NO133970B - - Google Patents

Description

The present invention relates to a method for the production of synthesis gas, i.e. a mixture of carbon monoxide and hydrogen, and is particularly suitable for the production of synthesis gas for the synthesis of oxygen-containing organic compounds, methanol or hydrocarbons.

Mixtures of hydrogen and carbon monoxide are produced by partial oxidation of hydrocarbon oils. The molar ratio C0/H2 in such synthesis gas is mainly a function of the ratio C/H in the oil. Only small amounts of steam and CC>2 can be supplied directly

to the reaction zone in the synthesis gas reactor without the reaction temperature dropping sharply or unwanted by-products being formed.

The product gas produced by direct partial oxidation of liquid hydrocarbons contains approximately equal volumes of carbon monoxide and hydrogen. The relative quantity ratio between hydrogen and carbon monoxide in such gas mixtures is increased by the reaction taking place over a water gas catalyst. The catalytic water gas reaction at a temperature of between approx. 204 and 538°C are well known. Typically, an iron chromium oxide catalyst is used in the upper temperature range and a zinc oxide-copper oxide catalyst in the lower range.

The present method deals with the production of two streams of synthesis gas - a product stream that has a molar ratio CO/H2 of approx. 1.0 or higher for use as oxo-synthesis gas, and

the second product stream with a lower molar ratio CO/Hg of approx.

0.5 to 1 for use as methanol or Fischer-Tropsch synthesis gas. In one embodiment, the two streams are produced simultaneously by dividing the waste gas from a flow-through gas generator into two streams. The first synthesis gas stream is mixed with C0£ which is produced later in the process and the mixture is reacted at a temperature of at least 8l6°C

in a non-catalytic reverse converter for the production of oxosynthesis gas. The second synthesis gas stream is mixed with steam and the mixture is reacted at a temperature of at least 816°C in a non-catalytic direct converter. C02 is recovered from this second product and introduced into the first reverse converter. If desired, oxygen can be added to the non-catalytic converter, as well as subjecting the converted gases to a renewed non-catalytic conversion reaction.

It is thus an aim of the present invention to provide a new continuous process for the production of synthesis gas without the use of catalysts where the process is independent of external heat input and which can be easily regulated with regard to the molar ratio C0/H2<

Another purpose of the invention is to simultaneously produce from a primary source of synthesis gas, two separate streams of synthesis gas - a first stream with a high molar ratio (C0/H2) and a second stream with a lower molar ratio (C0/H2) than the primary source.

Another purpose is to provide an economical non-catalytic method for producing the gas mixture for use as fuel gas or for the chemical synthesis of oxygen-containing hydrocarbons, methanol and hydrocarbons.

According to the present invention, there is thus provided a method for producing synthesis gas from a gas mixture comprising carbon monoxide and hydrogen to which steam and/or carbon dioxide have been added as an additional gas, and this method is characterized by the fact that the gas mixture is reacted in the conversion zone by a non-catalytic , adiabatic water-gas flow-through reactor at a temperature of 8l6 - 1538°C to a synthesis gas comprising CO, H2 > H20 and C02 whose molar ratio (CO/H2) relative to the molar ratio (C0/H2) of the gaseous starting mixture is greater when the make-up gas is C02 and is less when the make-up gas is steam.

The flow of feed gas essentially consists of

of a mixture of carbon monoxide and hydrogen with smaller amounts of HgO, C02 and a content of free carbon such as soot in an amount of approx. 0.01 to 10 percent by weight (based on the amount of carbon in the hydrocarbon oil) and this feed gas is preferably prepared by non-catalytic partial oxidation of hydrocarbon oil with oxygen, preferably relatively pure oxygen (ie 95 mole percent O2 or

higher) and possibly with steam at a temperature in the range of approx. 1093 to 1649°C. The atomic ratio between free oxygen (not bound oxygen) and carbon in the starting material (O/C ratio) is preferably between 0.80 and 1.5". The molar ratio (CO/l^) in the waste gas stream from the synthesis gas generator is usually approx. 0.56 when the hydrocarbon supplied to the generator is natural gas,

and up to approx. 1.2 when the hydrocarbon for the generator is vacuum residual oil. The reaction time in the gas generator is approx. 2 to 6 seconds.

