compress the synthesis gas to subsequent converter-reactor pressure. So far, such a thing has not been practical. It is achieved in the methanol conversion reactor in accordance with the invention, by injecting the gas to be reformed into the water space 215 through nozzles 22 3, 224 and 225. Water vapor is thereby evaporated directly into the gas during its passage up through the heated pipes. It is thus practical to drive the catalyst towards the gas outlet from there, at the equilibrium-favorable temperature as low as e.g. 260°C, while the water evaporates at about 24 3°C at a total pressure of

more than 60 ata, while the vapor pressure of water at 243°C is only about 35 ata. In contrast, if pure steam were developed in this converter and mixed with the gas to be reformed, the product synthesis gas with a relatively large volume would have to be compressed from below 35 ata to the reforming pressure of 60 ata, and this requires large and expensive compressors as well as energy.

Furthermore, even if the total pressure is constant throughout the water side of the reactor, the temperature in the catalyst bed can be controlled and varied to achieve a desired temperature profile. in the catalyst layer along its flow path. Such a desirable profile in an exothermic catalytic reaction limited by the reaction rate near the inlet zone and equal weight near the outlet zone, is that the temperature is relatively high in the inlet zone and relatively low in the outlet zone,

with intermediate temperatures in between. With the reactor in accordance with the invention, this is easily achieved by varying the relative gas flow to the tubes. The desired temperature profile is thus achieved by varying the size of the gas nozzles in a radial direction, the nozzles 225 supplying the pipes near the inlet zone of the converter gas being the smallest in the gas flow opening, the nozzles 224 being larger and 223 the largest. Of course, appropriate nozzle sizes will also be determined by the heat release rate of the catalyst that locally surrounds the tubes. As the reaction takes place along the radial converter gas flow path, and especially in the case of the desired temperature profile, the reaction rate will thus decrease, and the heat release rate proportionally. Preferably, this will be at least partially compensated

for by varying the pipe spacing. The pipes will thus be relatively closer together in the inlet zone. It can too

be desirable to provide external fins on at least some of the tubes, (it is noted that the preferred direction of the converter gas flow, i.e. radially inward, is also advantageous in that relatively fewer tubes towards the center line also provide a relatively larger flow cross-section and thus also helps equalize the converter gas velocities).

Also by setting up cylindrical dividing devices (not shown) in the water side of the reactor, it is clear that steam that is unmixed with gas can be developed to the extent that heat beyond that required to saturate the gas is available and that the pressure of such steam (and consequently the temperature inside the part of the catalyst layer that is cooled thereby) can be varied independently.

It is also seen that this reactor solves all problems with "flare-up", especially as the gas flow in the pipes will induce relatively rapid circulation of the water through the down-

the runs 220 and up through the sprinkled pipes. This circular

tion will cause the tubes, catalyst and converter gas to be brought up to at least close to the water temperature so as to provide rapid start-up and stability for the reactor.

Finally, it can be seen that the reactor according to the invention does not suffer from capacity limitations like other commonly used reactors. The capacity of the multi-layer cylindrical bed developed by Imperial Chemical Industries can only be enlarged, at a given diameter,

by increasing layer depths and using higher gas velocities,

as. both will tend to increase pressure drops too strongly. The so-called Lurgi reactor, in which the gases flow inside catalyst-filled tubes that are cooled by boiling water, can also be given increased capacity with a given diameter simply by lengthening the tubes and increasing the gas velocities, which again leads to excessive pressure drops. In contrast, in the case of the present reactor, increased capacities are achieved simply by increasing the height of the reactor and thereby the cross-section of the flow, without increasing the pressure drop for the converter gas (at the same time, with a relatively very large capacity, undoubtedly many times as large (as existing reactors, the increased bottom-to-top temperature difference of the cooling liquid/gas mixture flowing upwards may be undesirable and tend to limit further increases in size and capacity).

