WO2012084135A1

WO2012084135A1 - Process for reforming hydrocarbon

Info

Publication number: WO2012084135A1
Application number: PCT/EP2011/006179
Authority: WO
Inventors: Charlotte Stub Nielsen; Kim Aasberg-Petersen
Original assignee: Haldor Topsøe A/S
Priority date: 2010-12-22
Filing date: 2011-12-08
Publication date: 2012-06-28

Abstract

The invention relates to a process and plant for the production of synthesis gas from a hydrocarbon feedstock, in which the hydrocarbon feedstock is passed through a heat exchanger reformer and autothermal reformer; the hot effluent gas from the autothermal reformer is used as heat source for the reforming reactions occurring in the heat exchange reformer and a quench stream is added to said hot effluent gas from the autothermal reformer.

Description

Title I Process for Reforming Hydrocarbon

The present invention relates to a process and plant for production of gas rich in hydrogen, particularly

synthesis gas for the production of ammonia, methanol, dimethyl ether (DME) , hydrogen and hydrocarbons by

Fischer-Tropsch synthesis. The invention relates further to the production of synthesis gas by means of

particularly a series arrangement of heat exchange reforming and autothermal reforming stages, in which the heat required for the reforming reactions in the heat exchange reforming stage is provided by hot effluent synthesis gas from the autothermal reforming stage. More particularly, the invention relates to the introduction of a quench stream comprising steam to the hot effluent gas from the autothermal reforming stage.

The use of a product stream of reformed gas as a source of heat in heat exchange reforming is known in the art. Thus, EP-A-0033128 and EP-A-0334540 deal with parallel arrangements, in which a hydrocarbon feed is introduced in parallel to a tubular reformer and heat exchange reformer. The partially reformed gas from the tubular reformer is then used as heat source for the reforming reactions in the heat exchange reformer.

Other parallel arrangements combine heat exchange reforming and autothermal reforming. EP-A-0983963 , EP-A- 1106570 and EP-A-0504471 deal with processes in which a hydrocarbon feed is introduced in parallel to a heat exchange reformer and autothermal reformer. The hot product synthesis gas from the autothermal reformer is used as heat exchanging medium for the reforming

reactions occurring in the heat exchange reformer.

In EP-A-0440258 we disclose a process in which the hydrocarbon feed is first passed through a first heat exchange reformer to provide a partially reformed stream. The partially reformed stream is then introduced in parallel to a tubular reformer and second heat exchange reformer. The product streams from both reformers are combined and introduced to an autothermal reformer. The product gas from the autothermal reformer is used as heat source in the second heat exchange reformer, while the product gas from said second heat exchange reformer is used as heat source in the first heat exchange reformer.

Series arrangements are also known in the art. Our patent DK 148882 discloses a process for production of synthesis gas, in which the hydrocarbon feed is passed through a heat exchange reforming and autothermal reformer, and where the product gas from the latter is used as heat source in the heat exchange reformer . US Patent No .

4,824,658 and US Patent No.6, 296, 679 (Fig. 2 in both references) disclose a process in which the entire hydrocarbon feed is first introduced to a heat exchange reformer, then passed to a tubular reformer and finally to an autothermal reformer. The product gas from the autothermal reformer is used as heat source in the heat exchange reformer. US 4,376,717 and our US 2009/0184293 disclose a process in which a hydrocarbon feed is first passed through a tubular reformer; the partially reformed gas is then subjected to heat exchange reforming and finally

autothermal reforming. The product gas from the latter is used as heat source in the heat exchange reforming. In our US 2009/0184293 we found specifically that by

providing a process in which the entire hydrocarbon feed is passed through a tubular reformer, heat exchanger reformer and autothermal reformer in a series

arrangement, the risk of metal dusting is significantly reduced. In otherwise conventional processes such as heat exchange reformers in parallel or series with either a tubular reformer or autothermal reformer, metal parts of the heat exchange reformer may experience temperatures in the prohibitive range of metal dusting temperatures, i.e. about 400-800°C as hot effluent gas from the autothermal reformer at typically about 1000°C cools during its passage through the heat exchange reformer.

JP 59217605 discloses an apparatus having a shift

reaction part for CO-conversion in addition to a

reforming part in a body shell, capable of giving

hydrogen from hydrocarbons in a compact apparatus . The reforming reaction receives heat from a combustion catalyst bed. In our co-pending patent application DK 201000039 we describe a process in which the entire hydrocarbon feed is passed through a steam reformer, heat exchanger reformer and autothermal reformer in a series

arrangement, and a cooling medium is added to the heat exchange reformer separately from the actual process gas fed to this reformer. The use of heat exchange reformers in the production of synthetic fuels by the Fischer-Tropsch synthesis has potential for significant benefits. Potential benefits include reduced oxygen consumption, lower capital cost per unit of product, and a higher carbon and energy efficiency.

In plants for production of chemicals such as methanol, the use of heat exchange reformers may have similar advantages.

Heat exchange reformers may also be coupled in a similar manner with other partial oxidation reactors. In addition to autothermal reforming (ATR) , such reactors include non catalytic partial oxidation (POX) such as gasifiers, and catalytic partial oxidation (CPO) . ATR and CPO are provided with a fixed bed of catalyst.

However, for many processes such as those mentioned above, especially for large scale plants it may be preferred to operate with a low steam-to-carbon ratio: although operation at high steam-to-carbon ratios reduce the risk of metal dusting, it also means higher flow rates due to the increased amount of steam in the feed. In other words, operating at high steam-to-carbon ratios means that the capital cost due to the use of larger equipment may be prohibitively large, especially for large scale plants. Furthermore, high steam-to-carbon ratio means that a larger amount of carbon dioxide is formed in the process. This is in many cases a

disadvantage such as for example in plants for the production of synthetic fuels by the low temperature Fischer-Tropsch synthesis. In the low temperature

Fischer-Tropsch synthesis carbon dioxide is considered an inert and not a reactant . Operating at low steam-to-carbon ratios in plants

comprising a heat exchange reformer creates a number of challenges. One such challenge is metal dusting. Metal dusting is a type of metallic corrosion that may be encountered when gases containing carbon monoxide come into contact with metals above ca. 400°C, particularly in the range 400-800°C as described above. Metal dusting is described extensively in the literature.

Metal dusting is a highly complex corrosion process which is not completely understood. However, metal dusting is often represented by the following reaction:

CO + H₂ → C + H₂0 (1) The formed carbon leads to corrosion of the construction material possibly by a mechanism including carbide formation and/or dissolution of the carbon in the

material . It is clear from reaction (1) that the potential for metal dusting increases when partial pressure or

concentration of carbon monoxide and/or hydrogen increase and when the partial pressure or concentration of steam decreases . This may be expressed by the reaction

quotient, Q:

PH2O/(P_H2*PCO) where P_H2o is the partial pressure of steam in the gas, P_H2 the partial pressure of hydrogen and P_Co the partial pressure of carbon monoxide.

