WO2003062141A1

WO2003062141A1 - Process for preparing synthesis gas by autothermal reforming

Info

Publication number: WO2003062141A1
Application number: PCT/GB2003/000279
Authority: WO
Inventors: Erling Rytter; Arne GRISLINGÅS; Karina Heitnes Hofstad
Original assignee: Statoil Asa; Smaggasgale, Gillian, Helen
Priority date: 2002-01-23
Filing date: 2003-01-23
Publication date: 2003-07-31
Also published as: GB0201505D0

Abstract

A process for reforming feed gas in a reactor having an outer reactor shell wherein at least a portion of the partly or wholly reformed gas is recycled within the outer reactor shell.

Description

A PROCESS FOR PREPARING SYNTHESIS GAS BY AUTOTHERMAL REFORMING

The present invention relates to a process for reforming feed gas. More particularly the present invention relates to a process for the production of hydrogen or synthesis gas.

Synthesis gas, which is often referred to as syngas, is generally produced from a hydrocarbon feedstock such as natural gas in which the principal hydrocarbon present is methane. These processes are known collectively as 'reforming' processes. They are generally carried out in the presence of a suitable catalyst. The synthesis gas produced will comprise the desired products carbon monoxide and hydrogen and may additionally include some or all of carbon dioxide, steam, unconverted hydrocarbon and inerts such as nitrogen and argon. These additional components may be separated from the carbon monoxide and hydrogen prior to the synthesis gas being utilised in further reactions. However, in certain reactions it may be possible or desirable to retain some or all of the additional components.

Depending on the method by which the synthesis gas is produced, it is generally possible to adjust the ratio of hydrogen to carbon monoxide and/or to minimise the amount of other components present.

Synthesis gas of an appropriate composition has a number of uses and is important in the production of a range of valuable compounds and compositions including methanol, acetic acid, dimethyl ether, other oxygenates, ammonia, olefins, fuels and synthetic crude oil. The reaction of the synthesis gas to form some of these compositions will be via the Fischer-Tropsch reaction.

The ratio of hydrogen to carbon monoxide may be adjusted depending on the use to which the synthesis gas is to be put. Thus, for example, where hydrogen is the primary reactant in the subsequent reaction process, the production of the synthesis gas is adjusted so that the proportion of hydrogen present is increased. Hydrogen production, whether as the main product or as the predominant product in a synthesis gas mixture is important as hydrogen has many valuable uses. It can be used as the feed for fuel cells which may be used in, for example, power generation and to power vehicles. Hydrogen can also be used as a direct fuel source. Further, it is essential in a number of chemical processes including hydrodesulphurisation, hydrotreating, hydrogenation and hydrocracking reactions.

The following chemical reactions are fundamental in the production of hydrogen or synthesis gas from natural gas:

(1) CH₄ + H₂O → CO + 3H₂ Steam reforming

(2) CH₄ + v.o → CO + 2H₂ Partial oxidation

(3) CO + H₂O → CO₂ + H₂ Shift; reaction

The steam reforming reaction (1) is highly endothermic and the required heat is generally provided by external firing. In one alternative arrangement, the steam reforming reaction may be combined with the partial oxidation reaction (2) which is an exothermic reaction, giving autothermal or near autothermal reforming.

Generally, reformers operate at high temperature. For example, steam reformers and autothermal reformers typically operate in the region of about 800°C to about 1000°C with partial oxidation reformers operating at even higher temperatures. The use of high temperatures has the benefit of securing high methane conversion (low methane slip) and low carbon dioxide production. Lower temperatures may be utilised for prereformers and heat integrated reformers such as synthesis gas or flue gas heated reformers.

In many operations used for hydrogen or syntheses gas production, oxygen may be fed to the reaction zone. The oxygen can be provided by adding air to the reactor. The air will comprise nitrogen as the major component in addition to the oxygen. However, the addition of further volumes of gas to the system can involve increased costs in construction of the pipe-work to accommodate the larger gas volumes. In order to reduce the load on the system, the air may be separated in an air separation unit into oxygen and nitrogen so that only the neat oxygen is added to the system. Whilst this separation minimises the problems associated with the presence of high volumes of gas, the air separation units are themselves costly and may represent fifty percent or more of the capital investment.

