This invention relates to a process for the conversion of natural gas to synthesis gas which can be further converted and/or purified as required for production of hydrogen, carbon monoxide, mixtures of hydrogen and carbon monoxide, as well as for the production of methanol, ammonia, dimethyl ether (DME), and/or liquid hydrocarbons. In particu- lar the invention relates to a combined reforming process for the preparation of synthesis gas suitable for conversion of natural gas to such products in plants of large capacity capable of converting natural gas to at least two thousand metric tons per day of such products.

The synthesis gas section is the most expensive section of a plant for conversion of natural gas to end products such as ammonia, methanol, dimethyl ether (DME) , and/or liquid hydrocarbons, both in terms of energy requirements and capital costs. Therefore, major efforts have been focusing on finding alternative process schemes for synthesis gas production which are less expensive and/or more energy efficient .

It is known to provide a synthesis gas, that is, a hydrogen and/or carbon monoxide rich gas by subjecting a hydrocarbon feedstock to a desulphurization step followed by preheating and one or more reforming steps. In steam reforming synthesis gas is produced from hydrocarbon feedstock by re- actions (1) - (3) : C _n H _n , + H ₂ O = n CO + (n + m/2) H ₂ (-ΔH < 0) (1) CO + H ₂ O = CO ₂ + H ₂ (-ΔH = 41 kJ/mole) (2) CH ₄ + H ₂ O = CO + 3 H ₂ (-ΔH= -206 kJ/mole) (3)

During prereforming, reactions 1-3 are carried out in adia- batic reactors, also called prereformers, at relatively low temperatures, normally 350° to 650 ⁰ C. An adiabatic reactor is a reactor in which no heat is transferred to or from the reacting stream (except for heat loss to the surroundings) . The prereforming may also be conducted in so-called heated prereformers, in which the reactions are supported by heat supplied to the reactor by heat exchange with a hot gas, typically a flue gas or a product gas from a subsequent reforming step operating at high temperature. The feed to a prereformer or heated prereformer is normally desulphurised hydrocarbon feedstock mixed with steam.

During primary reforming, reactions 1-3 are carried out in an externally heated reactor, hereinafter referred as a primary reformer. The feed to the primary reformer may be desulphurised hydrocarbon feed mixed with steam or the partly converted product gas from a previous prereforming step. The primary reformer is often a fired tubular reformer consisting of catalyst filled tubes placed in a fur- nace heated by one or several burners and which operates at conditions where the outlet temperature from the catalyst filled tubes is relatively high, normally in the range 650 to 950 ⁰ C. The catalyst filled tubes in a fired tubular reformer may advantageously be straight tubes with an inlet for process gas at one end and an outlet at the other end or tubes closed in one end with process gas first contacting catalyst located in an annular space and then leaving through a central tube in what is known as bayonet tubes. Bayonet tubes are also used in convective reformers as for instance disclosed in our US Patent No. 5,429,809. Convective reformers are a particular form of heat exchange re- formers. The heat required for the reforming of the process gas is provided by convection with a hot flue gas running on the outside of the bayonet tubes containing the catalyst and from the reformed gas on the inside of said tubes. Heat exchange reformers used in such concepts are also known as exchanger reformers or gas heated reformers.

Autothermal reforming (ATR) is a technology used for the production of synthesis gas in which the conversion of a hydrocarbon feedstock or the conversion of a partly con- verted gas from a prereforming step into synthesis gas is completed in a single reactor by the combination of partial combustion and adiabatic steam reforming. Combustion of hydrocarbon feed is carried out with substoichiometric amounts of air, enriched air or oxygen by flame reactions in a burner combustion zone. Steam reforming of the partially combusted hydrocarbon feedstock is subsequently conducted in a fixed bed of steam reforming catalyst.

Secondary reforming is a process in which partly converted (partly reformed) feed from a primary reforming step is further converted by the combination of partial combustion and adiabatic steam reforming. The secondary reformer in ammonia plants will normally be air-blown while in methanol, DME and hydrocarbon synthesis plants it will be oxygen blown. Thus, to a certain extent ATR and secondary reforming processes or reactors resemble one another; however, substantial differences exist. For instance the volumetric ratio between the feed and oxidant is normally lower in autothermal reformers than in secondary reformers. In addition, the heat release and operating temperatures are higher in autothermal reformers and consequently the re- quirements to burner and reactor design in both types of reformers are different.

In autothermal reforming and secondary reforming the steam reforming reactions 1-3 are supplemented by partial combus- tion, which may be represented by reaction (4) :

CH ₄ + 3/2 O ₂ = CO + 2 H ₂ O (4)

It is conventional to produce synthesis gas by first pass- ing a hydrocarbon feedstock mixed with steam through a steam reforming step, which comprises passing the feedstock through a prereformer, then a primary reformer, and finally through a secondary reformer fired with air, enriched air or oxygen. This combination of steam reforming followed by secondary reforming is often referred as two-step reforming and is particularly suitable for the preparation of synthesis gas suitable for methanol and ammonia production. By controlling the amount of reforming occurring prior to the secondary reformer, a synthesis gas having the correct stoichiometry for methanol synthesis or a synthesis gas having the correct hydrogen-to-nitrogen ratio for ammonia synthesis or any other application can be prepared.