A reduction of Q increases the thermodynamic potential for metal dusting.

According to the present invention, one preferred layout of a plant for production of synthesis gas is a heat exchange reformer upstream and particularly in series with an autothermal reformer. The hydrocarbon feedstock (for example desulphurised natural gas) is mixed with steam and the resultant mixture is directed to the catalyst side of the heat exchange reformer. In the heat exchange reformer, the gas is then steam reformed according to the following reactions:

CH₄ + H₂0 <→ CO + 3H₂ (2)

CO + H₂0 <→ C0₂ + H₂ (3)

Reactions similar to reaction (2) may take place for any higher hydrocarbon (e.g. ethane, propane etc) present in the feed. The gas leaving the heat exchange reformer is close to chemical equilibrium for reactions (2) and (3) above. Typically, the exit temperature is 600-850°C or preferably 675-775°C. The partially reformed gas leaving the heat exchange reformer is passed to the autothermal reformer. In the autothermal reformer also oxygen and in some cases a small amount of steam is added. Synthesis gas is formed by a combination of steam reforming and partial oxidation in the autothermal reformer. The gas leaving the ATR is free of oxygen and generally reactions (2) and (3) are close to chemical equilibrium. The temperature of this hot effluent gas from the ATR is between 950 and 1100°C, typically between 1000 and 1075°C.

This hot effluent gas leaving the autothermal reformer comprises carbon monoxide, hydrogen, carbon dioxide, steam, residual methane, and various other components including nitrogen and argon. This synthesis gas is passed to the non-catalytic side (herinafter also called "shell side") of the heat exchange reformer. This gas is cooled by supplying heat to the catalytic side of the heat exchange reformer by indirect heat exchange . The exit temperature from this side of the heat exchange reformer would typically be in the range from 500-800°C. In cases where the synthesis gas is required for a

Fischer-Tropsch synthesis unit, a tail gas from the

Fischer-Tropsch synthesis unit is preferably added upstream the autothermal reforming reactor and downstream the catalytic side of the heat exchange reforming

reactor, i.e. the tail gas is added to the primary reformed gas .

The tail gas contains carbon monoxide, carbon dioxide, hydrogen, various hydrocarbons including olefins and a range of other components. The gas leaving the autothermal reforming reactor

contains carbon monoxide. The amount of carbon monoxide increases when the steam-to-carbon ratio is reduced (and the value of Q in eq. (1) also decreases) . As indicated above, the process economic benefit when using a heat exchange reformer is largest at low steam-to-carbon ratios . This creates the problem of being able to operate at low steam-to-carbon ratios while at the same time avoiding or minimising the extent of metal dusting.

Carbon formation via the exothermic reactions 2CO-> C + C0₂ (Boudouard reaction) and CO + H₂ -> C + H₂0 (CO- reduction) is a precursor for metal dusting (MD)

corrosion. The exothermic reactions are favoured at low temperatures. However, the reaction rates are higher at higher temperatures. As a result, the MD potential for a given gas will be highest in a medium temperature range. In the series arrangement of heat exchange reformer and ATR part of the heat exchange reformer shell side will indeed be in the temperature region where the potential for MD is significant.

In stand alone ATR plants (e.g. no heat exchange reformer upstream the ATR) the challenges of metal dusting are overcome by refractory lining the connection and waste heat boiler in the temperature region where the potential for MD is significant. This solution cannot be employed when using a heat exchange reformer in, for instance, a series arrangement with an autothermal reformer (as it would prevent the desired heat exchange with the gas on the catalyst side) . Either the reactor material has to be protected with expensive materials that can withstand metal dusting, or the potential for metal dusting must to be reduced so damage can be avoided. The present

invention addresses the latter approach. We have found that despite of an increase in heat

transfer area in the heat exchange reactor, the addition of a supplementary gas stream in the form of a steam quench stream to the effluent hot gas from the

autothermal reformer which is used as heating medium in the heat exchange reformer has a significant advantageous effect in the plant in terms of avoiding undesired metal- dusting while still operating at low steam-to-carbon ratios, thus avoiding the need of large equipment

elsewhere in the plant.

Hence, according to the invention a steam quench stream is introduced to the exit gas from an ATR or CPO or POX in order to reduce the metal dusting potential in at least the shell side of the heat exchange reformer. This so-called steam quench stream contains preferably more than 90% (by volume) of steam (H₂0 in the vapour phase) , more preferably more than 95%, and most preferably more than 99%. This is possible with a minimum of impact on the plant capital costs as the reaction conditions in the upstream heat exchange reforming and autothermal reforming stages are almost unaffected and the composition of the dry product gas (synthesis gas) is essentially unchanged. This is counterintuitive since the introduction of the steam quench stream will also reduce the temperature of the gas stream entering the shell side of the heat exchange reformer, which is not desired because it will increase the heat transfer area of the heat exchange reformer. Although there is a modest increase in heat transfer area of the heat exchange reformer and further downstream in particularly the waste heat boiler of a cooling train used to cool the synthesis gas, the savings obtained in equipment size throughout the whole plant far than compensate for the slightly larger heat exchange reformer and waste heat boiler.

Accordingly, we provide a process and plant for the production of synthesis gas from a hydrocarbon feedstock according to the following features of the invention in correspondence with the appended claims:

Features of the invention

1. Process for the production of synthesis gas from a hydrocarbon feedstock, comprising the steps of:

(a) reforming said hydrocarbon feedstock in an

endothermic reforming stage in a heat exchange reformer, where the hydrocarbon feedstock is

submitted to indirect heat exchange with a mixture stream comprising hot effluent synthesis gas from step (b) , withdrawing from the heat exchange

reformer said mixture stream as a stream of cooled synthesis gas, and withdrawing from the heat

exchange reformer an effluent stream of primary reformed gas ;

(b) passing the primary reformed gas from the heat

exchange reformer through an autothermal reforming stage (ATR) , catalytic partial oxidation stage (CPO) , or partial oxidation stage (POX) , and

withdrawing a stream of hot effluent synthesis gas of which at least a portion is used as heating medium in the heat exchange reformer of step (a) ; characterised in that the process further comprises forming said mixture stream comprising hot effluent synthesis gas from step (b) by adding a stream comprising steam to at least a portion of said hot effluent

synthesis gas.

By the term "hydrocarbon feedstock" is meant a stream fed to the process comprising hydrocarbon in its widest term, an organic compound comprising hydrogen and carbon. The hydrocarbons are as simple as e.g. methane CH₄, and may comprise more complex molecules.