Essentially all cheap hydrocarbons can be used as feed for this process. E.g. the charge can consist of gaseous, liquid or solid hydrocarbon fuel. Suitable gaseous fuels are natural gas, waste gas from refineries and by-product gas from Fischer-Tropsch conversion. Suitable liquid hydrocarbons cover the entire oil range from propane, petrol and solar oils to heavier residual oils, reduced crude oils and clean crude oils. Likewise, hydrocarbons such as coal, oil, shale oil and tar sands oil are included. Solid carbonaceous substances, possibly in a mixture with said liquid hydrocarbons, include substances such as lignite, bituminous coal and anthracite coal and petroleum coke.

The hydrocarbon oil is partially oxidized in a flow-through synthesis gas generator lined with refractory material under a temperature of 1093 to 1649°C, which is maintained by the heat of reaction, and a pressure of between 1 and 350 atmospheres, forming a primary synthesis gas stream. The synthesis gas generator is preferably a compact non-catalytic flow-through generator of steel, without filler material and lined with refractory material and designed for pressure, of the type described in U.S. Pat. patent 2.809.IO4. Preheating the feed to the syngas generator is not necessary, but usually beneficial. E.g. the charge of liquid hydrocarbon and steam can be preheated to a temperature of approx. 38 to 427°C, and the oxygen supply can be preheated to a temperature between 38 and 399°C. Introduction of steam to the synthesis gas generator can be carried out if desired and possibly depends on the type of hydrocarbon used. E.g. generally no steam is required in connection with gaseous hydrocarbons, approx. 0.1 to 1 part by weight of steam per weight part liquid hydrocarbons and approx. 0.5 to 2 parts by weight of steam' per part by weight solid carbonaceous slurries.

The reaction zone for the non-catalytic water gas converter reaction can consist of a flow-through pressure vessel made of steel without filler material designed for adiabatic operation and lined with refractory material, preferably without constrictions and of a suitable size to give a residence time of approx. 0.1 to 6 seconds.

A stream of additional steam for the direct conversion or carbon dioxide-rich additional gas containing approx. 25 to 95 mole percent CO2 or more for the reverse converter, preferably with temperature

of between 260 and 816°C, is mixed with hot waste gas from the partial oxidation zone and the mixture is introduced into a non-catalytic flow-through water gas direct converter where the reaction takes place at a temperature of at least 816°C. Although such premixing is preferred, the additional steam or the additional carbon dioxide gas can be supplied separately to the converter and mixed and reacted with the primary supply of synthesis gas which is also introduced separately. Although only one converter is shown in the drawing, a number of separate reactors can optionally be used.

Additional H^O or -COg supplied to the converter can be obtained from an external source or can be produced in the process itself.

The quantity is added to the CO2 that is supplied to the primary stream

of synthesis gas must be sufficient to fulfill heat and material balances. The equilibrium constant (K^) for the water-gas reverse converter as shown in Equation I below is a function of the reaction temperature and varies from 1.36 to 3>^5 over a temperature range of 927 - 1538°C.

where (H2), (C02), (CO) and (H20) denote mole fractions (or partial pressures) for the gases in the brackets.

The supplied C₂-containing gas can be a gas mixture containing C0₂, CO, H₂ and H₂O, where the additional gas preferably has at least 25 mole percent CO₂ and significant amounts of CH₂. In a particular embodiment, a C02~ stream containing at least 95 mole percent CO2 is recovered as a by-product from the direct conversion reaction between steam and hot synthesis gas, which simultaneously produces a stream of methanol synthesis feed gas with a mole ratio (CO/H^) of approx. 5 or less. That process is described in more detail below.

The gas from the water-gas reaction is introduced into the non-catalytic adiabatic flow-through reverse converter with a residence time of approx. 0.1 to 6 seconds at a temperature of

at least 816°C, and preferably 0.1 to 2 seconds at a temperature between 927 and 1538°C and under a pressure of approx. 1 to 35^ atmospheres.

In the non-catalytic water gas reverse converter, part of the carbon dioxide is reduced to carbon monoxide, while at the same time a stoichiometric amount of hydrogen is oxidized to water. The net result of the water gas reverse reaction is that the mole ratio (CO/Hg) in the product gas from the reverse converter increases. The elimination of the catalyst is thus a definite economic advantage.