Two-step reforming is described in for example Dybkjasr, I., Fuel Processing Technology 42 (1995), p. 85-107. The overall steam to carbon ratio (S/C-ratio) , defined as the molar ratio between the total amount of steam added to the proc- ess in the steam reforming step and secondary reforming step and the carbon contained in the hydrocarbon feed, is high in this process when Ni-based reforming catalyst is used in the steam reforming step, usually in the range 1.3 - 4, in order to prevent undesired carbon formation in the steam reforming step. When catalysts more resistant to carbon formation are used, e.g. noble metal catalysts, the steam to carbon ratio in the steam reforming step can be reduced, typically to 0.5 - 1.3.

The steam required for the process can be added as one stream before a prereforming step and another supplementary stream after the prereforming step; if no prereforming step is present the steam may be added before the steam reform- ing step.

In two-step reforming, the hydrocarbon feedstock is treated as a single feed stream, i.e. without splitting the feed stream into several parallel streams, and reforming is con- ducted in one or more steps arranged in series with respect to the feedstock stream, prior to entering the secondary reforming unit. Because all the hydrocarbon feedstock passes through the primary reformer, and the operating conditions are determined by the feed composition and required synthesis gas composition, two-step reforming may sometimes lead to very low exit temperatures from the primary reformer, which may be inexpedient. This may be the case in situations where limitations in the design of the steam reforming step are defined by the pressure drop through the catalyst filled reformer tubes and not by the heat transfer through the tube walls, thus leading to less optimal design of the steam reforming step. It is also known to produce a synthesis gas by splitting a hydrocarbon feed stream into separate streams prior to reforming in order to reduce the capacity required in the primary reformer and to reduce the overall steam-to-carbon ratio in the reforming process, thereby reducing the mass flow through the plant and consequently the size of equipment .

GB-A-2160516 mentions a method for expanding an existing ammonia plant in which ammonia synthesis gas is produced from a heavy feedstock (feedstock liquid at room temperature) by bypassing a fraction of the heavy feedstock over the steam reforming step. The bypassed stream is mixed with steam and passed through an adiabatic reforming step so as to produce a gas containing methane as the major hydrocarbon component, while the major stream is mixed with steam and subjected to steam reforming. The outlet gas from the steam reforming is mixed with the adiabatically treated bypass gas and then passed through a secondary reformer.

US Patent No. 5,496,859 discloses a combined process for production of synthesis gas suitable for the manufacturing of methanol in which a feed stream is divided into separate parallel streams; one stream is directed to a gasifier, i.e. a non-catalytic partial oxidation unit, and the other stream is directed to a primary steam reformer. Steam is added to each stream prior to introduction into either the gasifier or primary steam reforming. The streams are combined prior to entering a secondary catalytic reformer. This process is specifically tailored for the synthesis of methanol and for this purpose a gasifier operating at high pressure is required so as to produce a synthesis gas which does not require the otherwise needed external compression of synthesis gas for the subsequent methanol synthesis.

US Patent No. 4,888,130 discloses yet another combined re- forming process for the preparation of methanol synthesis gas, in which a desulphurised feed stream, e.g. desulphurised natural gas, is divided in two parallel streams; steam is added to a first stream which is directed to a primary steam reformer. The second stream is preheated to a tem- perature so that upon combination of the streams prior to entering a secondary reformer, the temperature of the combined stream is above 600 ⁰ C. The process allows for a lower overall steam-to-carbon-ratio than in for instance, two- step reforming.

DE-A-10 2006 023 248 discloses a process for the production of synthesis gas in which a hydrocarbon stream is desul- furised, passed through a pre-reformer and then split in two streams. One stream is steam reformed and the other stream is simply heated. The streams are then unified and fed to an autothermal reformer.

In some situations the splitting of the hydrocarbon feedstock into separate parallel streams in order to improve certain process features with respect to the above mentioned two-step reforming can lead to deterioration in other features. For instance, separation of the hydrocarbon feedstock into parallel streams as taught in the above- mentioned US Patent No. 5,496,859 enables the product syn- thesis gas to be suitable for methanol synthesis, but requires for this purpose the use of a gasifier operating at high pressure. This restricts considerably the flexibility of the process with respect to production of synthesis gas necessary for other applications, such as for ammonia synthesis or hydrocarbon synthesis.

In this specification the term "steam reformer", "fired tubular reformer" and "tubular reformer" are used interchangeably. The term "primary reformer" encompasses steam reformers and heat exchange reformers such as convection reformers of the bayonet type. The term "adiabatic oxida- tive reactor" is used generically to denote a secondary or autothermal reformer.

Separation of the hydrocarbon feedstock into parallel streams as taught in the combined reforming of US Patent No. 4,888,130 allows operation at low overall steam-to- carbon ratio. However, at low steam-to-carbon ratios it may be difficult to maintain a suitable temperature difference between the outlet temperature of the tubular reformer and the Boudouard temperature or the equilibrium temperature for CO reduction of the gas. The latter are the temperatures below which a carbon monoxide containing gas will have the potential for carbon formation following the Boudouard reaction 2 CO = C + CO ₂ or the CO reduction reaction CO + H ₂ = C + H ₂ O. When the partial pressure of carbon mon- oxide is high and the gas is in contact with a metal surface at a temperature below the Boudouard temperature or CO reduction temperature, the above reactions are catalyzed by the metal surface. If the temperature of the gas is so low that the metal temperature drops below the Boudouard or CO reduction temperature, highly undesired carbon deposition on the metal surface and/or metal dusting may occur. Metal dusting is tantamount to a catastrophic corrosion on equip- merit parts due to destruction of oxide layers of the Fe-Ni- Cr alloy used in such parts.