By "indirect heat exchange" is meant that there is no direct contact between the catalyst and the heating medium, and thereby between the flow passing through the catalyst and the heating medium, because these are separated by a metal wall, i.e. the wall of the tube containing the catalyst.

It would be understood for a person skilled in the art that when producing ammonia, the autothermal reforming stage (ATR) is actually a secondary reforming stage.

It would also be understood that adding the stream comprising steam, hereinafter also referred as quench stream, may take place either upstream the heat exchange reformer, i.e. before the mixture stream enters the heat exchange reformer, or inside the heat exchange reformer, e.g. by adding steam to the hot effluent synthesis gas at a given height within the reactor, specifically at a position corresponding to 50% or more of the height of the heat exchange reformer. At such heights the risk of metal dusting is often the highest.

Preferably, the stream comprising steam comprises only steam. Preferably the temperature of said hot effluent synthesis gas is 950-1050°C, more preferably 1025°C while the steam added is preferably 271°C at 55 barg thus resulting in a temperature of the mixture stream of 900-990°C. Hence, according to the invention and contrary to

conventional thinking, the addition of water in the form of steam is (i) delayed till after the process gas being reformed has passed the ATR, CPO or POX, and (ii) the steam is added only on the non-catalytic side of the heat exchange reformer. The water is not mixed with the process gas which in turn has to undergo reforming reactions, but rather the water is added later in the process where it turns out to have a much better effect. Typically, steam is added upstream the heat exchange reformer. In a sense, water is borrowed for the process without it being used in reforming reactions. As a result, the heat exchange reactor and downstream first waste heat boiler become modestly larger, but this is compensated by the huge savings in plant equipment size in the reforming section resulting from the low amount of steam in the process gas being reformed, and which is reflected by low process and total steam-to-carbon ratios as defined below.

More specifically, such steam addition downstream the ATR results in unchanged equipment volume upstream the point of such steam addition (quench point) and a significantly smaller flow throughput and equipment size downstream the point of steam addition, i.e. heat exchange reformer shell side and subsequent cooling train.

2. Process according to feature 1 further comprising a step before the step (a) of:

reforming the hydrocarbon feedstock by adding steam to said hydrocarbon feedstock to form a

hydrocarbon-steam mixture, and passing the hydrocarbon- steam mixture through at least one adiabatic pre- reforming stage.

Adiabatic pre-reforming is preferably conducted in a fixed bed reactor containing a reforming catalyst, thereby converting all higher hydrocarbons into a mixture of carbon oxides, hydrogen and methane. This endothermic process is accompanied by the equilibration of exothermic methanation and shift reactions. Removal of higher hydrocarbons allows a higher preheat temperature to the subsequent steam reforming.

3. Process according to features 1 or 2 , wherein the ratio of steam in the mixture stream is in the range of 0.02 to 0.13 kg pr Nm³ hot effluent synthesis gas. Preferably the range is 0.05 to 0.10 kg/Nm³ hot effluent synthesis gas, for instance 0.06, 0.07, 0.08, 0.09 kg/Nm³ hot effluent synthesis gas. This is also represented in the application as quench ratio, RAT, as illustrated in the Examples. . Process according to any of the features 1 - 3 , wherein the process comprises a further step after the mixture stream has passed through the heat exchange reformer of: cooling the mixture stream to a temperature sufficiently low for the steam to condensate, and

separating the synthesis gas from said condensate. By the term "temperature sufficiently low for the steam to condensate" is meant a temperature lower than the dew- point temperature for the steam at the pressure in the system at the position in question. This is expediently effected by further cooling the synthesis gas in cooling train and separation units to a temperature of e.g. 40-

70°C, and separating the thus obtained synthesis gas from the condensate, i.e. water and dissolved gases. The sensible heat may be used for steam and preheating purposes .

5. Process according to any of the features 1 - 4 further comprising mixing the primary reformed gas with tail gas from a Fischer-Tropsch synthesis stage. The addition of such tail gas to the synthesis gas production section enables that there is sufficient carbon dioxide during the reforming to achieve the desired H₂/CO molar ratio, typically about 2.0 for

Fischer-Tropsch synthesis.

As used herein "tail gas" means off-gas from the Fischer- Tropsch synthesis stage which is not re-used in said stage .

6. Process according to any of the features 1 - 4

comprising mixing the primary reformed gas with a gas stream comprising at least 90 vol% C0₂.

Thus, instead of using tail gas a separate stream rich in C0₂, i.e. with at least 90 vol% C0₂ may be added to the primary reformed gas .

7. Process according to any of features 1 to 6, wherein all the hot effluent synthesis gas from step (b) is used as heating medium in the heat exchange reformer of step (a) .

8. Process according to any of features 1 to 7 , wherein the heat exchange reformer is selected from a tube and shell heat exchanger, and double-tube reactor with catalyst disposed inside the double tubes, catalyst disposed outside the double tubes, and/or catalyst disposed outside and inside the double tubes.

9. Process according to any of features 1 to 7 , wherein the heat exchange reformer is a bayonet tube type reactor. Thus, in a particular embodiment of the invention the heat exchange reactor is a bayonet tube type reactor, in which at least one catalyst tube is provided in the form of an outer and an inner tube, the outer tube being a U- shaped tube and provided with a reforming catalyst, the inner tube being adapted concentrically to withdraw an effluent stream of partly reformed hydrocarbon from the outer tube, the outer tube being concentrically

surrounded by a sleeve spaced apart the outer tube and being adapted to pass the mixture stream comprising the hot effluent synthesis gas in indirect heat exchange relationship with reacting feedstock in the outer tube by conducting said mixture stream in the space between the sleeve and the outer tube.

By the term "catalyst tube" is meant a tube filled with particulate catalyst thereby forming a fixed bed, or a tube in which the catalyst is adhered as coating or coated in a foil adapted to the inner perimeter of the tube, or a tube in which the catalyst is coated or impregnated on structural elements such as straight channel or cross-corrugated monoliths adapted within the tube . 10. Process according to feature 4 further

comprising converting the separated synthesis gas into ammonia synthesis gas, methanol synthesis gas, dimethyl ether (D E) synthesis gas, synthesis gas for production of hydrocarbons by Fischer-Tropsch synthesis, or

synthesis gas for the production of hydrogen, and further converting said synthesis gas into the respective product in the form of ammonia, methanol, DME, liquid

hydrocarbons , or hydrogen.

11. Process according to any of the features 1 to 10, wherein a portion of the hydrocarbon feedstock is led directly as a by-pass stream to the ATR, CPO or POX.

In a particular embodiment, tail gas is added to this bypass stream.

12. Process according to any of features 1 to 10 wherein a portion of the hydrocarbon feedstock is led directly as a by-pass stream to the primary reformed gas to form a combined stream before entering the ATR, CPO or POX.

