The present invention relates to a process for regenerating a Fischer-Tropsch catalyst which may be used in a process for treating natural gas to produce a liquid product.

It is well known that most oil wells also produce natural gas. At many oil wells natural gas is produced in relatively small quantities along with the oil. When the quantities of this associated gas are sufficiently large or the well is close to pre-existing gas transportation infrastructure, the gas can be transported to an off-site processing facility. When oil production takes place in more remote places it is difficult to introduce the associated gas into existing gas transportation infrastructure. In the absence of such infrastructure, the associated gas has typically been disposed of by flaring or re- injection. However, flaring the gas is no longer an environmentally acceptable approach, while re-injection can have a negative impact on the quality of the oil production from the field.

Gas-to-liquids technology can be used to convert the natural gas into liquid hydrocarbons and may follow a two-stage approach to hydrocarbon liquid production comprising syngas generation, followed by Fischer-Tropsch synthesis. In general, syngas (a mixture of hydrogen and carbon monoxide) may be generated by one or more of partial oxidation, auto-thermal reforming, or steam methane reforming. Where steam methane reforming is used, the reaction is endothermic and so requires heat, and a catalyst such as platinum/rhodium. The syngas is then subjected to Fischer- Tropsch synthesis. For performing Fischer-Tropsch synthesis the optimum ratio of hydrogen to carbon monoxide is about 2:1 , and steam reforming has a benefit of providing more than sufficient hydrogen for this purpose. As regards the Fischer- Tropsch process, a suitable catalyst uses cobalt on a ceramic support. Such a process is described for example in WO 01 / 51 194 (AEA Technology) and WO 03/048034 (Accentus pic).

Under some circumstances the Fischer-Tropsch catalyst may require regeneration. Methods of regeneration differ for different types of Fischer-Tropsch reactors. For example, in a conventional, large scale slurry bubble reactor, catalyst may be removed and replaced, and the spent catalyst may be regenerated remotely from the GTL plant and then subsequently returned thereto. Some smaller scale Fischer-Tropsch reactors have supported catalysts. These catalyst may be supported on the walls of the reactor structure, or alternatively, on dedicated supports that may be inserted into a reactor. However, even if the catalyst is not applied to a structural part of the reactor, removal and replacement of catalyst supports is time consuming and therefore a method whereby the catalyst could be regenerated while remaining in situ would provide clear advantages.

According to the present invention there is provided a process for regeneration of a Fischer-Tropsch catalyst in situ in a Fischer-Tropsch synthesis reactor for generating hydrocarbons from a synthesis gas containing hydrogen and carbon monoxide, the process comprising the following steps:

(a) while continuing to pass synthesis gas through the reactor to produce liquid hydrocarbons, increasing the hydrogen :CO ratio of the synthesis gas to greater than 2.0;

(b) then ensuring the temperature within the reactor is at least 220 °C;

(c) holding the reactor at these operating conditions for at least one day;

(d) then gradually changing the hydrogen:CO ratio, the pressure, and the temperature back to normal operating conditions. It will be appreciated that this regeneration process can be carried out in situ, and without modifying the associated plant. It has been found that this process is particularly suitable for regenerating catalyst that have been poisoned by ammonia, but it may also be effective for regenerating catalysts that have been deactivated in other ways, e.g. deactivation over time, deactivation by other poisons, deactivation by water, oxidation of active metal, interaction between active metal and support, fouling, or metal agglomeration/sintering.

A further advantage of the present invention is that the production of liquid products may continue under the regeneration conditions. The productivity of C5+ during regeneration may be as high as the productivity that had been obtained immediately before regeneration.

Prior to the step of increasing the hydrogen:CO ratio of the synthesis gas, the process may further include the step of gradually decreasing the pressure within the reactor to a lower level pressure which is less than 1 .0 MPa (gauge pressure). The process is particularly suitable for a compact catalytic Fischer-Tropsch reactor defining a multiplicity of first flow channels for the Fischer-Tropsch reaction arranged in proximity to a multiplicity of second flow channels for a heat exchange fluid, so there is heat exchange between the respective flow channels. A catalyst may be provided on the walls of the flow channels for the Fischer-Tropsch reaction, or alternatively each channel for the Fischer-Tropsch reaction contains a removable catalyst structure to catalyse the reaction. Each catalyst structure may comprise a metal substrate and incorporate an appropriate catalytic material. Each such catalyst structure may be shaped so as to subdivide the flow channel into a multiplicity of parallel flow sub-channels. Each catalyst structure may include a ceramic support material on the metal substrate, which provides a support for the catalyst.