Further, in US Patent No. 4,888,130 the fraction of the hy- drocarbon feedstock bypassing the primary reformer goes unconverted before mixing with the other fraction of the feedstock prior to entering the secondary reformer. This limits the possible temperature of the stream bypassing the primary reformer. The limited heat input in the bypass stream increases the size of the primary reformer. We find that as the relatively cold bypass stream with its content of higher hydrocarbons and at temperature about 500 or 600 ⁰ C is combined with the relatively hot exit gas from the primary reformer having normally temperatures above 650 ⁰ C, and for steam reformers in particular about 85O ⁰ C, two significant problems arise. Firstly, there is a risk for thermal cracking of the higher hydrocarbons when the temperature of the bypassed stream is increased. Thermal cracking may take place by pyrolysis of the higher hydrocarbons into olefins and then into coke (carbon) . Secondly, the cooling of the partly converted hydrocarbon feed from the primary reforming may bring the temperature of the combined stream below the Boudouard or CO temperature, thereby promoting carbon formation and/or metal dusting.

It would be desirable to ameliorate the deficiencies in the combined reforming processes of the related art as described above without significantly loosing the benefit provided by splitting a hydrocarbon feed stream into sepa- rate hydrocarbon streams prior to any adiabatic oxidative reforming. In particular, it is an object of the present invention to be able to combine the major and bypass stream with a minimum risk of thermal cracking and metal dusting whilst at the same time being capable of achieving a higher plant capacity in terms of for instance methanol or ammonia production .

Normally, one of either the size of the steam reformer, the supply of oxygen from an air separation unit to the adia- batic oxidative reactor or the exit flow of the adiabatic oxidative reactor represents the bottleneck for capacity in terms of for instance methanol production from the produced synthesis gas. It would be desirable to obtain a higher plant capacity for the same steam-reformer size or duty when for instance the duty of the steam reformer represents the bottleneck for capacity. Similarly, it would be desir- able to obtain a higher plant capacity for the same exit flow from the adiabatic oxidative reactor when this represents the bottleneck for capacity.

It is therefore another object of the invention to provide a process which is more flexible than prior art processes in that it makes possible the full use of the various process units involved, in particular the steam reformer, the adiabatic oxidative reactor and when relevant the air separation unit supplying oxygen to said adiabatic oxidative reactor.

It is a further object of the invention to provide a process that enables a higher exit temperature from the steam reformer than in conventional two-step reforming processes. It is yet another object of the invention to provide a process in which the required duty transferred in the primary reforming is reduced compared to prior art processes.

These and other objects are achieved in the present invention by providing a prereforming step on the first major stream of hydrocarbon feedstock prior to splitting this stream into a second and third streams. Steam is added to the first hydrocarbon stream in connection with said prere- forming step, and additional steam is added to a second stream subsequent to the division of the first hydrocarbon stream. The second stream is preheated to a temperature of 400 to 700 ⁰ C, then passes through a steam reforming step conducted in one or more reforming units, and is subse- quently mixed with a third stream which bypasses said steam reforming step. The second and third stream combine to form a single major stream that is subsequently treated in an adiabatic oxidative reforming step, from which synthesis gas having a given H ₂ /CO/C0 ₂ -ratio is withdrawn.

Hence, according to the invention we provide a process for the production of synthesis gas from a hydrocarbon feedstock comprising: (a) forming a first hydrocarbon stream by desulphurising the hydrocarbon feedstock; (b) mixing said first hydrocarbon stream with steam; (c) forming a partly converted first hydrocarbon stream by passing said first hydrocarbon feedstock stream mixed with steam through a prereforming step; (d) splitting said partly converted first hydrocarbon stream into at least a partly converted second and third stream; (e) mixing said partly converted second hydrocarbon stream with steam; (f) forming a further converted second stream by passing said partly converted second hydrocarbon stream mixed with steam through one or more reforming steps; (g) forming a single major stream by mixing said partly converted third stream and said further converted second stream, (h) contacting said single major stream with an oxygen containing stream in an adiabatic oxidative reactor, and (i) withdrawing from said adiabatic oxidative reactor a stream of synthesis gas containing hydrogen and carbon monoxide, charactherised in that step (e) further comprises preheating of said partly converted second stream to a temperature in the range 400 to 700 ⁰ C prior to entering the reforming step.

By the combined reforming of the invention we have found that it is possible by simple means to increase plant capacity and to combine a cold by-pass stream (third stream) with a hot second stream from a primary reforming step without risking significant carbon deposition due to thermal cracking. At the same time it is possible to reduce the duty transferred in primary reforming compared to prior art processes such as in DE-A-10 2006 023 248.

By the invention it is now possible to optimize the operating conditions in the reforming step of the second stream, so that it is possible to reform a large first hydrocarbon stream at low temperature or a smaller first hydrocarbon stream at higher temperature, whilst at the same time avoiding the disadvantages associated with undesired thermal cracking of hydrocarbon and/or potential metal dusting in the adiabatic oxidative reforming unit. The invention enables at the same time the production of synthesis gas in a process which is less expensive and more energy efficient than in prior art processes.

Normally, one of either the steam reformer, oxygen supply to the ATR or exit flow of the ATR is the bottleneck for capacity, as described above. We find that it is now possible to reach a situation where all these units operate at maximum achievable capacity.