In a particular embodiment, tail gas is added to this bypass stream and the thereby formed stream is combined with the primary reformed gas . 13. Process according to any of features 1 to 12, wherein the process steam-to-carbon ratio (S/C_pr0cess) is in the range 0.4-1.2. Preferably this steam-to-carbon ratio is in the range 0.55-0.90, more preferably 0.60- 0.80. At such low process steam-to-carbon ratios best results in terms of metal dusting are obtained.

Particularly at 0.60-0.80, more particularly at about

0.60, metal dusting in the heat exchange reformer is prevented and without increasing plant equipment size. The process steam-to-carbon ratio, S/C_pr0cess_/ means the number of moles steam added to the process divided by the number of moles of hydrocarbon carbon added to the process. The number of moles of steam includes all the steam added to the feedstock, e.g. natural gas, upstream any or all the reforming reactors. Steam added as quench following the ATR, CPO or POx is not included. The hydrocarbon carbon includes both the hydrocarbons present in the feedstock and the hydrocarbons, including methane, ethane, propane, olefins etc. present in e.g. the tail gas from the Fischer-Tropsch synthesis stage. The total steam-to-carbon ratio S/C_tot/ is preferably in the range 0.80-1.30, more preferably 0.90-1.25.

The total steam-to-carbon ratio, S/C_tot/ means number of moles steam added to the process divided by the number of moles of hydrocarbon carbon added to the process. The number of moles of steam includes both the steam added to the feedstock, e.g. natural gas upstream any or all the reforming reactors and steam added downstream the ATR, CPO or POx i.e. to the hot effluent gas from the ATR, CPO or POx. The hydrocarbon carbon includes both the

hydrocarbons present in the feedstock and the

hydrocarbons, including methane, ethane, propane, olefins etc. present in e.g. the tail gas from the Fischer- Tropsch synthesis stage.

By the invention, S/C_tot can be kept at a lower value compared to a conventional situation where steam is not added to the hot effluent gas from the ATR, CPO or POx. As a result it is now possible to reduce the equipment size in the plant with respect to conventional processes, while at the same time it is now also possible to

increase the quality of the synthesis gas product (product purity, X_H2₊co_/ see Examples) . Hence, compared to conventional processes, not only is metal dusting

mitigated, but plant equipment size is reduced and synthesis gas quality for downstream applications, particularly Fischer-Troposch synthesis, is significantly enhanced.

14. Process according to any of the features 1 to 13, wherein the H₂/C0-molar ratio in said stream of cooled synthesis gas in step (a) is in the range of 1.7 to 2.3.

This is the H₂/C0-molar ratio of the synthesis gas product used in downstream processes, such as Fischer- Tropsch synthesis, ammonia synthesis, methanol synthesis, DME synthesis, and hydrogen synthesis. The ratio can be tailored accordingly depending on the downstream process, for instance H₂/CO-molar ratio of about 2 for Fischer- Tropsch synthesis.

15. Plant for the production of synthesis gas from a hydrocarbon feedstock, comprising an arrangement of:

• a heat exchange reformer for producing a primary

reformed gas by reforming of said hydrocarbon feedstock in indirect heat exchange with a mixture stream comprising hot effluent synthesis gas,

• an authothermal reformer (ATR) , a catalytic partial oxidation apparatus (CPO) , or a partial oxidation apparatus (POX) for producing said hot effluent synthesis gas from at least a portion of the primary reformed gas from said heat exchange reformer, means for adding steam to said hot effluent

synthesis gas thereby forming said mixture stream, means downstream the heat exchange reformer for condensating steam from said mixture stream thereby forming a steam condensate,

means for separating synthesis gas from the steam condensate .

16. Plant according to feature 15 further comprising an adiabatic pre-reformer upstream the heat exchange reformer .

The invention is further illustrated by reference to the attached drawings Figs. - 1-4 which are described in the following and each showing one particular embodiment of the invention. Fig. 5 shows one general embodiment of the invention including a pre-reformer.

In Fig. 1, a mixture of hydrocarbon containing feedstock and steam 10 is passed to the heat-exchange reformer 25 where it is catalytically steam reformed and thereafter leaves the heat-exchange reformer as stream 30. The primary reformed gas 30 is fed to an autothermal reformer 75 to which oxidant 80 is also supplied. The primary reformed gas is partially combusted and brought towards equilibrium over reforming catalyst in the autothermal reformer 75. The hot effluent synthesis gas 90 from the autothermal reformer 75 is mixed with a H₂0 stream 100 to form what is here defined as mixture stream comprising hot effluent synthesis gas 110. Heat is recovered from this mixture stream 110 by passing the mixture 110 to the heat exchange reformer 25. The mixture stream is cooled by heat exchange with the gas undergoing reforming over the catalyst in the heat-exchange reformer 25. The thus cooled synthesis gas leaves the heat exchange reformer as stream 120. The mixture stream 120 now as cooled

synthesis gas is further cooled in the cooling train and separation units 125 and separated into the product synthesis gas 130 and process condensate 140.

In Fig. 2, a mixture of hydrocarbon containing feedstock and steam 10 is passed to the heat-exchange reformer 25 where it is catalytically steam reformed and thereafter leaves the heat-exchange reformer as stream 30. The primary reformed gas stream 30 is mixed with Fischer- Tropsch tail gas 60 forming the ATR feed stream 70. The ATR feed stream 70 is fed to an autothermal reformer 75 to which oxidant 80 is also supplied. The ATR feed stream is partially combusted and brought towards equilibrium over reforming catalyst in the autothermal reformer 75. The hot effluent synthesis gas 90 from the autothermal reformer 75 is mixed with a H₂0 stream 100 to form what is here defined as mixture stream comprising hot effluent synthesis gas 110. Heat is recovered from this mixture stream 110 by passing the mixture 110 to the heat

exchange reformer 25. The mixture stream is cooled by heat exchange with the gas undergoing reforming over the catalyst in the heat-exchange reformer 25. The thus cooled synthesis gas leaves the heat exchange reformer as stream 120. The mixture stream 120 now as cooled

synthesis gas is further cooled in the cooling train and separation units 125 and separated into the product synthesis gas 130 and process condensate 140. In Fig. 3, a mixture of hydrocarbon containing feedstock and steam 10 is divided into two streams 20 and 40. The first stream 20 is fed to the heat-exchange reformer 25 where it is catalytically steam reformed and thereafter leaves the heat-exchange reformer as primary reformed gas 30. The second stream 40 is preheated in a heat exchanger 45 and bypasses the heat exchange reformer. The primary reformed gas 30 is mixed with the preheated second stream 50 forming the ATR feed stream 70. The ATR feed stream is fed to the autothermal reformer 75 to which oxidant 80 is also supplied. The ATR feed stream is partially combusted and brought towards equilibrium over reforming catalyst in the autothermal reformer 75. The hot effluent

synthesis gas 90 from the autothermal reformer 75 is mixed with a H₂0 stream 100 to form what is here defined as mixture stream comprising hot effluent synthesis gas 110. Heat is recovered from this mixture stream 110 by passing the mixture 110 to the heat exchange reformer 25. The mixture stream is cooled by heat exchange with the gas undergoing reforming over the catalyst in the heat- exchange reformer 25. The thus cooled synthesis gas leaves the heat exchange reformer as stream 120. The mixture stream 120 now as cooled synthesis gas is further cooled in the cooling train and separation units 125 and separated into the product synthesis gas 130 and process condensate 140.