A high temperature is maintained in the water gas reaction zone so that the adiabatic water gas reverse conversion reaction proceeds quickly without a catalyst. The term "adiabatic" in connection with the water-gas converter means that essentially no heat exchange takes place with the surroundings apart from a minor unavoidable heat loss through the reactor walls.

In one embodiment of the invention, the main stream of outgoing synthesis gas is introduced directly into said non-catalytic adiabatic reverse converter essentially at the temperature and pressure that the gas had in the reaction zone of the non-catalytic flow-through synthesis gas generator, i.e. a temperature of between approx. IO93 and 2093°C and a pressure of between approx. 1 and 35O atmospheres. This satisfies the temperature requirements for the next step in the process, i.e. non-catalytic adiabatic water gas reverse condensing. In this way, the costs of heating and compressing the incoming supply gas to the water-gas conversion reaction are saved. This results in a significant financial advantage.

The gases leave the non-catalytic reverse converter at a temperature of over 816°C, and can be cooled in a cooling zone such as a water boiler or similar equipment to a temperature of between approx. 204 and 427°C• The steam produced in the kettle can be used elsewhere in the process or can be led out of the plant. For a detailed description of the cooling system for synthesis gas using waste heat boilers and gas washing towers, reference is made to U.S. patent 2,980,523. A suitable "dipieg" gas-liquid contactor is shown in U.S. Pat. patent 2,896,927.

Entrained solid particles can be showered out from the cooled reformed gas exiting the waste heat boiler by direct contact with cooling water in a gas-liquid contactor, e.g. a shower tower, venturi nozzle or spray contact, a bubble plate or packed column, or combinations of such u ^ ..yr. Common venturi nozzles or jet contactors are described in Chemical Engineers' Handbook, Fourth Edition published by J.H. Perry, New York, McGraw-Hill Co., I963, pp. 18-55, 56.

Excess carbon dioxide can be removed from the reformed gas stream by a conventional regenerative gas washing process, e.g. monoethanolamine, the hot-carbonate or rectisol process. These processes will also remove any H^S present in the product gas. This CO^ can be recycled as part of the additional CO2 introduced into the water gas reverser.

Oxosynthesis gas with a molar ratio (CO/H^) of approx. 1.0 or higher can be produced according to the present method provided that the waste gas from the synthesis gas generator has a correspondingly lower ratio CO/B^ than the desired product gas.

In another embodiment of the invention, a small amount of additional oxygen in free form, preferably relatively pure oxygen (95 mole percent 0^ or more), is introduced into the adiabatic non-catalytic water gas reverse converter zone in an amount sufficient to maintain the temperature at least 816°C and preferably between 927 and 1538°C.

In yet another embodiment, the water gas converter stream undergoes a series, preferably two, successive stages of water gas reverse sorption reactions. The first reverse conversion is carried out in a non-catalytic adiabatic water gas reverse converter, and this gas then undergoes further conversion in a second non-catalytic water gas reverse converter as described for the first reaction.

Additional COg gas is introduced into the first converter and, if desired, into the second converter together with the flow from the first converter. Supplemental oxygen, preferably pure oxygen (95 mole percent 0^ or more) can be introduced into the first or second converter or both. Such multi-stage operation is recommended in special cases when a product gas with the desired composition cannot be easily produced by a single direct or reverse conversion process.

If the water gas feed gas temperature is too low for entry to the first or second conversion reaction, or both, a small amount of supplemental oxygen, preferably pure oxygen (95 mole percent Cv, or higher) can be introduced in admixture with the supplemental CC^.

In such cases, part of the H2 and CO in the converter gas mixture will react with the additional oxygen so that the temperature in the reaction zone is increased to at least 816°C, preferably to between 927 and 1538°C.

The amount of additional H^O supplied to the primary feed gas must satisfy heat and material balance. The equilibrium constant (K^) for the water-gas direct converter is listed in Equation I below. K~ is a function of the reaction temperature cg varies between 0.74 and 0.26 over a temperature range from 927 to 1538°C.

where: (H2), (COr,), (CO) and (H20) denote molar mass ions (or partial pressure) for the gas in question.