In one particular embodiment of the invention the process further comprises preheating of said first hydrocarbon stream mixed with steam in step (b) to a temperature in the range 400 to 550 ⁰ C prior to passing through the prereform- ing step. This enables increased conversion in the prere- forming step and thus reduction of the required duty of the unit or units utilised in the reforming of the second hydrocarbon stream, for example the steam reformer. Preferably this prereforming is conducted adiabatically .

The prereforming step of the first hydrocarbon feedstock prior to its splitting allows operating the reforming step of the second hydrocarbon stream at more severe conditions and makes it possible to heat the third stream (by-pass stream) to higher temperatures than in prior art processes. The inventive process may further comprise heating the partly converted third stream to a temperature in the range 500 to 750 ⁰ C, preferably 500 to 65O ⁰ C, prior to mixing with the further converted second stream. This enables that the temperature difference between the third and second stream be reduced, whereby the risks of thermal cracking and its concomitant undesired effects are further prevented. In addition, any heat added as preheat duty will lead to a cor- responding reduction in the required duty of the primary reformer.

The prereforming step provides the possibility of heating the second and third stream to higher temperatures than in processes operating without the prereforming step and further provides simplicity and economy of scale compared to a situation where the two streams are prereformed separately after splitting of the dry hydrocarbon feed stream.

We have also found that the splitting of the total amount of steam into two streams, one added before the prereforming of said first hydrocarbon stream and a second added to said second stream, surprisingly allows the operation of the synthesis process at a reduced overall steam to carbon ratio compared to a situation in combined reforming where all steam is added prior to the prereforming of said first hydrocarbon stream, while still being able to operate with higher plant capacities.

The maximum capacity in a plant is related to the amount of product such as methanol or ammonia which can be produced in a process scheme for synthesis gas production consuming a given amount of oxygen and normally involving a steam re- forming step with a given transferred duty. In some cases the maximum capacity may be determined by the maximum allowable pressure drop in the adiabatic oxidative reactor (ATR or secondary reformer) of the maximum practical size. The pressure drop in a given ATR reactor or secondary re- former depends on the total volumetric flow of product gas from the adiabatic oxidative reactor, i.e. its exit flow, so that the maximum achievable capacity is with good ap- proximation defined by a maximum allowable volumetric flow of wet product gas from the adiabatic oxidative reactor. By the invention, process parameters such as steam-to-carbon ratio, split ratio of the hydrocarbon feedstock, reformer duty in a heated steam reformer, and amount of oxygen added in the adiabatic oxidative reactor for the final conversion of the combined stream into synthesis gas, can be expediently adjusted so as to obtain a maximum amount of synthesis gas with predetermined H2/CO/CO ₂ ratios which are suit- able for the production of large amounts (several thousand metric tons per day (MTPD) ) of products in any given downstream application, such as methanol and/or DME synthesis, ammonia synthesis and synthesis of liquid hydrocarbons.

As used herein the term steam-to-carbon ratio (s/c) differs from the term overall steam-to-carbon ratio (S/C) in that the former is defined by the molar ratio between the amount of steam added to the process gas in a given steam reforming step and the carbon contained in the hydrocarbon feed to said reforming step, whereas overall steam-to-carbon ratio is defined by the molar ratio between the total amount of steam added to the process in the steam reforming step, including prereforming and secondary reforming step, and the carbon contained in the total hydrocarbon feed, i.e. first stream.

The amount of steam added in the process to the first hydrocarbon stream is adjusted so that the steam to carbon ratio after step (b) is in the range 0.1 to 10.0, prefera- bly 0.2 to 3.0, more preferably between 0.3 and 1.2, for example 0.35 or 1.1. The steam added to the partly converted second hydrocarbon stream in step (e) is adjusted so that the steam to carbon ratio is about 2 to 10 times greater than that of step (b) , preferably 2 to 5 times greater. For instance, the steam to carbon ratio after step (b) may be 0.35 while in step (e) it may be as high as 2.0, for example 0.6, 0.9, 1.3 or 1.6.

The reforming in step (f) of the second stream may advantageously be performed by passing said stream through a series of one or more reforming steps. The steam reforming may thus be conducted in one or more heated steam reforming stages in series and/or in one or more adiabatic steam reforming stages in series with intermediate heating of feed stock gas leaving the one or more adiabatic steam reforming stages.

The heat required for the reforming in step (f) is preferably provided by indirect heat exchange with the gas withdrawn from the adiabatic, oxidative reformer of step (i) .

The catalytic activity for steam reforming in the one or more reforming units can be obtained either by conventional fixed beds of pellet catalysts, by catalysed hardware, or by structured catalysts. In the case of catalysed hardware, catalytic material is added directly to a metal surface. The catalytic coating of a metal surface is a known process; a description is given in e.g. Cybulski, A., and Moulijn, J. A., Structured catalysts and reactors, Marcel Dekker, Inc, New York, 1998, Chapter 3, and references herein. The appropriate material such as a ferritic steel containing Cr and/or Al, is heated to a temperature preferably above 800°C in order to form a layer of Cr and/or Al oxide. This layer facilitates good adhesion of the ceramic to the steel. A thin layer of slurry containing the ceramic precursor is applied on the surface by means of for instance spraying, painting or dipping. After applying the coat the slurry is dried and calcined at a temperature usu- ally in the region 350-1000°C. Finally, the ceramic layer is impregnated with the catalytic active material which may be based on Ni or a noble metal such as Ru or Rh or any other material with suitable activity for the reforming reactions. Alternatively the catalytic active material is ap- plied simultaneously with the ceramic precursor. Catalysed hardware can in the present invention either be catalyst attached directly to a channel wall in which the process gas flows or catalyst attached to a metallic structured element forming a structured catalyst. The structured ele- ment serves to provide support to the catalyst. Further, catalyst hardware may be in the form of catalyst being deposited in metallic or ceramic structure, which is adhered to wall of the reactor.