The metal substrate provides strength to the catalyst structure and enhances thermal transfer by conduction. The metal substrate may be of a steel alloy that forms an adherent surface coating of aluminium oxide when heated, for example a ferritic steel alloy that incorporates aluminium (eg Fecralloy (TM)). The substrate may be a foil, a wire mesh or a felt sheet, which may be corrugated, dimpled or pleated; a suitable substrate is a thin metal foil for example of thickness typically between 50 μηι and 200 μηι, for example 100 μηι, which is corrugated to define the longitudinal sub-channels.

The reactor may comprise a stack of plates. For example the first and second flow channels may be defined by grooves in respective plates, the plates being stacked and then bonded together. Alternatively, the flow channels may be defined by thin metal sheets that are castellated and stacked alternately with flat sheets; the edges of the flow channels may be defined by sealing strips. The flow channels may instead be defined by flat metal sheets spaced apart by spacer strips. To ensure the required good thermal contact both the first and the second gas flow channels may be between 10 mm and 2 mm high (in cross-section); and each channel may be of width between about 3 mm and 25 mm. The stack of plates forming the reactor block is bonded together for example by diffusion bonding, brazing, or hot isostatic pressing.

In the optional pressure reduction step before step (a) of the regeneration process, the lower level pressure may be less than 0.6 MPa (gauge pressure), for example 0.5 MPa (g) or less, but would usually be above atmospheric pressure. The pressure is decreased gradually, preferably at no more than 20 kPa/min, for example at 10 kPa/min. The pressure may be decreased continuously, or stepwise; for example the pressure may be decreased stepwise by 100 kPa every ten minutes. However the reactor conditions such as temperature should be monitored during this process, and if a rapid change is detected, the pressure decrease should be held until the change ceases. Step (a) of the regeneration process normally requires an increase in the syngas

H ₂ /CO ratio, because a typical ratio during normal operation would be between 1 .9 and 2.0. This change should also be performed gradually, for example taking place over a period of at least 30 minutes, more preferably at least 2 hours. The ratio may be raised to 3,0 or even 3.5, for example, although higher ratios are also envisaged.

Step (b) may also require a change of temperature, although this depends on the normal operating temperature. In step (b) the temperature is preferably raised to a temperature between 220 °C and 235 ^q C, for example 225°C or 230 °C; and this is achieved by slowly changing the temperature for example at a rate of less than

0.5°C/hour, for example at 0.2 ^< O/hr.

Step (c) brings about regeneration of the catalyst, due to the low pressure, high temperature and high syngas H ₂ /CO ratio. This step may be performed for at least two days, for example for between 60 and 72 hours.

After performing step (c), and before performing step (d), it may be beneficial to perform a step of decreasing the temperature gradually down to between 200 ° and 215 ^q C, for example to 210°C. This would be performed gradually, for example at no more than 1 .0°C/min, for example at Ο.δ'ΌΛηίη.

The final step, step (d) involves returning the operating conditions to the normal values. This may be performed as successive steps: (d1 ) ensuring the syngas H ₂ /CO ratio has its normal value; (d2) then increasing the pressure back to the normal operating value, at no more than 20 kPa/min, for example at up to 10 kPa/minute; and (d3) then ensuring the temperature is at its normal operating value, changing the temperature gradually, if required, at a rate of less than O.S'O/hour, for example at 0.2°C/hr or less. In each of these steps the performance of the reactor is monitored, and if any rapid change is observed, the varying parameter (for example the pressure) is held or the temperature decreased, until the rapid change ceases.

In each case any changes may be performed continuously, or stepwise. For example when increasing the pressure, the pressure may be increased in steps of 100 kPa, and held for at least 10 minutes at each value.

The invention will now be further and more particularly described, by way of example only, and with reference to the accompanying drawings, in which:

Figure 1 shows a schematic flow diagram of a gas-to-liquid plant and associated equipment, including a Fischer-Tropsch reactor; and

Figure 2 shows a diagrammatic sectional view of a reactor block suitable for use in the Fischer-Tropsch reactor. 1 . Gas-to-Liquid Plant Overview

The invention is of relevance to a chemical plant and process for converting natural gas (primarily methane) to longer chain hydrocarbons. The plant is suitable for treating associated gas, which is natural gas that is produced along with crude oil, and is then separated from the crude oil. The first stage of the chemical process involves the formation of synthesis gas. This may be achieved for example by steam reforming, by a reaction of the type:

H ₂ 0 + CH ₄ → CO + 3 H ₂ (1 )