In Fig. 4, a mixture of hydrocarbon containing feedstock and steam 10 is divided into two streams 20 and 40. The first stream 20 is fed to the heat-exchange reformer 25 where it is catalytically steam reformed and thereafter leaves the heat-exchange reformer as primary reformed gas 30. The second stream 40 is preheated in a heat exchanger 45. The preheated stream 45 is mixed with Fischer-Tropsch tail gas 60 forming stream 65. The primary reformed gas 30 is mixed with the mixture of Fischer-Tropsch tail gas and preheated second stream 65 forming the ATR feed stream 70. The ATR feed stream is fed to the autothermal reformer 75 to which oxidant 80 is also supplied. The ATR feed stream is partially combusted and brought towards equilibrium over reforming catalyst in the autothermal reformer 75. The hot effluent synthesis gas 90 from the autothermal reformer 75 is mixed with a H₂0 stream 100 to form what is here defined as mixture stream comprising hot effluent synthesis gas 110. Heat is recovered from this mixture stream 110 by passing the mixture 110 to the heat exchange reformer 25. The mixture stream is cooled by heat exchange with the gas undergoing reforming over the catalyst in the heat-exchange reformer 25. The thus cooled synthesis gas leaves the heat exchange reformer as stream 120. The mixture stream 120 now as cooled

The accompanying Fig. 5 shows a general embodiment for production of synthesis gas for Fischer-Tropsch synthesis using an autothermal reformer in series with a heat exchange reformer, as in Fig. 2. Clean (hydrogenated and free of sulphur and other poisons to reforming catalysts) hydrocarbon feed gas 10 such as natural gas or other hydrocarbon containing gas source is mixed with process steam 81, optional partly via saturator/humidifier . The mixture 11 is preheated and pre-reformed adiabatically in pre-reformer 30 in order to convert higher hydrocarbons into H_2/ CO, C0₂ and CH₄. This mixture 12 is fed to the catalyst side of the heat exchange reformer 25, where it is reformed substantially to equilibrium at the outlet temperature. The outlet temperature is often in the range 600°C-850°C, more often between 675°C and 775°C.

Downstream the catalyst side of the heat exchange

reformer 25 the primary reformed gas 30 is mixed with tail gas 60 (containing unconverted hydrocarbons (CH₄, higher HC olefins etc.) and other gases (H₂, CO, C0₂, N₂ etc.) from the FT-synthesis . In a front-end for e.g. a methanol plant there is no tail gas. This process gas mixture, oxygen 80 and protection steam 81 is fed to the ATR 75 where it is partly combusted and further

catalytically reformed to equilibrium.

A stream 100 comprising steam (quench) is added to the hot effluent synthesis gas 90 from the ATR. The quenched synthesis gas 110, defined herein as mixture stream comprising hot effluent synthesis gas, is led to the heat exchange reactor 25 where it delivers heat for the reforming reaction on the catalyst side by heat exchange. Finally the cooled synthesis gas 120 leaving the heat exchange reformer 25 is further cooled in cooling train and separation units 125 to a low temperature, e.g. 40- 70°C, and separated as synthesis gas 130 from the

condensate, i.e. water and dissolved gases. The sensible heat may be used for steam and preheating purposes . EXAMPLES

In the following Examples the compositions of natural gas and tail gas are as follows:

Table 1

In all the Examples it is assumed that the gas is in chemical equilibrium at the outlet of the catalyst side of heat exchange reformer and at the outlet of the autothermal reformer . In the tables, the following ratio is used to assess the severity of the gas entering the heat exchange reformer after passing through the autothermal reforming stage and after steam addition (stream 110 in Figs. 1-4)

P_H2o: Partial pressure of steam (atm abs)

Pco: Partial pressure of carbon monoxide (atm abs)

P_H2: Partial pressure of hydrogen (atm abs) It is an advantage for avoiding or reducing the rate of metal dusting, that the value of Q is as high as

possible .

Alternatively (see Example 17) the equilibrium

temperature for CO reduction ( _CQ._RED) in the actual gas is used as a comparable severity factor for the risk of metal dusting. Higher T_CO-_RED indicates higher potential for MD. As already described above the process steam-to-carbon ratio, S Cprocess / means number of moles steam added to the process divided by the number of moles of hydrocarbon carbon added to the process. The number of moles of steam includes all the steam added to the feedstock (here natural gas) upstream any or all the reforming reactors. Steam added as quench following the ATR is not included. The hydrocarbon carbon includes both the hydrocarbons present in the feedstock and the hydrocarbons including methane, ethane, propane, olefins etc. present in the tail gas from the Fischer-Tropsch synthesis stage. The total steam-to-carbon ratio, S/C_tot/ means number of moles steam added to the process divided by the number of moles of hydrocarbon carbon added to the process . The number of moles of steam includes both the steam added to the feedstock (here natural gas) upstream any or all the reforming reactors and the steam added downstream the ATR. The hydrocarbon carbon includes both the

hydrocarbons present in the feedstock and the

hydrocarbons including methane, ethane, propane, olefins etc. present in the tail gas from the Fischer-Tropsch synthesis stage.

Other definitions: Quench ratio, RAT: kg H₂0 in quench (stream 100 in Figs.

1-4) /Nm³ hot effluent gas leaving ATR (stream 90 in Figs. 1-4) .

Specific inlet feed flow, F_in/CO: relative values of Nm³ process gas (stream 10 in Figs. 1-4) /Nm³ CO in the synthesis gas product (stream 130 in Figs. 1-4) .

Specific raw product flow, F_ATR,_out/CO : relative values of Nm³ hot reformed synthesis gas after quench point (stream 110 in Figs. 1-4)/ Nm³ CO in the synthesis gas product (stream 130 in Figs. 1-4) .

Product purity, X_H2+co: dry mole % of H₂+CO in product synthesis gas (stream 130 in Figs. 1-4) .

H₂/CO ratio: means the number of moles of hydrogen divided by the number of moles of CO in a stream. Example 1 (Comparative Example)

In this Example no steam is added to the effluent stream from the autothermal reformer.

Natural gas mixed with steam is sent to the heat exchange reformer. A primary reformed gas is withdrawn from the heat exchange reformer at 7 50°C. This stream is mixed with tail gas from the Fischer-Tropsch synthesis unit.