Recently it has been suggested that oxygen transport membranes may be usefully integrated into the plant for production of hydrogen or synthesis gas such that the oxygen may be introduced without the high capital expenditure associated with the air separation units. In oxygen transport membrane devices, air is fed to one side of the membrane, which is generally a dense ceramic sheet which at elevated temperatures allows oxygen ions to pass through the membrane to the second side of the membrane. The oxygen ions can then react with the natural gas or other feed gas supplied to the second side. The membrane may be provided in the form of a tube where the air is supplied though the centre of the tube and the feed gas is supplied to the reactor which contains the tube such that the fed gas is located outside the tube.

On the air side of the membrane (cathode side) the reaction can be represented as:

(4) V₂ O₂ + N₂ + 2e → O²- + N₂ and that on the feed gas side (anode side) can be represented as:

(5) CH₄ + O²- - CO + 2H₂ + 2e

The reaction between the feed gas and the oxygen is rapid and this leads to a low partial pressure of oxygen on the anode side of the membrane such that a concentration gradient across the membrane is established. This concentration gradient drives further oxygen ion transport across the membrane.

The equilibrium requirements mean that a high temperature, typically in the region of about 600°C to about 1200°C, or more typically of about 800° to about 1000°C, will be utilised. Steam may be added to the anode chamber. This may be added to reduce coking or, for an autothermal or near autothermal reaction, may be added as an essential component of, for example, the steam reforming reaction. It will therefore be understood that the reaction on the anode side of the membrane may be more complex than that detailed above. For example, in one simplified reaction scheme the reactions may include:

(6) CH₄ + H₂O → CO + 3H₂ and

(7) H₂ + O²- → H₂O + 2e

One problem associated with the process for the production of hydrogen or synthesis gas is the deposition of carbon or carbonaceous species in the reactor. This is known as coking. The presence of such carbon deposits may reduce the efficiency of the reactor and may cause any catalyst present to be deactivated more rapidly then would be expected.

The main reactions which may form carbon from a synthesis gas type gas composition may be summarised as follows: (8) CH₄ → C + 2H₂ Methane decomposition

(9) 2CO → C + CO₂ Boudouard reaction

(10) CO + H₂ → C + 3H₂O CO reduction.

The corresponding equilibrium constants for a given temperature may be calculated from standard tabulated thermodynamic data as:

K _{g; eq} = (P _H2 / P _H4) _eq »^"

K io, eq ^— (P H20 ' PIG Pco) eq where P_x is the partial pressure of the gas x. These values may be compared with the stoichiometric constants for the actual gas composition and temperature. If coking is to be avoided the stoichiometric constants for the actual compositions should be close to or above the equilibrium values. Unfortunately, as the reactions progress, the stoichiometric constants for at least reactions (9) and (8) move such that coke formation is favoured. It should be noted that reaction (8) is endothermic and therefore the equilibrium constant increases with temperature which facilitates coke formation. Reactions (9) and (10) exhibit the reverse behaviour.

In addition, even where the feed gas is natural gas, it may comprise appreciable amounts of higher hydrocarbons and carbon dioxide in addition to the methane. The presence of these may further accelerate the coking problem. One solution will be to incorporate a prereformer into the process scheme to reduce the concentration of at least the higher hydrocarbons in the feed. However, the incorporation of this additional process step will increase both the capital and operating costs of the plant.

In an alternative arrangement for the production of hydrogen or synthesis gas, the feed gas may comprise higher hydrocarbons such as naphtha or paraffins, oxygenates such as methanol or gasified biomass. The problem of coking of the reactor is particularly severe for such feeds.

In GB 1460187 it was suggested that carbon formation could be minimised by a process in which a proportion of the synthesis gas produced which is at least equal in volume to the volume of gaseous hydrocarbons fed to the reactor should be separated from the product stream and then recycled to the inlet of the reaction zone.