EXAMPLE

A plant for the production of methanol from methane in accordance with the preferred embodiment of the invention, and which is operated at approximately 60 ata pressure and with a production rate of 1000 tonnes per hour. 24 hours uses approximately 506 tonnes per day. 24-hour oxygen obtained by cryogenic separation of air compressed by means of a gas turbine, and compressed by means of centrifugal compressors driven by a steam turbine supplied with steam from the evolution of steam in the process and by economically exploiting sensible heat contained in the exhaust gases from the gas turbine. gas turbine

the fuel is supplied from fuel gas at approximately 17 ata, mainly obtained from the necessary flushing of inert components from the exhaust gas from the catalytic converter.

In the case of direct partial oxidation, substantial heat is provided for the process, together with over approximately 110 tonnes per day methanol. Autothermal synthesis gas formation over the secondary reforming catalyst 100 takes place at about 60 ata down to about 1093°C, and this gas is then cooled by the rotary regenerator 109 to about 316°C. In this still moist reformed off-gas the residual methane content is about 0.75 mol% which requires a steam/atom to total carbon mole ratio, in the total feed to the reformer vessel 31 of only about 0.99, and yet the reformer gas composition is in good distance from any tendency to deposit carbon on the reforming catalyst 100, in part because that approximately 24.5% of the carbon atoms supplied to the reformer vessel 31 are recycled carbon dioxide.

The total yield of methanol from this process is over approximately 93.2% of the total input and fuel methane, including the provision of all necessary fuel for the gas turbine. This yield corresponds to less than approximately 5.85 million kcal less than approximately 5.85 million kcal lower heating value for supply and fuel per tonnes of product methanol.

Although in this presentation the more preferred means

for carrying out the purpose of the invention is described in detail, it will be clear to the expert that for-

different substitutes can also be used within the scope of the invention. With regard to the process aspect of the invention, it is e.g. clear that although a rotating regenerator is preferred because the stated advantages, other means for preheating, reaction and cooling as well as fixed layer, exchange regenerators or recuperators can be used, even if this becomes less advantageous. The invention is thus not limited in any way other than as stated in the subsequent patent claims.

Claims (19)