The tail gas amount is adjusted to give an H₂/CO-ratio at the outlet of the autothermal reformer of 2 . 0 0 by volume.

The resulting gas mixture and oxygen are added to the autothermal Reformer. The amount of oxygen is adjusted to give the desired exit temperature from the autothermal reformer. The hot effluent gas leaving the autothermal reformer is passed to the shell side of the heat exchange reformer without addition of steam. Base line S/C_process = 0 . 6 .

Example 2

The process in this Example is similar to the process described in Example 1 . The only difference is that steam is added to the hot effluent gas leaving the autothermal reformer .

The results of the calculations for Examples 1 & 2 are provided in Table 2 : Table 2: Results for Examples 1 and 2

It is seen that increased Q-values (reduced potential for metal dusting) can be obtained by increasing the S/C_pr0Cess ratios without steam addition downstream the ATR (RAT=0) , cases la, lb, and lc. Corresponding increased Q-values can be obtained by keeping the S/C_process constant (e.g. 0.6 in this Example) while adding steam to the reformed gas downstream the ATR ( RAT>0 ) , cases 2b and 2c. In cases lb and lc the flow throughput of the plant (Fi_n/CO and FA R,out CO) increases significantly resulting in larger plant equipment. In addition the product quality (X_H2+co) decreases. Increasing the Q-value by steam addition downstream the ATR (cases 2b and 2c) results in unchanged equipment volume upstream the quench point (F_in/CO is constant) and a significantly smaller flow throughput and equipment downstream the quench point (heat exchange reformer shell side and cooling train) . It is furthermore seen that the concentration of hydrogen and carbon monoxide in cases 2b and 2c are the same as in case la and higher than in cases lb and lc. It is thus clear that H₂0 from the quench stream added downstream the ATR can replace the H₂0 from a higher process S/C ratio with respect to preventing metal dusting corrosion, but with considerably less impact on the plant, and while at the same time maintaining high product quality. Example 3 (Comparative Example)

The process in this Example is similar to the process described in Example 1. This Example differs from Example 1 by having a higher S/C_pr0cess ( =0.8) as base line. Example 4

The process in this Example is similar to the process described in Example 3. The only difference is that steam is added to the hot effluent gas leaving the autothermal reformer .

The results of the calculations for Examples 3 & 4 are provided in Table 3 :

Table 3: Results for Examples 3 and 4

It is seen that the advantages when comparing Examples 3 and 4 are analogous to comparison between Example 1 and 2. Example 5 (Comparative Example)

The process in this Example is similar to the process described in Example 1. This Example differs from Example 1 by the temperature (= 700 ^aC) of the primary reformed gas withdrawn from the heat exchange reformer (stream 30 in Figs . 1-4) .

Example 6

The process in this Example is similar to the process described in Example 5. The only difference is that steam is added to the hot effluent gas leaving the autothermal reformer .

The results of the calculations for Examples 5 & 6 are provided in Table 4:

Table 4: Results for Examples 5 and 6

It is seen that the results from comparing Examples 5 and 6 are analogous to comparison between Example 1 and 2.

Example 7 (Comparative Example)

The process in this Example is similar to the process described in Example 1. This Example differs from Example 1 by the H₂/CO-ratio (1.8 by volume) at the outlet of the autothermal reformer, i.e. hot effluent gas leaving the autothermal reformer (stream 90 in Figs. 1-4).

Example 8

The process in this Example is similar to the process described in Example 7. The only difference is that steam is added to the hot effluent gas leaving the autothermal reformer .

The results of the calculations for Examples 7 & 8 are provided in Table 5 :

Table 5: Results for Examples 7 and 8

It is seen that the results from comparing Examples 7 and 8 are analogous to comparison between Example 1 and 2. Example 9 (Comparative Example)

The process in this Example is similar to the process described in Example 1. This Example differs from Example 1 by the H₂/CO-ratio (2.2 by volume) at the outlet of the autothermal reformer; i.e. hot effluent gas leaving the autothermal reformer (stream 90 in Figs. 1-4). Example 10

The process in this Example is similar to the process described in Example 9. The only difference is that steam is added to the hot effluent gas leaving the autothermal reformer .

The results of the calculations for Examples 9 & 10 are provided in Table 6 :

Table 6: Results for Examples 9 and 10

It is seen that the results from comparing Examples 9 and 10 are analogous to comparison between Example 1 and 2.

Example 11 (Comparative Example)

In this Example no steam is added to the hot effluent gas from the autothermal reformer. A stream of natural gas mixed with steam is divided into two streams. The first stream (85% of the volumetric flow) is sent to the heat- exchange reformer. The remaining 15% of the volumetric flow (stream 40 in Fig. 3) bypasses the heat-exchange reformer catalyst side. A primary reformed gas is

withdrawn from the heat-exchange reformer at 750°C. The by-pass stream is heated to 620 ^aC and thereafter mixed with tail gas from the Fischer-Tropsch synthesis unit. The hot by-pass gas/tail gas mixture (stream 65 in Fig. 4) is mixed with the primary reformed gas from the heat- exchange reformer. The resulting gas mixture (stream 70 in Fig. 4) and oxygen are added to the autothermal reformer. The tail gas amount is adjusted to give an H₂/CO-ratio at the outlet of the autothermal reformer of 2.00 by volume .

The amount of oxygen is adjusted to give the desired exit temperature from the autothermal reformer. The gas leaving the autothermal reformer is passed to the shell side of the heat exchange reformer without addition of steam. Base line S/C_process = 0.6.

Example 12

The process in this Example is similar to the process described in Example 11. The only difference is that steam is added to the hot effluent gas leaving the autothermal reformer.

The results of the calculations for Examples 11 & 12 are provided in Table 7 :

Table 7: Results for Examples 11 and 12

Case S/C_process s/c_tot RAT Q Fi_n/CO FATR.OU./CO XH2+ (mol/mol ) (mol/mol ) relative relative (Dry mol%)

11a 0.6 0.6 0 0.0400 1.000 1.000 88.44 lib 1.0 1.0 0 0.0706 1.163 1.119 85.15

12b 0.6 0.96 0.074 0.0706 1.000 1.092 88.44 It is seen that the results from comparing Examples 11 and 12 are analogous to comparison between Example 1 and 2 .