US 5837034 and US 6106591 also suggest that recycling of a proportion of the product gas to the inlet of the reactor may be advantageous in the reduction of coking effects. These documents suggest that the recycle may be utilised in a process incorporating an oxygen transport membrane. Whilst the processes described in the above-mentioned documents go some way to addressing the problems associated with coking in reforming systems, the capital costs of the pipe- work associated with recycling a proportion of the reactants to the inlet substantially increases the plant costs. Further, in these arrangements, it is usually necessary to treat the recycle stream to ensure that it is at an appropriate temperature and pressure to be reintroduced to the reactor. Therefore the plant will generally require the inclusion of additional heat exchangers and compressors.

Thus, according to the first aspect of the present invention, there is provided a process for reforming feed gas in a reactor having an outer reactor shell wherein at least a portion of the partly or wholly reformed gas is recycled within the outer reactor shell.

Whilst the reaction of the present invention may be applied to any reactor which is suitable for use in reforming feed gas, it is particularly suitable for use with autothermal reforming reactors and reactors including oxygen transport membranes.

Surprisingly, the process of the present invention allows for the efficient production of hydrogen or synthesis gas. In particular, the danger of coking or the amount of coke produced is substantially reduced in comparison with known arrangements. Further, since the recycle is carried out internally of the reactor shell, the requirement to adjust the temperature and/or pressure of the recycle stream is minimised and may even be obviated.

The presence of the internal recycle may also reduce the net amount of steam required to produce the same volume of product gas than was required in processes available heretofore.

The feed gas is preferably natural gas but may be any suitable gaseous feed gas, for example, feed gas comprising higher hydrocarbons such as naphtha or paraffins, oxygenates such as methanol or biomass. By "feed gas" we mean any suitable feed gas whether fed directly to the reactor or feed gas which has been first subjected to a prereforming process. For ease of reference, the gaseous feed, whether fed directly to the reactor or pre-treated in any manner, will be referred to as "feed gas" .The reforming process will preferably occur in the presence of steam.

Whilst the process of the present invention may be utilised in the reforming of any feed gas, it is particularly suitable for use in the production of hydrogen or synthesis gas. It may be applied to any suitable arrangement for the production of hydrogen or synthesis gas including autothermal reforming, steam reforming, combined reforming, partial oxidation, compact reforming, convective reforming, heat integrated reforming and gas heated reforming. It is particularly suitable for use in a steam reforming process. It is also particularly suitable for use in a process in which oxygen or an oxygen containing gas is fed to the reactor. In a most preferred arrangement gas is contacted with the feed gas via an oxygen transport membrane.

Any suitable oxygen transport membrane design of reactor may be used in the preferred arrangement of the process of the present invention. Examples of designs of membranes and reactors incorporating such membranes can be found in US 4791079 and US 5846641 which are incorporated herein by reference. In particular, the membrane may be tubular, flat, corrugated or have an honeycomb shape. However, an oxygen transport membrane in the form of one or more tubes is preferred. The oxygen or oxygen containing gas is supplied to one side of the membrane and the feed gas to the other side. The oxygen side and the feed gas side of the membrane may be interchanged if required. However, in a most preferred arrangement, the oxygen containing gas is supplied to the inside, i.e. tube side, of the membrane.

The presence of the internal recycle is particularly advantageous for reactors including an oxygen transport membrane, since the recycled stream will act to heat the feed within the reactor to the desired temperature for oxygen flux, thereby minimising or even obviating the requirement to pre-heat the feed gas.

This is particularly significant as the performance of oxygen transport membranes generally increases with temperature. The critical lower temperature will vary with the materials employed and the overall reactor design but in general, a temperature above 700°C, preferably above 750°C, more preferably above 800°C is required at the position where the feed gas first sees the membrane.