The water gas charge is introduced into a non-catalytic adiabatic flow converter with a residence time of between approx. 0.1 and 5 seconds at a temperature of at least 816°C, preferably between 0.1 and 2 seconds at a temperature of between 927 and 1538°C, and a pressure of approx. 1 and 350 atmospheres.

In the water-gas converter reaction, part of the E^O in the feed is reduced to hydrogen, while at the same time a stoichiometric amount of carbon monoxide is oxidized to carbon dioxide. The net result of the water-gas conversion reaction is a reduction in the mole ratio (C0/H2) in the product gas leaving the direct converter.

A high temperature is maintained in the water-gas converter zone so that the direct conversion reaction takes place quickly without . catalyst. According to one embodiment of the invention, the main flow of synthesis gas is introduced directly into said non-catalytic adiabatic direct converter zone with essentially the same temperature and pressure as prevailed in the reaction zone in the non-catalytic flow-through synthesis gas generator without filler material, at a temperature of between 1093 and 1538 °C and a pressure of between approx. 1 and 35O atmospheres, which satisfies the heat requirements for the next step in the process, i.e. the non-catalytic adiabatic direct conversion reaction. In this way, heating and compression costs are saved.

The gas leaving the non-catalytic water gas direct converter with a temperature above 816°C can be cooled in a cooling zone such as a waste heat boiler or direct immersion tube cooler to a temperature of approx. 204 t^1 427°C. The steam produced in the waste heat boiler or immersion tube digester can be used to advantage as all or part of the additional H^O supplied to the water gas direct converter. Surplus steam can be used elsewhere in the process or can be led out of the plant. Entrained free carbon as soot can be showered from the cooled reformed gas that comes out of the waste heat boiler by direct contact with shower water in the gas-liquid contactor as previously described.

Excess C02 and H^S can be removed from the product gas in the usual suitable regeneration equipment as previously described. Methanol synthesis gas with a molar ratio (C0/H2) of approx. 0.5 or more and Fischer-Tropsch synthesis gas with a molar ratio (CO/H^) of approx. 0.5 .oij. 1.0: can be produced according to the present method provided that the outgoing gas from the synthesis gas generator has a correspondingly higher CO/H^ ratio than the desired product gas.

In another embodiment of the invention, a small amount of additional free form oxygen, preferably relatively pure oxygen (95 mole percent or higher), is introduced into the adiabatic non-catalytic water-gas direct converter in an amount sufficient to maintain a temperature in the converter of at least 8l6 °C and preferably 927 to 1538 °C.

According to yet another embodiment, the water gas supply undergoes a number and preferably two successive stages of water gas conversion reactions. The first water-gas reforming reaction is carried out in a non-catalytic adiabatic reactor, and the reformed gas undergoes further reforming in a second non-catalytic water-gas direct reformer as described above for the first reformer. Such multi-stage operation is recommended in special cases when a product gas with a desired composition cannot be easily produced by a single reforming process. Furthermore, if the temperature of the feed gas is too low for the non-catalytic second reformer reaction, a smaller amount of free form supplemental oxygen can be introduced to the second reformer either directly or in admixture with supplemental steam. In this case, part of the H2 and CO in the feed gas will react with the introduced additional oxygen in the direct converter and raise the temperature in the second converter zone to at least 800°C, preferably 927 to 1538°C.

Furthermore, part of the oxo synthesis gas produced by reverse conversion as described above and part of the methanol synthesis gas produced by direct conversion can be mixed into a synthesis gas which has an average composition.

One will understand the invention better in connection with the attached schematic drawing which illustrates the method of the invention. The description and drawing also serve as a particular example according to the invention.

The drawing shows a particular example where both oxo synthesis gas and methanol synthesis gas are produced simultaneously from separate parts of outgoing synthesis gas from a non-catalytic flow-through synthesis gas generator. The oxo synthesis gas has a higher molar ratio C0/H2 than the outgoing synthesis gas, while the molar ratio C0/H2 in the methanol synthesis gas is relatively lower than this content in the outgoing synthesis gas.

EXAMPLE I

The example shows a preferred embodiment in connection with the production of oxosynthesis and methanol synthesis gas from off-gases from an oil refinery.