Structured elements are devices comprising a plurality of layers with flow channels present between the adjoining layers. The layers are shaped in such a way that placing the adjoining layers together results in an element in which the flow channels can, for instance, cross each other or can form straight channels. Structured elements are further described in for instance US Patent Nos. 5,536,699, 4,985,230, EP patent application Nos. 396,650, 433,223 and 208,929. Two types of structured elements are particularly suitable for the inventive process - the straight- channelled elements and the cross-corrugated elements. The straight-channelled elements require adiabatic conditions and various geometries of these elements are possible. For example, straight channel monoliths are suitable for use in the process of the invention in the adiabatic reactor (s) . Cross-corrugated elements allow efficient heat transfer from the reactor wall to the gas stream. They are also suitable for use in the process of the invention especially in the sections with heat exchange. Other catalysed structured elements can also be applied in the process of the invention such as high surface structured elements. Examples of structured catalysts include catalysed monoliths, catalysed cross-corrugated structures and catalysed rings, e.g. pall-rings.

Both with catalysed hardware applied directly to the wall of the reactor and with structured catalysts, the amount of catalyst can be tailored to the required catalytic activity for the steam reforming reactions at the given operating conditions. In this manner the pressure drop is lower and the amount of catalyst is not more than needed which is especially an advantage if costly noble metals are used. In more conventional applications with pellets, the steam reforming reactors are often designed to maximise heat transfer and pellets are simply placed in channels where the process gas flows often resulting in a vast excess of catalytic activity.

According to the invention, the partly converted second stream obtained after the splitting of the first stream is preheated to a predetermined temperature in the range 400 to 700 ⁰ C, preferably about 65O ⁰ C, prior to entering the primary reforming step. This offers the advantage that the required duty transferred in the subsequent primary reforming step is correspondingly reduced. In one particular em- bodiment, the preheated partly converted second stream is subsequently passed through an additional adiabatic prere- forming step. The resulting stream is then reheated to a temperature in the range 600 to 700 ⁰ C, preferably about 65O ⁰ C, before passing it through the subsequent primary reforming step. The adiabatic prereforming step in the second stream enables i.a. the reduction of inexpedient steam export that otherwise may result from the substantial heat surplus in the convection section of the steam reformer. The inclusion of adiabatic prereforming and subsequent reheating in the second stream increases the amount of heat added from external sources, thereby increasing the total capacity of the process. The reheating of the partly converted second stream to the 600-700 ⁰ C range after said adiabatic reforming step prior to primary reforming may advantageously be effected in a coil containing a reforming catalyst as catalyzed hardware in the form of a coating on its inner wall, where said coil is in contact with hot flue gas in the convective section of the steam reformer. As de- scribed above, catalysed hardware can in the present invention either be catalyst attached directly to a channel wall in which the process gas flows or catalyst attached to a metallic structured element forming a structured catalyst. The structured element serves to provide support to the catalyst.

The term split-ratio as used herein means the amount of volume flow of the third feed stream (by-pass stream) divided by the amount flow of the first stream prior to splitting. Thus, a split ratio of one implies that all hydrocarbon feedstock is bypassed and not treated by a reforming step prior to entering the adiabatic oxidative re- former, while a split ratio of zero implies no by-pass and accordingly that all hydrocarbon feedstock is passed to a reforming step prior to entering said adiabatic oxidative reformer. The first possibility corresponds to a conven- tional synthesis gas process as described in our US patent 6,375,916 based on prereforming of a hydrocarbon feedstock followed by autothermal reforming. The latter corresponds to a conventional two step reforming as described by for example Dybkjasr, I., Fuel Processing Technology 42 (1995), p. 85-107. The split-ratio according to the invention is preferably in the range 0.20 to 0.80, more preferably in the range 0.30 to 0.70, most preferably in the range 0.35 to 0.65. Suitable split-ratios include 0.38, 0.49, 0.57, 0.61 and 0.64. The use of split-ratios lying within the above ranges enables optimisation of the process layout for production of synthesis gas with a given H ₂ /CO/CO ₂ ratio, while at the same time optimising the layout of the reforming step with respect to feed flow of the second hydrocarbon stream and operating conditions. This will be understood by those skilled in the art in that the H ₂ /CO/CO ₂ ratio in the synthesis gas is determined by the ratio between the total amount of heat supplied from external sources to the second and third feed stream by preheating feed streams and by steam reforming, as well as the heat supplied by internal combustion in the adiabatic oxidative reformer.

Any of the preheating steps of partly converted second and third streams may be conducted by indirect heat exchange of said partly converted second and/or third stream with hot flue gas in the convective section of a steam reformer and wherein the tubing inside which the partly converted second and/or third stream is passed is provided as a coil containing a reforming catalyst in the form of a coating on its inner wall or any other suitable structured element. The structured element is preferably selected from the group consisting of catalysed monoliths, catalysed cross- corrugated structures and catalysed rings, such as pall- rings .