This internal heating of the feed has the added advantage of reducing the propensity for coking to occur in the inlet tubes since the feed gas may be passed through the inlet tubes at a lower temperature than that at which coking is likely to occur. Further advantages are noted where the present invention is applied to the reactors including oxygen transport membranes. First, the mixed feed gas (fresh feed + recycle gas) contains a substantial amount of hydrogen that rapidly will react with the oxygen that is transported through the membrane (Equation 7), and thus enhance the temperature further towards the optimum operational temperature. This factor is particularly important if the natural gas or other hydrocarbon feeds are not prereformed, as they then will not contain any hydrogen at all unless gas recycle is employed. Second, a reduced proportion of the methane will be present in the mixed feed for reaction with steam according to Equation 6. As this latter reaction is endothermic, it will tend to cool down the membrane material. In total, the reduced temperature and concentration gradients due to recycle as the mixed feed gas propagates along the oxygen transport membrane(s) in the reactor, will give rise to less thermal and compositional stress of the membrane material(s), and more even oxygen transport and anode reactions. It is thus clear that the use of the internal recycle of the present invention will allow the design and operation of the reactor system to be simplified.

Whilst recycling may require the inclusion within the reactor shell devices having moving parts such as a pump to drive the recycle, in a most preferred arrangement such devices are not required and the recycle may be driven by means of a mechanical arrangement having no moving parts.

Any suitable mechanical arrangement may be used to drive the recycle. However, in one preferred embodiment at least one gas driven ejector may be used. These ejectors, which are also know as venturi type jets, may be used as simple and efficient compressor units which will provide the necessary pressure differential to drive the recycle loop. The feed gas, which will generally have been preheated, will generally serve as the motive gas for the or each gas driven ejector.

Whichever method or arrangement is utilised to facilitate the internal recycle of reformed gas, the system may include a single recycle loop or a plurality of recycle loops which may be arranged in any suitable configuration.

Any suitable proportion of reformed gas may be recycled in accordance with the present invention. However, the benefits of the present invention may be obtained when the recycle ratio is from about 0.3 to about 3.0, more preferably from about 0.5 to about 1.5.

The reforming reaction may be carried out in the presence of a catalyst. The choice of catalyst will depend on the reforming reaction which is to be carried out. Suitable catalysts for use in steam reforming reactions of natural gas include, but are not limited to, nickel on alumina supports. Some catalyst may contain amounts of oxides, typically less than 20 wt%, more typically less than 10 wt% of other oxides including MgO, CaO, SiO₂, K₂O and La₂O₃. These may be in the form of mixed spinels, for example a calcium-alumina spinel. Where required, promoters, and in some arrangements sulphur compounds, may be added. Any suitable support may be used. Whilst alumina is commonly used, hydrotalcites may also be suitable.

The present invention will now be described, by way of example, by reference to the accompanying drawings, in which: Figure 1 is a schematic diagram of one arrangement of the present invention;

Figure 2 is a schematic diagram of one alternative arrangement of the present invention;

Figure 3 is a schematic diagram of a second alternative arrangement of the present invention; and

Figure 4 is a graph of modelled results.

It will be understood by those skilled in the art that the drawings are diagrammatic and that further items of equipment such as heat exchangers, distillation columns, pumps, vacuum pumps, compressors, gas recycle compressors, temperature sensors, pressure sensors, pressure relief valves, control valves, flow controllers, level controllers, holding tanks, storage tanks, and the like may be required in a commercial plant. The provision of such ancillary items of equipment forms no part of the present invention and is in accordance with conventional chemical engineering practice.

For convenience, the process of the present invention will be described with reference to the production of synthesis gas in an oxygen transfer membrane reactor.

As illustrated in Figure 1, in the most simple arrangement feed gas, such as natural gas and steam which has been preheated to preferred reaction temperatures, and is preferably heated to a maximum temperature of from about 600°C to about 650°C and most preferably not above 700°C, is fed via line 1 to a reaction zone 2 located within a pressure shell 3. The reaction zone 2 will generally house a catalyst such as alumina supported nickel 4. An oxygen transport membrane in the form of a tube 5 will extend into the reaction zone 2 within the pressure shell 3. Air or oxygen is fed in line 6 into the cathode side of the membrane such that oxygen ions pass through the membrane to the reaction zone where they react with the feed gas. A proportion of the synthesis gas is recycled in line 7 and the remainder is removed from the reactor shell in line 8.