With reference to the drawing, enter 159» 6 m°l per hour refinery gas with a temperature of 53§°c °S pressure 38»7 kg/cm and the following composition, to synthesis gas generator 1 through line 2 and burner 3:

105.6 mol per hour pure oxygen with 95 mole percent

purity and temperature 149°C and pressure 38.7 kg/cm<2> is introduced into gas generator 1 through pipeline 4 °g is reacted with the refinery gas in the reaction zone 5 with refractory lining in generator 1 during production of a primary synthesis gas supply. 491»9 mo1 Per hour primary synthesis gas supply with temperature 1427°C and pressure 35 kg/cm<2> and with a mole ratio (C0/H2) equal to 0.56 leaves the gas gene-

rator 1 through line 6. The waste gas through pipeline 6 also contains 0.05 weight percent free carbon as soot (the amount of carbon on the basis of carbon in the refinery feed gas).

Using the control valves 7 and 8, outgoing synthesis gas is distributed through line 6 to lines 9 and 10 in the ratio 1 to 3 respectively. 123.0 mol per hour outgoing gas with a temperature of 1427°C in line 9 is passed through valve 8, line 11 and to pipeline 12 where it is mixed with 25.7 m°l per hour of gas containing 100% carbon dioxide with a temperature of 538°C from line 13. The gas mixture which contains approx. 0.21 volumes of C02 from line 13 per unit volume of synthesis gas containing essentially H2 and CO from line 11, is introduced through line 14 to the non-catalytic water gas reverse converter JO with a pressure of 35 kg/cm<2>. There it is formed by an endothermic reaction between C02

and H2, CO and H2O. An analysis of the mixed gases leaving this reverse converter 70 through conduit 15 at a temperature of

about. 1113°C is shown in Table 1.

Although not necessary in this example, free oxygen (95 mole percent O2 or more) may be introduced through line 71 into the reverse converter "JO" in sufficient quantities to maintain the temperature at least 816°C and preferably 927 to 1538°C.

148.7 mol per hour gases through line I5 are cooled to a temperature of 302°C in the boiler 16 and leave this through pipeline 17. 133 mo1 per hour of water at 252°C is introduced to the boiler through line l8 which is converted into saturated steam with a pressure of 42 kg/ cm in the boiler 16 and exits this through line 19. Approx. 113.2 mol per hour excess steam is taken out of the system for taking out the plant through lines 20 and 21 and valve 22.

In a conventional carbon rinsing plant 23, all soot except 2 dpm on a dry gas weight basis is removed in line 17, and this soot can be drained from the system through line 24 as a slurry of soot and water. One can use any carbon extraction system, e.g. washing the gas in line 17 with water in a regular venturi washer.

The soot-free gas in line 25 is introduced into a regular carbon dioxide recovery zone 26 as described above. 112.0 mol per hour oxosynthesis gas is fed out through line 27 and has a ratio C0/H2 of 1.00, which is significantly higher than the ratio

C0/H2 equal to 0.57 in line 6. Gas analysis of stream 27 is listed

in table 1. 12.3 m°i Per« hour 100 percent C02 is recovered and recycled to the converter 70 through the pipelines 28 and 29, the compressor 30, pipeline 31» heat exchanger 32 and lines 13,

12 and 14. The pressure in the C02 flow through line 13 is increased so that it is higher than the pressure in the supply gas flow through line 13, by means of the compressor 30"

About. 368.9 mol per hour supply synthesis gas through line 10 with a temperature of 1427°C is passed through line 33 °S and mixed in line 34 with 19»8 mol per hour of steam from line 35 at a temperature of 538°C. The mixture of steam and synthesis gas in line 34", which essentially consists of H2 and CO, is passed through line 36 to the non-catalytic water gas direct converter 72 where the gases react under a pressure of approx. 35 kg/cm. The reformed outgoing gas stream leaves the converter 72 through line 37 at a temperature of 1352°C and is cooled to a temperature of 302°C in boiler 38 by indirect heat exchange with water entering through line 39 and leaves the plant as steam through line 40. The excess steam is led out of the plant through pipeline 40. The reformed waste gases leave the heat exchanger 38 through line 41 and are further cooled in the heat exchanger 32 by indirect heat exchange with the C02 flow from line 31 as described.

Although it is not required in this example, free oxygen can optionally be introduced through line 73 to the direct converter 72 in a sufficient amount to keep the temperature at least 816°C.