1. Process for producing a synthesis gas comprising carbon monoxide and hydrogen from a hydrocarbon-containing gas, characterized in that deb hydrocarbon-containing gas is preheated by heat exchange with said synthesis gas to a temperature of more than approximately 816°C, the hydrocarbon-containing gas is reacted with oxygen in the presence of steam so that the hydrocarbon-containing gas is further heated, and the hydrocarbon-containing gas is reformed so that the synthesis gas is formed. 2. Method as stated in claim 1, characterized in that water vapor is added to the hydrocarbon-containing gas by passing the hydrocarbon-containing gas together with water in liquid form in heat exchange conditions to a heat-producing catalytic reaction of gases derived from synthesis gas. 3. Method as stated in claim 2, characterized in that the aforementioned heat-producing catalytic reaction is the conversion of synthesis gas to methanol. 4. Process for the production of methanol from hydrocarbon-containing gas, characterized in that the gas is reformed in the presence of steam and carbon dioxide into a synthesis gas, unreacted steam is condensed from there, carbon dioxide is absorbed from synthesis^ the gas by means of a liquid absorbent so that a carbon dioxide-rich absorbent is thereby produced, then most of the synthesis gas is catalytically converted into methanol , carbon dioxide is stripped from the carbon dioxide-rich adsorbent, and carbon dioxide is mixed with the aforementioned hydrocarbon-containing gas. 5. Method as stated in claim 4, characterized in that a part of the synthesis gas after the aforementioned reaction which catalytically converts most of the synthesis gas to methanol, is contacted in countercurrent with the carbon dioxide-rich adsorbent so that carbon dioxide is stripped from it and undeformed methanol is returned to the aforementioned reforming stage. 6. Process for the production of methanol from a synthesis gas comprising carbon monoxide, hydrogen, inert gases and sulfur-containing gases, characterized in that the synthesis gas is countercurrently brought into contact with methanol produced in the aforementioned process so that sulphur-containing gases are absorbed from the synthesis gas and produces sulphur-containing methanol and then the said sulphur-containing methanol is heated and the sulphur-containing methanol is stripped with a gas comprising the said inert gases after these are flushed out from the process. 7. Process for the production of methanol under an elevated pressure from a synthesis gas comprising the components carbon monoxide, hydrogen and unreformed methane, characterized in that crude methanol containing the aforementioned gas components in solution is depressurized and vaporized to a lower pressure than the aforementioned elevated pressure so that methane is selectively flushed out from the process, and since at least a part of the aforementioned raw methanol is returned after the depressurization evaporation to further contact with gases resulting from the synthesis of the methanol, 8. Process for the production of methanol, characterized by the steps with 1) heat and methanol are produced by partial oxidation of a hydrocarbon-containing gas with a stream comprising compressed oxygen and by absorbing the gaseous methanol produced by said partial oxidation using a volatile absorbent to form a heated methanol-rich absorbent 2) the hydrocarbon-containing gas is then reformed into a synthesis gas 3) the aforementioned synthesis gas is catalytically converted to methanol and 4) methanol is stripped from said methanol-rich absorbent using at least part of said heat. 9. Method as set forth in claim 8, characterized in that a carbon dioxide-rich absorption agent is formed by bringing the synthesis gas in counterflow into contact with an absorption agent and carbon dioxide is stripped from the carbon dioxide-rich absorption agent using at least part of the aforementioned heat. 10. Method as stated in claim 8, characterized in that impurities in the methanol from step (3) are removed by fractionation using steam produced in step (4). 11. Method as stated in claim 8, characterized in that contaminants in the methanol-rich absorbent in step (1) are stripped from it by bringing the methanol-rich absorbent in countercurrent contact with residual synthesis gas containing inert components from step (3) so that a fuel gas is produced consisting of the aforementioned contaminants and the aforementioned inert constituents. 12. Method as stated in claim 11, characterized in that the fuel gas is used as fuel for energy devices to produce the compressed oxygen needed in step (1). 13. Process for the production of methanol from a hydrocarbon-containing gas comprising methane, with a relatively high yield from the methane in the said hydrocarbon-containing gas, characterized in that the said hydrocarbon-containing gas is partially oxidized to produce methanol and a partially reformed gas, the methanol produced thereby being recovered, the said partially reformed gas is reformed into synthesis gas, and the thus produced synthesis gas is catalytically converted into methanol. 14. Method for producing synthesis gas from a hydrocarbon-containing material in the presence of water vapor in a reforming reaction zone and for consuming at least part of the synthesis gas in another, heat-producing reaction zone, characterized by passing a gas together with water in liquid form in heat exchange conditions into the aforementioned second zone so that the reaction temperature in the aforementioned second zone is thereby controlled and heat is recovered from there in the form of water vapor in the aforementioned gas, and then the aforementioned gas is introduced into the aforementioned first zone so that water vapor is thereby supplied to the first zone. 15. Method as stated in claim 14, characterized in that devices are arranged to maintain a temperature profile in the said second zone so that the outlet temperature from the said second zone is kept at a lower temperature than the average temperature in the said second zone. 16. Apparatus for controlling the temperature in an exothermic catalytic reaction zone, characterized by first and second separate chamber means defining a generally annular catalytic reaction zone, a mass of particulate catalyst in the zone, a plurality of heat transfer tubes embedded in the catalyst mass and arranged generally coaxially in said zone, a source of reactant gases connected to one of said chamber^ devices and a reactant gas outlet from said second chamber device. 17. Apparatus as specified in claim 16, characterized by devices connected to the pipes for passing heat-absorbing liquid through the pipes. 18. Apparatus as stated in claim 16, characterized in that devices are connected to the said pipes for passing a heat-absorbing liquid together with a gas through the pipes. 19. Apparatus as stated in claim 16, characterized in that cylindrical screen devices are arranged to contain the catalyst mass in the mentioned zone and devices are arranged to lift at least part of the mentioned screen devices so that used catalyst can be easily released from the layer .