The arrangement of Figure 2 illustrates an alternative arrangement by which the internal recycle may be arranged. In the specific arrangement illustrated, three small recycle loops are present. The number of loops used will depend on the architecture and design of the oxygen transfer membrane reactor.

As shown in Figure 2, the reactor comprises three oxygen transfer membranes, 9a, 9b, and 9c. The feed gas will be introduced in stream 10 and will then pass downwardly through reaction zones 11a, lib, and lie which contain a suitable catalyst. Product gas and unreacted gas will then flow outwardly from the apertures each identified as 12 at the base of the reaction zones 1 la, 1 lb, 1 lc into conduits 13a, 13b, 13c and 13d. Here flow will generally be in an upwardly direction. Gas traveling upwardly through the conduits will then pass through corresponding apertures 14 which will enable some or all of the gas to be recycled to the corresponding reaction zone and eventually removed via line 15.

One alternative arrangement is illustrated in Figure 3. Here, two oxygen transfer membranes 16a and 16b are located within reaction zones 17a and 17b. Air is provided to the cathode side of the membrane in lines 18a and 18b. Illustrated in this embodiment are gas driven ejectors which are used as simple units for applying the necessary pressure differential in the recycle loop. Here feed gas fed in line 19 passes through a narrow region 20 in the inlet tube 21. The narrow region is formed by obstacles 22a, 22b, 22c and 22d. In general, gas flows downwardly in inlet tube 21 and enters the reaction zones 17a and 17b via apertures 23a and 23b where reaction will occur between the feed gas and the oxygen passing through the oxygen transport membranes 16a and 16b. As the velocity of the feed gas in inlet pipe 21 increases when the gas flows across the obstacles cause by 22a, 22b, 22c, and 22d, the gas pressure drops and reformed gas and/or unreactive gas is sucked in from the shell side of the reactor through the apertures 23a and 23b and is then subjected to recycle. The product stream is removed from the reactor in line 24.

The pressure increase in the recycle loop will depend on the pressure drop through the oxygen transport membrane reactor and this in turn will depend on the design of the reactor and the catalyst deployment system. If the motive gas for the ejector which, in the arrangement of Figure 3 is the primary feed in line 19, is available at pressure substantially above the reactor outlet's pressure, the pressure of the recirculated gas, i.e. the suction gas, in the ejector can be increased depending on the recycle ratio. Thus in the arrangement of Figure 3, if the feed gas in line 19 has a pressure in the region of from about 10 bar to about 20 bar above the pressure in product stream 24, the pressure of the recirculated gas can be increased by an amount typically in the region of 1 to 2 bar. Where higher pressure increases are required, more complex ejector units may be utilised. The ejector units may be formed from any suitable material provided that they are able to withstand the high temperatures utilised within the reactor. If coking in the feed line may occur, for example if the process is to be operated at a low steam to carbon ratio in the gas such as less than 1.2 or more likely less that 0.8 by mole, the feed line and/or the ejectors may be made wholly or in part from a ceramic material or may have a ceramic lining.

The ability to recycle internally product gas will reduce the net steam feed requirement to the reactor.

The present invention will now be described with reference to the following examples. Process simulations were conducted using the standard modeling tool HYS YS for a number of conditions. The simulations are based on the assumption that equilibrium is reached when the gaseous reaction products of the synthesis gas are leaving the catalyst section of an oxygen transfer membrane type reactor. This assumption will be approximately correct for a suitable catalyst employment and loading. Further, it is assumed that the reactor is adiabatic i.e there is no heat loss to the surroundings. The exit temperature of the synthesis gas is calculated from the chemical reactions taking place with a specified feed composition and feed inlet temperatures. Specifically, all calculations assume that the air feed to the inside of the oxygen transfer membrane tubes has a temperature of 900°C. This temperature can be obtained by heat exchange with the effluent gases. Further, the temperature of the depleted air is fixed at 950°C. Utilisation of oxygen in the air is set to 80% unless specified separately. The feed gas is natural gas with the molar composition being 85% methane, 8.5% ethane, 3.6% propane and 2.9% butanes, unless specified otherwise which has been prereformed at 470°C to give essentially a mixture of methane, carbon dioxide and hydrogen. The prereformed gas is heated to 650°C before entering the synthesis gas reactor, with the exception of one case where the effect of reducing the temperature to 550°C is explored. Equilibrium calculations where performed for an exit pressure of 15 bar with an assumed pressure drop of 2.5 bar through the reactor.