The reformed waste gas from line 42 is introduced into

a normal soot removal plant 43 which may be of the same type as the plant 23 described above. Exit gas is freed of all soot except 2 ppm and is drained from the system through line 44* The soot-free exit gas through line 45 is introduced to a C02 recovery zone 46 which is of the same type as zone 26 described above. You take out approx. 3^0.0 mol per hour reformed synthesis gas with a C0/H2 ratio of 0.50, which is smaller than the molar ratio (C0/H2) in the outgoing synthesis gas stream from line 6, via line 47 from C02 recovery zone 46. An analysis of the gas is reproduced in table 1. A current of 14»5 mol per hour additional C02 is taken out from the C02 recovery zone 46 through line 48 and 49* Approx. 13»4 mol per hour C02 from line 50 is mixed in line 29 a stream

of 12.3 mo1 per hour C02 from line 28. Approx. 1.1 mol per hour COg surplus is taken out of the system through lines 51, 52 and valve 53* ^e combined CCv, flows of a total of 25.7 mol per hour additional C02 is heated in the heat exchanger 32 and fed to the non-catalytic water gas converter 14 through lines 13 and 12 as described earlier.

Although the particular drawing and the particular example describe a system where two product gases are produced simultaneously, namely an oxosynthesis gas and a feed gas for methanol synthesis, the same principles can be used to produce a product gas stream. By way of example, an embodiment producing only oxo synthesis gas is illustrated in Example 2. An embodiment producing only methanol or Fischer-Tropsch synthesis gas is

described in Example III.

EXAMPLE II

With reference to the drawing, 491»9 m°l Pr-time synthesis gas from the non-catalytic flow-through generator 1 passes through line 6 as described in example I. With valve 7 closed and 8 open, this entire gas flow will be led through lines 6, 9 °g H and is mixed in line 12 with 53 > 8 m°l Pr-time 100 percent pure carbon dioxide from an external source which is connected to the system through line 54> valve 55 °g the lines 56 > 49 °g 5<0>, and is mixed in line 29 with 49>1 m°l Per« hour recycled stream of CO^ from the COg recovery zone 26 through lines 28 and 29. By means of the compressor 30, the mixed stream of additional CO,-, in line 29 is passed through line 31 °g to heat exchanger 32 where the temperature is increased to 53^°0 and then the additional CO^ stream is fed

.to wires 13, 12 and 14.

The water gas reverse reaction takes place in the non-catalytic reverse converter 70 at a temperature of approx. 1113°C and a pressure of approx. 35 kg/cm. 44& m°l Per« hour oxosynthesis gas leaves the COg recovery zone 26 through line 27, and has analytical values as shown in table II.

EXAMPLE III

This example describes another embodiment in connection with the production of methanol synthesis gas from residual oil after distillation and shows a multi-stage non-catalytic adiabatic direct conversion and supply of additional oxygen to the water-gas converter.

133.2 mol per hour of residual oil with a temperature of 399°C is reacted with 133»2 m°l Per* hour of pure oxygen as a 99 mol percent stream 02 with a temperature of 149°C, and 116.5 mol per hour steam with a temperature of 399°c« The oil residue has an API specific gravity of 4>5°> a total heating effect of 9700 calories/kg and the following elemental analysis: C 85.8, H2 9.1, S 5 and axis 0, 1.

All outgoing gas from the synthesis gas generator, i.e. 605 mol per hour with temperature approx. 1371°C and pressure approx. 35 kg/cm<2>, and with a composition as shown in table II below, mixed with 390 ml per hour of steam that holds 260°C. The mixture is introduced immediately into a first non-catalytic adiabatic flow-through direct converter where HgO in the gas stream reacts with CO in this and forms more Hg and COg.

995" in mol per hour converted synthesis gas leaves the first non-catalytic adiabatic direct converter with a temperature of 954°C and a pressure of approx. 35 kg/cm<2>. This outgoing gas has a molar ratio CO/Hg equal to 0.54", which constitutes a 53% reduction in relation to the ratio CO/Hg in outgoing gas from the synthesis gas generator. The composition of this gas stream is shown in Table II below.