Example 13 (Comparative Example)

In this Example no steam is added to the hot effluent gas from the autothermal reformer. The Example is analogous to Example 1 except no tail gas is added, and therefore the H₂/CO ratio in the hot effluent gas leaving the autothermal reformer is not adjusted. Natural gas mixed with steam is sent to the heat exchange reformer. A primary reformed gas is withdrawn from the heat exchange reformer at 7 00°C. The primary reformed gas and oxygen are added to the autothermal Reformer . The amount of oxygen is adjusted to give the desired exit temperature from the autothermal reformer. The hot effluent gas leaving the autothermal reformer is passed to the shell side of the heat exchange reformer without addition of steam. Base line S /Cp_roCess = 0 . 6 . Example 14

The process in this Example is similar to the process described in Example 13. The only difference is that steam is added to the hot effluent gas leaving the autothermal reformer .

The results of the calculations for Examples 13 & 14 are provided in Table 8 : Table 8: Results for Examples 13 and

It is seen that increased Q-values (reduced potential for metal dusting) can be obtained by increasing the S/C_pr0cess ratios without quenching with steam (RAT=0 ) , cases 13a, 13b, and 13c. Corresponding increased Q-values can be obtained by keeping the S/C_pr0cess constant (e.g. 0.6 in this Example) while adding steam quench to the reformed gas (RAT>0 ) , cases 14b and 14c. In cases 13b and 13c the flow throughput of the plant (Fi_n/CO and F_ATR,_ou /CO) increases significantly resulting in larger plant

equipment. In addition the product quality (X_H2₊co)

decreases. Increasing the Q-value by quenching (cases 14b and 14c) results in unchanged equipment volume upstream the quench point (Fi_n/CO is constant) and significantly smaller flow throughput and equipment downstream the quench point (heat exchange reformer shell side and cooling train) . It is furthermore seen that the

concentration of hydrogen and carbon monoxide in cases 2b and 2c are the same as in case la and higher than in cases lb and lc. Example 15 (Comparative Example)

In this Example no steam is added to the effluent stream from the autothermal reformer . The Example is analogous to Example 11 except that no tail gas is added, and therefore the H₂/CO-ratio outlet the autothermal reformer is not adjusted.

A stream of natural gas mixed with steam is divided into two streams. The first stream (85% of the volumetric flow) is sent to the heat exchange reformer. The

remaining 15% of the volumetric flow (stream 40 in Fig. 3) bypasses the heat exchange reformer catalyst side. A primary reformed gas is withdrawn from the heat exchange reformer at 750°C. The bypass stream is heated to 620 ^eC and thereafter mixed with the primary reformed gas from the heat exchange reformer. The resulting gas mixture (stream 70 in Fig. 3) and oxygen are added to the

autothermal Reformer. The amount of oxygen is adjusted to give the desired exit temperature from the autothermal reformer. The gas leaving the autothermal reformer is passed to the shell side of the heat exchange reformer without addition of steam. Base line S/C_process = 0.6. Example 16

The process in this Example is similar to the process described in Example 15. The only difference is that steam is added to the hot effluent gas leaving the autothermal reformer.

The results of the calculations for Examples 15 & 16 are provided in Table 9 : Table 9: Results for Examples 15 and

It is seen that increased Q-values (reduced potential for metal dusting) can be obtained by increasing the S/C_proCess ratios without quenching ( RAT=0) , cases 15a and 15b.

Corresponding increased Q-values can be obtained by keeping the S/C_proCess constant (e.g. 0.6 in this Example) while adding steam quench to the reformed gas ( RAT>0) , case 18b. In case 15b the flow throughput of the plant ( Fin/CO and FA out ) increases significantly resulting in larger plant equipment. The product quality (X_H2+co) s slightly decreased. Increasing the Q-value by quenching with steam (case 16b) results in unchanged equipment volume upstream the quench point ( Fi_n/CO is constant) and significantly smaller flow throughput and equipment downstream the quench point (heat exchange reformer shell side and cooling train) . The product composition is unchanged.

Example 17

Series arrangement of adiabatic-prereformer, heat

exchange reformer and autothermal reformer (Fig. 5) For fixed feed gas flow and compositions of the feed- and tail-gases, the main process variables for obtaining a specific product synthesis gas H₂/CO ratio are:

• Overall S/C ratio

• Heat exchange reactor catalyst exit temperature

• ATR exit temperature (temperature of hot effluent gas from ATR)

• Tail gas flow (Fischer-Tropsch front-end only)

The selected values of these variables impose the

required amount of oxygen to the ATR (in analogy to this a scenario with fixed oxygen flow would impose the feed gas flow) .

Fixing the catalyst exit temperatures (high catalyst exit temperatures are selected because this gives a higher conversion) leaves two main variables: the S/C ratio and the tail gas flow for controlling the product composition (H₂/CO ratio) . A higher S/C ratio demands a higher tail gas flow. The resulting volumetric flow through the plant will also be higher. This is illustrated in Table 10 below for a typical grassroots Gas-to-Liquids (GTL) layout based on series arrangement of heat exchange reformer and ATR:

Table 10: Relative values of specific tail gas import and total gas flow pr unit CO in product for varying S/C ratios * ( S /C ) process 0.6 0.8 1.0 1.25 m^J Ftail gas/Nm^J

COprod / relative % 100 123.4 145.5 173.0

Nm^J Fexit ATR/ itr*

COprod relative % 100 105.5 111.1 118.2

* Fixed natural gas (NG) feed flow, typical (fixed) NG & tail gas (TG) composition, T_exit Heat exchange reformer cat = 750^aC, T_exit ATR (hot effluent gas from ATR) =

1025 ^aC, product H₂/CO ratio = 1.925, and no quench steam

Increasing the (S/C)_pr0cess e.g. from 0.6 to 1.0 requires an increase in tail gas flow (and the equipment volume which handles the tail gas- recycle compressor, heating, piping etc.) of 45% pr unit CO produced when maintaining the product H₂/CO ratio. The resulting volumetric flow of the hot effluent gas from the ATR (Nm³ F_exit ATR/ΝΠΙ³ C0_prOd, relative %) increases by 11% and so does the gas volume handled by the downstream equipment. The gas volume will also increase considerable between the steam addition point and the tail gas addition point. An increased (S/C)process therefore has considerable effect on the equipment size and therefore cost of the plant.

Critical process characteristics:

The Fischer-Tropsch synthesis requires a synthesis gas with an H₂/C0 ratio of about 2.0 (1.9 - 2.1). In

accordance with the above it is generally accepted that it is economically attractive to operate the process with high catalyst exit temperatures and a low (S/C)_process of around 0.6. However, this results in rather aggressive conditions (gas composition and temperature) with respect to risk of metal dusting (MD) of the piping and equipment downstream the ATR reactor.

Significance of the quench:

As is also shown below, H₂0 from the quench stream can replace the H₂0 from a higher (S/C) _proCess with respect to preventing metal dusting corrosion, but with considerably less impact on the plant.