In the following study of the effect of recycle ratio (RR), the steam to carbon ratio (S/C) is set to 2.0. This S/C ratio serves to demonstrate the effect of RR, and represents a steam feed that is intermediate between what in practice normally would be used for hydrogen production (higher steam, typically in the S/C range 2.5-4.0) and Fischer-Tropsch or methanol synthesis (lower steam, typically in the S/C 0-1.5). Figure 4 illustrates varying oxygen to carbon ratio (O₂/C) a number that is roughly proportional to the area needed of the oxygen transport membrane, or even to the cost of the reactor. From Figure 4 it is apparent that the inlet temperature of the mixed gases can be increased significantly. An RR of about 1.0 appears to be sufficient to raise the temperature to a level where adequate ionic oxygen transport through the membrane can be achieved.

Further, a critical temperature interval for coke formation and metal dusting is circumvented. In this connection, reference can be made to Table 1, where the gas composition after mixing fresh preformed feed with recycled gas is documented. Table 1. Dry gas composition (mol%) after mixing prereformed feed with recycle gas, and H₂O/C ratio, compared with exit composition.

Oxygen/Carbon Recycle Ratio (RR) H2 CO CO2 CH4 Steam/Carbon Temp (°C)

0.45 0 17 0.2 8.3 74 1.80 650

0.45 0.5 40 8.7 9.7 41 1.74 711

0.45 1.0 49 12 10.3 29 1.71 745

0.45 3.0 60 16 10.9 13 1.65 805

0.45 ReactorExit 68 19 11 1.2 1.58 874

0.50 0 17 0.2 8.3 74 1.80 650

0.50 0.5 40 9.5 9.4 41 1.76 736

0.50 1 49 13 9.8 29 1.75 785

0.50 3.0 59 18 10.3 13 1.72 867

0.50 Reactor Exit 68 21 11 0.2 1.67 966

0.55 0 17 0.1 8.3 74 1.80 650

0.55 0.5 39 10.2 9.2 42 1.80 771

0.55 1.0 48 14 9.5 29 1.80 838

0.55 3.0 58 19 9.9 13 1.79 951

0.55 Reactor Exit 67 23 10.3 0.0 1.79 1086

Table 2. Comparison of equilibrium constants for coke forming reactions with corresponding stoichiometric constants for selected O₂/C and RR values.

Recycle shifts the gas composition towards the equilibrium composition, or more specifically, increases hydrogen and reduces the methane content, thus preventing carbon formation according to reaction (8) above.

For a comparison of the equilibrium constants for coke forming reactions, with corresponding stoichiometric constants for selected oxygen to carbon ratios and recycle ratio values, reference may be made to Table 2. From this table, it can be seen that the stoichiometric constants are roughly one order of magnitude larger than the equilibrium constants for carbon monoxide hydrogenation and the Boudouard reaction. These reactions are therefore unlikely to occur in the presence of the recycle and therefore carbon deposition is reduced. As far as the methane decomposition route for obtaining carbon is concerned, the values at the inlet (0.3 and 4.0) without recycle seem unfavourable, but industrial practice with a proper selection of materials shows that coking can be avoided up to 600-650°C due to kinetic constraints. It is clear that further heating of the prereformed gas to about 800°C, creates a considerably larger gap to the equilibrium constant simultaneously with the increased gas diffusion rates and therefore there is severe danger of coking of the reactor materials. However, recycle of the product gas diminishes this gap significantly thereby reducing the kinetic driving force for coking.