111 moles per hour of additional steam at a temperature of 260°C

and approx. 2 moles per hour additional oxygen" with a temperature of 149°C is mixed with these 995"! m°l Per« hour converted gas. The gas mixture is introduced directly into a second non-catalytic adiabatic flow-through direct converter for water gas where the added oxygen reacts exothermically with Hg, C or CO and maintains a temperature in the conversion zone of over 927°C > so that the water-gas conversion can take place quickly.At the same time, more CO is converted to COg.

1106.1 mol per hour exit gas leaves the second non-catalytic direct converter with a temperature of 954°C and a pressure of 35 kg/cm<2>. Exiting product gas has a molar ratio (CO/Hg) equal to 0.48, which is a 59 percent reduction compared to the molar ratio CO/Hg in outgoing gas from the gas generator. The composition of this gas stream is shown in Table II. The gas stream is introduced into a heat exchanger boiler where it is cooled to a temperature of approx. 3l6°C by evaporation of 887 mol per hour of water which turns into saturated steam with a temperature of 260°C and a pressure of 4^.8 kg/cm<2>. As described, approx. 617 mol per hour of this saturated steam partly as supply to the synthesis gas generator and thus as supply to the first and second non-catalytic adiabatic direct converters. About. 270 moles per hour of this saturated steam can be used for other purposes.

Claims (8)

. 1. Process for producing synthesis gas from a gas mixture comprising carbon monoxide and hydrogen to which steam and/or carbon dioxide have been added as an additional gas, characterized in that the gas mixture is reacted in the conversion zone of a non-catalytic adiabatic water gas flow-through reactor at a temperature of 8l6 - 1538°C to a synthesis gas comprising CO, Hg, HgO and COg whose molar ratio (CO/Hg) relative to the molar ratio (CO/Hg) in the gaseous starting mixture is greater when the additional gas is COg and is smaller when the additional gas is steam. 2. Method according to claim 1, characterized in that the additional oxygen in free form (95 mole percent Og or more) is introduced into said non-catalytic adiabatic water gas converter zone in an amount that is sufficient to keep the reaction temperature in the converter zone in the range 927 - 1538°C . 3. Method according to claim 1, where the additional gas is COg, characterized in that essentially all COg found in the synthesis gas product stream is removed by treating the product gas in a conventional carbon dioxide extraction plant for the production of a gas stream containing at least 95 mole percent COg, and recycles this COg stream back to the non-catalytic adiabatic water gas reformer zone as at least a portion of said additional COg stream. 4. Method according to claim 3" characterized in that at least part of the flow of additional carbon dioxide is supplied from an external source comprising at least 25 mole percent COg. 5. Method according to claim 1, where the additional gas is COg, characterized in that another adiabatic non-catalytic water gas converter zone is arranged downstream of said reactor, and that further COg is introduced into the second conversion zone and is reacted there with Hg in the synthesis gas product stream at a temperature of at least 8l6°C and at a pressure in the range 1 - 350 atmospheres. 6. Method according to claim 3> where the additional gas is steam, characterized in that the synthesis gas mixture produced in the converter zone is introduced into a waste heat boiler in indirect heat exchange with a flow of water to form at least part of the flow of additional steam. 7. A method according to any one of the preceding claims, wherein a first portion of a feed gas stream comprising CO and Hg is introduced together with a stream comprising additional carbon dioxide into a non-catalytic, adiabatic water gas reverse condensing zone to form a stream of oxo synthesis gas, while another part of said feed gas flow with a flow of additional steam is introduced into a non-catalytic, adiabatic water gas direct conversion zone to form a flow of methanol synthesis gas, characterized in that substantially all COg is removed from the methanol synthesis gas and supplied as at least a part of the additional carbon dioxide to the non-catalytic, adiabatic water gas reverse somf ormer zone. 8. Method according to claim 1, where the additional gas is steam, characterized in that another non-catalytic, adiabatic water gas converter zone is arranged downstream for said reactor and that additional HgO is introduced into said second converter zone and reacted there with CO from said synthesis gas product stream at a temperature of at least 8l6°C. 9« Method according to claim 8, characterized in that additional oxygen is introduced into said second adiabatic, non-catalytic water gas converter zone in an amount sufficient to maintain the reaction temperature in said second non-catalytic water gas direct converter zone at a temperature of at least 927 °C•