Alternatively and in some cases more advantageously, the heat for said preheating step can be provided by heat exchange with the hot exit gas from the adiabatic oxidative reforming step where the primary reformer is a gas heated reformer or exchanger reformer.

In the process of the invention the primary reforming step may be carried out by a suitable reactor such as a steam reformer. In such reformer heat is supplied by external combustion by means of a number of burners arranged in the furnace wall at different levels operated with a low sur- plus of air, typically 5-20% above the stoichiometric amount, i.e. the amount of air which contains exactly the amount of oxygen required for complete combustion of all combustible components in the fuel, so as to provide for a high adiabatic flame temperature. By adiabatic flame tem- perature is meant the temperature that would be achieved from the fuel and air or oxygen containing gas if there is no exchange of enthalpy with the surroundings, for example 2000 °C or higher. The heat for the reforming reaction is thereby supplied by radiation from the hot gas and from the furnace walls to the reformer tubes wherein solid catalyst is disposed and to minor extent by convection from the flue gas, which leaves the furnace at high temperature, typi- cally about 1000°C. In many practical situations steam is of little value, and steam export is therefore often not desirable. When using a fired steam reformer it is not possible to adjust the conditions in such a way that produc- tion of excess steam is avoided. In addition, only about 50% of the fired duty is transferred to the reformer tube wall, thus requiring constant external fuel input. Thermal efficiency in the steam reforming process is accordingly low.

Thus, in a preferred embodiment the above mentioned primary reforming step of the partly converted second steam is conducted by convection reforming. When conducting the primary reforming with a convection reformer, the heat required is provided by hot flue gas from a dedicated burner, which is conveniently adapted at the bottom of said convection reformer. Several convection reformers may be arranged in series with respect to hot flue gas, so that the heat required in the first reformer is provided by the flue gas from the burner, while the heat required in the second reformer is provided by flue gas from its own burner adapted at the bottom of the reformer whilst operating with the flue gas from the previous reformer as combustion air. The partly converted second stream is then fed in parallel to said convection reformers, as described in our US Patent No. 5,925,328. This enables easy control of the flame temperature in the reformers. Hence, in the inventive process the reforming step of the second stream is conducted in one or more convection reformers in which at least part of the heat required is provided by hot flue gas from a burner. Compared to a situation in which the primary reforming is conducted in a steam reformer, where the exit temperature of the second stream normally is about 850 ⁰ C, the provision of a convection reformer with for instance bayonet tubes as primary reformer enables that the exit temperature of the second stream from this reformer is below the exit tempera- ture from the catalyst bed. This exit temperature is normally in the range 600 to 65O ⁰ C. Since the preheated bypass stream normally has a temperature in the range 500 to 650 ⁰ C, the temperature difference between these streams is significantly reduced, or even eliminated, thus further re- ducing the risk of thermal cracking upon combination of these streams prior to entering the adiabatic oxidative reactor.

The steam required in the process may be provided by steam produced from waste heat in downstream process steps. At least part of the steam required in the mixing of said first hydrocarbon stream (step (b) ) is provided by a satu- rator, whereby the hydrocarbon stream is contacted directly with hot water such as process condensate in a suitable ap- paratus which may be heated indirectly by low pressure stream or by other means. In this manner, the hydrocarbon stream is saturated with water at the pressure and temperature prevailing in the apparatus. This enables replacement of steam produced from heat at high temperature with steam produced from heat at lower temperature. Furthermore, when the hot water is process condensate, the saturator provides a convenient means of returning volatile impurities and dissolved gases in the condensate to the process.

By the invention it is also possible to achieve a high degree of flexibility in terms of the composition of the synthesis gas to be prepared for a particular application, in particular for very large plants, i.e. plants capable of producing several thousand MTPD, for instance at least 2000 MTPD or more, such as 10000, 11000 MTPD or more of methanol, ammonia, DME (dimethyl ether) or liquid hydrocarbons or any combination of these products. Synthesis gas for the production of methanol is also suitable for the production of liquid hydrocarbons by Fischer-Tropsch synthesis, while methanol, DME or methanol/DME mixtures are suitable oxygenates for the production of liquid hydrocarbons such as gasoline according to for instance our integrated gasoline synthesis described in US Patent 4,481,305.

The invention encompasses therefore also a process according to claim 1 further comprising converting the synthesis gas from step (i) into ammonia, methanol, DME, liquid hydrocarbons, or combinations thereof.

The invention is illustrated by the accompanying drawings, in which Figure 1 shows a process scheme according to the process of US 4,888,130. Figure 2 shows a process scheme according to a conventional two-step reforming process, and Figure 3 and 4 shows process schemes according to particular embodiments of the present invention as described below.