The process of the present invention is applicable in many fields and in each of the fields in which prior art arrangements were utilised. If the present method is applied primarily for the production of hydrogen, a high steam/carbon ratio will be chosen as the shift reaction will maximise hydrogen production. The carbon content of the feed will be converted to a mixture of CO and CO₂ that may be treated further with a high, and optionally low, temperature shift converter in order to produce even more hydrogen. The oxygen feed to the oxygen transport membrane can be relatively moderate although the following lower equilibrium temperature will increase the amount of unconverted methane (methane slip). For an S/C ratio of 3.0 and O₂/C of 0.45, the methane slip is 2.7 mol% whereas by increasing O₂/C to 0.475 only 1.2 mol% of the carbon in the feed is left as unconverted methane. The level of purity required in the hydrogen product will depend on the end use to which the hydrogen is to be put.

Where the product of the reforming reaction is synthesis gas for subsequent use in methanol production, a rather moderate oxygen consumption will maximise the so-called stoichiometric number which is defined as (P_H2-Pco₂)/(Pco⁺Pco₂) ^^ut again due consideration must be taken to the methane slip. However, the stoichiometric number which defines the possible utilisation of the synthesis gas in the methanol synthesis will not reach the theoretical value of 2.0. In practice, this value should be even higher as a hydrogen surplus in the gas recycle loop around the methanol reactor prevents coking and by-product formation and some hydrogen unavoidable will be loss through necessary purging. Therefore, where the synthesis gas is to be used in methanol production, it may be advantageous to remove some carbon dioxide from the synthesis gas or add a smaller steam reformer to enhance the stoichiometric number value.

The synthesis gas produced in accordance with the present invention is also suitable for use in Fischer-Tropsch Synthesis. There are several variations known in the literature and these are selected depending on the desired products and the type of hydrocarbon feed. Currently, much attention is given to the conversion of natural gas to long chained paraffins in the presence of a cobalt containing catalyst with subsequent hydroisomerisation and hydrocracking to diesel fuel, petrochemical naphtha and base oil. The theoretical overall H₂/CO ratio need for such a process will be in the region of 2.05 which means that the stream to carbon ratio should be reduced as much as possible, preferably below 1.0 provided that coking can be prevented. Somewhat higher oxygen consumption and lower steam addition may be required than for hydrogen and methanol production. This will reduce H₂/CO slightly by increasing carbon monoxide production at the expense of carbon dioxide and make methane slip negligible. Results for a Fischer-Tropsch case, a methanol production and a hydrogen production case are set out in Table 3. In each case, a dry gas consumption (mol%) after mixing prereferred feed with recycle gas and a stream/carbon ratio are coupled with reactor exit compositions. The S/C and O₂/C values refer to the ratios fed to the prereformer and to the oxygen transport membrane reactor respectively.

In an alternative arrangement, reactor designs may be utilised in which the oxygen transfer membrane feed tubes are changed for a direct air or oxygen feed system such as in autothermal reformers or combined reforming and partial oxidation reactors.

Where a dry gas composition (mol%) is fed to an autothermal reformer after mixing prereformed feed with recycle gas with a steam to carbon ratio of 2.0 in the feed to the prereformer and an oxygen (preheated to 260°C) to carbon ratio of 0.5 in the feed to the autothermal reformer, the results are set out in Table 4a and a comparison of equilibrium constants for coke forming reactions with corresponding stoichiometric constants for selected oxygen to carbon ratios and recycle ratios to the autothermal reformer are shown in Table 4b.