Table 11 below illustrates the impact on T_CO-_RED of

(S/C) process and of quench. The relative impact on the exit flow from the ATR after the quench point (and thereby on the cooling train) is also included:

Table 11: Impact of overall S/C ratio and quench on T_Co red and wet synthesis gas flow for a GTL plant with H₂/CO ratio = 1.925

Parameter \ Base 1A IB 2A 2B Case No No No Quench Quench quench quench quench

(S/C) process 0.6 1.0 1.25 0.6 0.6

Quench ratio, 0 0 0 0.068 0.111 kg/Nm³ ATR_exi_t-

2)

flow

TCO-RED after 891 848 828 848 828 quench , ^SC

Process gas 100 114.9 124.8 100 100 flow³' /unit CO

prod, Relative

%

Wet syngas 100 111.1 118.2 108.4 113.8 flow⁴' /unit CO

prod, Relative

%

1) Including H₂0 taking part in the reactions, and excluding steam from quench

2) Hot effluent gas from ATR before steam addition

(quench)

3) Volumetric process gas flow downstream main process steam addition inlet the first reforming reactor, here an adiabatic pre-reformer .

4) Volumetric synthesis gas flow downstream the ATR reactor after the quench point

The Figs, in Table 11 show that e.g. a 43 ^SC reduction in T_CO-_RED from 891 ^SC in the base case to 848 ^AC (case 1A) requires an increase in overall S/C ratio from 0.6 to 1.0. Keeping the H₂/CO ratio at 1.925 this results in a 15% increase in the volumetric process gas flow per unit of produced CO inlet the first reforming reactor

(prereformer) and 11% increase in the wet synthesis gas flow per unit of produced CO downstream the ATR and throughout the cooling train.

The same reduction in T_CO-_RED (case 2A) would require addition of 0.068 kg quench steam per Nm³ ATR exit flow. This corresponds to increasing the volume flow exit the ATR by 8.4% (the quench corresponds to 7.8% of the resulting synthesis gas flow) . As the dry gas composition is equal in case 2A and 1A and in CO production in particular is unchanged, the resulting specific wet synthesis gas flow (per unit of produced CO) would increase by 8.4%, thus 2.7% less than in case 1A. The process gas flow upstream the quench point would remain unchanged and thus be approx 15% lower than in case Al . This has a significant impact on the size and cost of the equipment.

Further reduction in the T_CO-_RED would result in an even larger benefit from using quench in stead of increased (S/C) process · This illustrated by cases 2A and 2B where Τ_∞- R_ED is reduced by 63 ^aC compared to the base case.

Introducing a steam quench downstream the ATR instead of increasing the (S/C) _pr0cess has the following main

advantages :

· The potential for metal dusting corrosion on the

shell side of the heat exchange reactor represented by the T_CO-RED is reduced The added H₂0 does not take part in the reactions, and the synthesis gas composition is unchanged The volumetric flows in the reforming section are significantly reduced; e.g. for a change in metal dusting potential corresponding to increasing the (S/C) process from 0.6 -> 1.0 the volume reductions in the calculation Example are:

o 15% in the adiabatic pre-reformer and HTER-s catalyst

o 2.7% in all the equipment downstream the ATR

Claims

(a) reforming said hydrocarbon feedstock in an

exchange reformer an effluent stream of primary reformed gas ;

(b) passing the primary reformed gas from the heat

exchange reformer through an autothermal reforming stage (ATR) , catalytic partial oxidation stage (CPO) or partial oxidation stage (POX) , and withdrawing a stream of hot effluent synthesis gas of which at least a portion is used as heating medium in the heat exchange reformer of step (a) ;

characterised in that the process further comprises forming said mixture stream comprising hot effluent synthesis gas from step (b) by adding a stream comprising steam to at least a portion of said hot effluent

synthesis gas.

2. Process according to claim 1 further comprising a step before the step (a) of:

- reforming the hydrocarbon feedstock by adding steam to said hydrocarbon feedstock to form a

3. Process according to claims 1 or 2 , wherein the ratio of steam in the mixture stream is in the range of 0.02 to 0.13 kg pr Nm³ hot effluent synthesis gas.

4. Process according to any of the claims 1 - 3, wherein the process comprises a further step after the mixture stream has passed through the heat exchange reformer of: cooling the mixture stream to a temperature sufficiently low for the steam to condensate, and separating the synthesis gas from said condensate.

5. Process according to any of the claims 1 - 4 further comprising mixing the primary reformed gas with tail gas from a Fischer-Tropsch synthesis stage.

6. Process according to any of the claims 1 - 4

7. Process according to any of claims 1 to 6, wherein all the hot effluent synthesis gas from step (b) is used as heating medium in the heat exchange reformer of step (a) .

8. Process according to any of claims 1 to 7 , wherein the heat exchange reformer is selected from a tube and shell heat exchanger, and double-tube reactor with catalyst disposed inside the double tubes, catalyst disposed outside the double tubes, and/or catalyst disposed outside and inside the double tubes.

9. Process according to any of claims 1 to 7,

wherein the heat exchange reformer is a bayonet tube type reactor .

10. Process according to claim 4 further comprising converting the separated synthesis gas into ammonia synthesis gas, methanol synthesis gas, dimethyl ether (DME) synthesis gas, synthesis gas for production of hydrocarbons by Fischer-Tropsch synthesis, or synthesis gas for the production of hydrogen, and further

converting said synthesis gas into the respective product in the form of ammonia, methanol, DME, liquid

hydrocarbons , or hydrogen.

11. Process according to any of the claims 1 to 10, wherein a portion of the hydrocarbon feedstock is led directly as a by-pass stream to the ATR, CPO or POX.

12. Process according to any of claims 1 - 10 wherein a portion of the hydrocarbon feedstock is led directly as a by-pass stream to the primary reformed gas to form a combined stream before entering the ATR, CPO or POX.

13. Process according to any of claims 1 to 12, wherein the process steam-to-carbon ratio (S/C_pr0cess) is in the range 0.4-1.2, preferably 0.55-0.90, more

preferably 0.60-0.80.

14. Process according to any of the claims 1 to 13, wherein the H₂/CO-molar ratio in said stream of cooled synthesis gas in step (a) is in the range of 1.7 to 2.3.

- a heat exchange reformer for producing a primary reformed gas by reforming of said hydrocarbon feedstock in indirect heat exchange with a mixture stream

comprising hot effluent synthesis gas,

- an authothermal reformer (ATR) , or catalytic partial oxidation apparatus (CPO) or a partial oxidation

apparatus (POX) for producing said hot effluent synthesis gas from at least a portion of the primary reformed gas from said heat exchange reformer,

- means for adding steam to said hot effluent synthesis gas thereby forming said mixture stream,

- means downstream the heat exchange reformer for

condensating steam from said mixture stream thereby forming a steam condensate,

- means for separating synthesis gas from the steam condensate .

16. Plant according to claim 15 further comprising an adiabatic pre-reformer upstream the heat exchange reformer .