Fig. 1 shows a process scheme for the production of synthesis gas by a combined reforming process according to US Patent No. 4,888,130. Relevant process conditions are shown in Table 1. A dry hydrocarbon feed stream 1 is split into a first stream 2 and a second (bypass) stream 3. The first stream is mixed with a stream 4 of steam to produce a stream 5 with a predetermined steam to carbon ratio (s/c) . Stream 5 is preheated in a preheater 6 to produce stream 7 which is passed through prereformer 8 to produce partly converted stream 9. Stream 9 is passed through second preheater 10 to produce stream 11 which is then fed to steam reformer 12, where it is reacted to produce reformed stream 13. The second stream 3 is heated in a heater 14 to produce stream 15 which is then mixed with the reformed stream 13 to produce a combined stream 16. The combined stream 16 is reacted with a stream 17 of oxygen from an air separation unit (not shown) and a small additional amount of steam 18 in an adiabatic oxidative reactor 19 to produce a stream 20 of wet synthesis gas. The additional steam 18 of steam is added to protect the burner parts in the adiabatic oxidative reactor. The synthesis gas stream 20 is after cooling and removal of condensed reconverted steam, further treated for production of e.g. pure hydrogen and/or carbon monoxide, methanol, DME, liquid hydrocarbons or ammonia.

Fig. 2 shows production of synthesis gas by one embodiment of the so-called two step reforming process according to the description in e.g. Dybkjasr, I., Fuel Processing Technology 42 (1995), page 85-107. Relevant process conditions are shown in Table 1. The dry hydrocarbon feed stream 1 is mixed with a stream 2 of steam to produce a wet feed stream 3 with a predetermined steam to carbon ratio (s/c) . Steam 3 is preheated in preheater 4 to produce stream 5, which is reacted in a prereformer 6 to obtain stream 7. Stream 7 is heated in heater 8 to produce prereformed stream 9, which is then fed to steam reformer 10 to produce reformed stream 11. Stream 11 is reacted with oxygen stream 12 from an air separation unit (not shown) and a small additional amount of steam 13 in adiabatic oxidative reactor 14 to produce a stream 15 of wet synthesis gas. The additional stream 13 of steam is added to protect the burner parts in the reactor. The synthesis gas stream 15 is after cooling and removal of condensed reconverted steam, further treated for production of e.g. pure hydrogen and/or carbon monoxide, methanol, DME, liquid hydrocarbons or ammonia.

Fig. 3 shows a process scheme for the production of synthesis gas according to one preferred embodiment of the inven- tion. Relevant process conditions are also shown in Table 1. The dry hydrocarbon feed stream 1 is mixed with a stream 2 of steam to produce a major wet feed stream 3. Steam 3 is preheated in preheater 4 to produce wet feed stream 5 which is reacted in prereformer 6 to produce stream 7. Stream 7 is subsequently split to produce a second stream 8 and a third (by-pass) stream 9. Stream 8 is mixed with a stream 10 of steam to produce stream 11 which is then preheated in heater 12 to produce stream 13. Stream 13 is fed to steam reformer 14, where it is reacted to produce reformed stream 15. The divided third stream 9 is heated in heater 16 to produce stream 17 and is then mixed with the reformed stream 15 to produce a combined stream 18. Stream 18 is reacted with oxygen stream 19 from an air separation unit (not shown) and a small additional amount of steam 20 in an adiabatic oxidative reactor 21 to produce a stream 22 of wet synthesis gas. The additional steam 20 is added to protect the burner parts in the adiabatic oxidative reactor. The synthesis gas stream 22 is after cooling and removal of condensed reconverted steam, further treated for production of e.g. pure hydrogen and/or carbon monoxide, methanol, DME, liquid hydrocarbons or ammonia. Fig. 4 shows another preferred embodiment of the invention for which relevant process conditions are also shown in Table 1. The difference with respect to the embodiment of Fig. 3 is that in Fig. 4 heater 12 produces a stream 12a, which is reacted in a second adiabatic prereformer 12b to produce a further prereformed stream 12c. Stream 12c is heated in heater 12d to produce prereformed stream 13 which is further processed as explained in Fig. 3. The inclusion of the second prereformer 12b and the second heater 12d in- creases the amount of heat added from external sources, thereby increasing the total capacity of the process.

Table 1 compares results obtained by combined reforming as illustrated in Fig. 1 or Two-step reforming as illustrated in Fig. 2 with the results obtained by the invention as illustrated in Fig. 3 and 4.

Table 1

In all cases in Table 1, the dry hydrocarbon feed is lean natural gas, i.e. gas with above 90 vol% methane. Furthermore, the fired duty of the steam reformer is 150 Gcal/h, and the amount of oxygen fed to the adiabatic oxidative re- actor is 4000 MTPD. In Fig. 1, 3 and 4, the split ratio of the dry hydrocarbon feed stream in the combined reforming case of Fig. 1 and the wet hydrocarbon feed stream in the invention, Fig. 3 and 4 have been adjusted so that the outlet temperature of the steam reformer is 850°C. The tem- perature of the preheated stream is 550°C for the dry bypass stream in Fig. 1; 520°C for non-reformed feed streams to prereformers or fired reformers and 650 ⁰ C for prere- formed gases. The outlet temperature from the steam reformer and the outlet temperature from the adiabatic oxida- tive reactor and the composition of the synthesis gas produced in the latter are calculated in all cases. The operating parameters in each case are adjusted in such a way that the content of methane in the produced synthesis gas is 2 dry mole %. Finally, the potential production of methanol from the synthesis gas is calculated for all cases assuming 95% conversion of the carbon oxides (CO + CO ₂ ) contained in the synthesis gas, and the specific consumption of hydrocarbon feed, and oxygen (per ton of methanol product) are calculated for all cases.