Table 3. Dry gas composition (mol%) after mixing prereformed feed with recycle gas, and steam/carbon ratio, compared with reactor exit composition. S/C and O2/C refers to the ratios fed to the prereformer and in the OTM reactor, respectively .-

FT-case: S/O0.5, O2/C=0.525, H2/CO(product)= =2.07

Oxygen/ Recycle Ratio (RR) Steam/Carbon Temp. CC-

H2 CO C02 CH4 Carbon

0.525 0 9.4 0.4 6.5 83.7 0.35 650

0.525 1.0 38.8 16.8 5.0 39.4 0.38 823

0.525 ReactorExit 64.9 31.4 3.6 0.1 0.45 1128

MeOH-case: S/C=1.0, O2/C=0.45, SN(product)=1.84

Oxygen/ Recycle Ratio (RR) Steam/Carbon Temp. (°C⁾

H2 CO C02 CH4 Carbon

0.45 0 12.7 0.2 7.3 79.8 0.83 650

0.45 1.0 43.7 14.6 7.2 34.5 0.80 753

0.45 Reactor Exit 66 25 7 2 0.74 911

H2-case: S/C=3.0, O2/C=0.45

Oxygen/ Recycle Ratio (RR) Steam/

H2 CO C02 CH4 Temp. ("C-* Carbon Carbon

0.45 0 21.4 0.1 9.1 69.4 2.77 650

0.45 1.0 52.5 9.9 12.4 25.2 2.64 740

0.45 Reactor Exit 69.7 15.2 14.3 0.8 2.46 851

Table 4a. Dry gas composition (mol%) for an autothermal reformer (ATR) after mixing prereformed feed with recycle gas, and Steam/Carbon ratio, compared with reactor exit composition. S/C and O2/C refers to the feeds to the prereformer and ATR, respectively.

ATR-case: S/O2.0, 02/C=0.5

Oxygen/Carbon Recycle Ratio (RR) H2 CO CO2 CH4 Steam/Carbon Temp. (°C)

0.5 0 17.5 0.2 8.3 74 1.80 650

0.5 1.0 48.6 12.7 10.1 28.6 1.75 772

0.5 Reactor Exit 68 20.6 11.1 0.3 1.66 936

Table 4b. Comparison of equilibrium constants for coke forming reactions with corresponding stoichiometric constants for selected O₂/C and RR value for ATR.

For comparison, a dry gas composition (mol%) and the temperature at the reactor exit for a reformer after first mixing pre-reformed feed and externally recycled synthesis gas was modelled. All conditions were set as in Table 1 except that the temperature of the recycle gas and fresh feed is adjusted to 650°C. The results for a steam to carbon ratio of 2.0 and an oxygen to cabon ratio of 0.5 are set out in Table 5.

Table 5

By comparison with Table 1 it can be seen that where the external recycle is zero, ie no recycle is present, the same synthesis gas composition and exit temperature are noted as where noted for any internal recycle ratio. However, where external recycle is introduced, more unconverted methane is noted (increased methane slip) and further increased production of carbon dioxide at the expense of carbon monoxide.

It can be noted that where the external recycle ratio is 1.0 an exit gas composition is obtained that is close to the values obtained with internal recycle with an oxygen to carbon ratio of 0.45. It will therefore be seen that moving from external to internal recycle saves 20% of the oxygen consumption which in turn offers a significant cost saving.

Claims

1. A process for reforming feed gas in a reactor having an outer reactor shell wherein at least a portion of the partly or wholly reformed gas is recycled within the outer reactor shell.

2. A process according to Claim 1 wherein the reactor is an authothermal reforming reactor or a reactor including one or more oxygen transport membranes.

3. A process according to Claim 1 or 2 wherein the feed gas is natural gas.

4. A process according to any one of Claims 1 to 3 wherein the reforming reaction is carried out in the presence of steam.

5. A process according to any one of Claims 1 to 4 for the production of hydrogen or synthesis gas.

6. A process according to any one of Claims 1 to 5 wherein oxygen or an oxygen containing gas is fed to the reactor.

7. A process according to any one of Claims 1 to 6 wherein the recirculating is driven by a mechanical arrangement having no moving parts.

8. A process according to Claim 7 wherein the mechanical arrangement is one or more gas driven ejector.

9. A process according to Claim 8 wherein the feed gas serves as the motive gas for the or each gas driven ejector.

10. A process according to any one of Claims 1 to 9 wherein the recycle ratio is from about 0.3 to about 3.0.

11. A process according to any one of Claims 1 to 10 wherein the recycle ratio is from about 0.5 to about 1.5.