From Table 1 it is seen that higher capacities in terms of methanol production (MTPD) can be obtained with two step reforming (Fig. 2) and with the embodiments of the invention illustrated in Fig. 3 and especially Fig. 4. Further- more it is seen that the inlet temperature to the adiabatic oxidative reactor (ATR) is quite similar in the different cases, but that it is obtained by different means: (i) In the combined reforming of Fig. 1, the inlet temperature to the ATR of 721°C is obtained by mixing a dry, unconverted natural gas stream at 550°C with a steam reformer effluent stream at 850°C. This involves risk of carbon for- mation or corrosion by heating of the higher hydrocarbons contained in the natural gas and/or by inadvertently cooling the reformed gas in the second stream to a temperature significantly below the Boudouard temperature or the equilibrium temperature for CO reduction of the gas. (ii) In the two-step reforming of Fig. 2, the inlet temperature of 717°C to the ATR corresponds to the exit temperature of the steam reformer. The exit stream from the steam reformer is not combined. There is no heating of gas containing higher hydrocarbons and no cooling of reformer effluent gas. Thus, there is no risk of carbon formation or corrosion .

(iii) In the two embodiments of the invention Fig. 3 and 4, the inlet temperatures of 734°C and 740°C are obtained by mixing a wet, prereformed by-pass stream at 650 ⁰ C with a reformed gas stream at 850°C. This involves no risk of carbon formation from heating of higher hydrocarbons, since these have previously been removed by steam reforming. Further, the risk of carbon formation or corrosion by inadvertently cooling this gas to temperatures significantly below the Boudouard temperature or the equilibrium temperature for CO reduction of the gas is significantly reduced. The skilled person would appreciate that where the inlet temperature to the ATR and the Boudouard temperature of the ATR feed as defined in Table 1 are about the same (± 1O ⁰ C) the risk of metal dusting is not significant. From Table 1 it is further seen that the processes according to the invention allows the operation of the synthesis process at a reduced overall steam to carbon ratio compared to a situation where all steam is added prior to the prere- forming of said first hydrocarbon stream as in the combined process of Fig. 1. Both prereforming and adiabatic oxidative reforming can now be operated at a lower steam to carbon ratio than that normally required in primary reforming in order to avoid undesirable carbon formation. It is also apparent that the process of the invention, particularly the embodiment of Fig. 4, results in a higher plant capacity for methanol production.

In the process schemes of Fig. 1 to 4 the fired duty of the steam reformer and oxygen supply represent the bottleneck for capacity. By the embodiment of Fig. 4 a higher capacity for the same reformer size is obtainable compared to the two-step reforming of Fig. 2, but also compared to the combined reforming of Fig. 1. By the embodiment of Fig. 3 a higher capacity for the same reformer size is obtainable with respect to the combined reforming of Fig. 1.

Compared to the two-step reforming process of Fig. 2, the invention allows that the exit temperature from the steam reformer be higher, and thereby improved design of the steam reformer in situations where the pressure drop through the catalyst filled tubes is limiting rather than heat transfer through the walls of the catalyst tubes.

Table 2 compares the results obtained with the process of Fig. 4 with different values of the steam-to-carbon ratio obtained after step b (stream 3) and step e (stream 11) . Table 2

In all cases in Table 2, the dry hydrocarbon feed is also lean natural gas, the fired duty of the steam reformer is 150 Gcal/h, and the amount of oxygen fed to the adiabatic oxidative reactor is 4000 MTPD. The split ratio of the wet hydrocarbon feed stream has been adjusted so that the outlet temperature of the steam reformer is 850 ⁰ C. The preheat temperature is 430°C for non-reformed feed streams to pre- reformers operating at steam-to-carbon ratio 0.35, 520 ⁰ C for prereformers working at higher steam to carbon ratio, and 650 ⁰ C for prereformed gases. The outlet temperature from the adiabatic oxidative reactor and the composition of the synthesis gas produced is calculated in all cases. The operating parameters in each case are adjusted in such a way that the content of methane in the synthesis gas is 2 dry mole%. The production of methanol from the synthesis gas is calculated for all cases assuming conversion of 96% of the carbon oxides (CO + CO ₂ ) contained in the synthesis gas to methanol by known means, and the specific consump- tion of hydrocarbon feed, and oxygen (per ton of methanol product) are calculated for all cases. Finally, the potential production of ammonia is calculated assuming conversion of 95 % of the hydrogen and carbon monoxide contained in the synthesis gas to ammonia.

It is seen from Table 2 that the production of methanol and ammonia decreases slightly as the steam-to-carbon ratios decrease. However, the invention enables to adjust the process so that the total wet flow of synthesis gas out of the adiabatic oxidative reactor is the limiting factor for the obtainable capacity. It is therefore now possible to lower the steam-to-carbon ratios and at the same time increase the production of ammonia or methanol as shown in Table 3, which is based on the same cases as those shown in Table 2, but with varying reformer duty and oxygen flow to the ATR. Cases of columns 3 in Table 2 and 3 are the same.

Hence, the invention enables high production capacities and at the same time a low overall steam-to-carbon ratio, which is often expedient as it reduces the mass flow through the plant and accordingly the size and cost of plant equipment. Table 3

The process according to the invention as shown by the data of Tables 2 and 3 is therefore highly flexible and allows the tailoring of the process in order to achieve the highest possible capacity in terms of for instance MTPD ammonia or methanol from the produced synthesis gas for such very large plants. Normally, one of either the steam reformer, oxygen supply to the ATR or exit flow of the ATR is the bottleneck for capacity, as described above. By the invention it is possible to reach a situation where all these units operate at maximum achievable capacity.