This reaction is endothermic, and may be catalysed by a rhodium or platinum/rhodium catalyst in a first gas flow channel. The heat required to cause this reaction may be provided by catalytic combustion of a gas such as methane or hydrogen, which is exothermic, in an adjacent channel, or by heat exchange with exhaust gases from a separate combustion reactor. The combustion may be catalysed by a palladium catalyst in an adjacent second gas flow channel in a compact catalytic reactor. In both cases the catalyst may be on a stabilised-alumina support which forms a coating typically less than 100 μηι thick on a metallic substrate. Alternatively, the catalyst may be applied to the walls of the flow channels or may be provided as pellets within the flow channel. The heat generated by the combustion would be conducted through the metal sheet separating the adjacent channels. As shown in equation (1 ) the resulting syngas H ₂ /CO ratio is 3, although the exact value depends on reactor conditions, and on the ratio of steam to methane provided to the reactor, and for example the ratio may be 3.5 if a higher proportion of steam is provided.

The gas mixture produced by the steam/methane reforming is then used to perform a Fischer-Tropsch synthesis to generate a longer chain hydrocarbon, that is to say: n CO + 2n H ₂ → (CH ₂ ) _n + n H ₂ 0 (2) which is an exothermic reaction, occurring at an elevated temperature, typically between ^~ \ 90 °C and 280 °C, for example 230 °C, and an elevated pressure typically between 1 .8 MPa and 2.7 MPa (absolute values), in the presence of a catalyst. Whilst Fe based catalysts can be used, metallic Co promoted with precious metals such as Pd, Pt, Ru or Re doped to 1 wt% are preferred when operating at lower temperatures as they have enhanced stability to oxidation. The active metals are impregnated to 10- 40 wt% into refractory support materials such as Ti0 ₂ , Al ₂ 0 ₃ or Si0 ₂ which may be doped with rare earth and transition metal oxides to improve their hydrothermal stability.

It will be appreciated from the equations above that, if steam/methane reforming is used to produce the synthesis gas, there is an excess of hydrogen. A hydrogen-rich gas stream can therefore be separated either from the synthesis gas stream before performing Fischer-Tropsch synthesis, or from the tail gases that remain after performing Fischer-Tropsch synthesis. Such a separation may use a membrane separator.

Referring to figure 1 , there is shown a gas-to-liquid plant 10 of the invention. A natural gas feed 5 consists primarily of methane, but with small proportions of other gaseous hydrocarbons, hydrocarbon vapours, and water vapour. The gas feed 5 may for example be at a pressure of 4.0 MPa (40 atmospheres) and 35 °C, following sea water cooling from an initial temperature of 90 °C, and may constitute associated gas from a well producing crude oil.

The natural gas feed 5 is supplied to a pretreatment system 25, in which it is subjected to treatment which may comprise one or more of the following: changing its pressure; changing its temperature; and removing impurities such as sulphur. It is then mixed with steam in a mixer 26.

2. Making Synthesis Gas The gas/steam mixture, preferably at a temperature of about 450 °C, is then fed into a catalytic steam/methane reformer 30. The first section of the reformer 30 may be a pre-reformer in which any ethane or higher hydrocarbons are converted to methane. The reformer 30 consists of a compact catalytic reactor formed from a stack of plates defining two sets of channels arranged alternately. One set of channels are for the reforming reaction, and contain a reforming catalyst on removable corrugated metal foil supports, while the other set of channels are for the provision of heat. In a modification the pre-reformer and the reformer are separate reactors.

In this embodiment the heat is provided using a separate burner 32, the exhaust gases from the burner 32 at about 850 °C being passed through the reformer 30 in counter-current to the flow of the steam/methane mixture. The reaction channels of the reformer 30 may contain a nickel catalyst in an initial part of the channel, of length between 100 and 200 mm, for example 150 mm, out of a total reaction channel length of 600 mm. In the first part of the channel, where the nickel catalyst is present, pre- reforming takes place, so any higher hydrocarbons will react with steam to produce methane. The remainder of the length of the reaction channels contains a reformer catalyst, for example a platinum/rhodium catalyst, where the steam and methane react to form carbon monoxide and hydrogen.

The heat for the steam/methane reforming reaction in the reformer 30 is provided by combustion of a fuel gas from a fuel header 34 in a stream of combustion air. In this example the fuel gas is primarily hydrogen. The combustion air is provided by a blower 36 and is preheated in a heat exchanger 38, taking heat from the hot exhaust gases from the combustion after they have passed through the reformer 30. In addition a mixture of steam and alcohol vapour 40 is introduced into the combustion air upstream of the burner 32. After passing through the heat exchanger 38 the exhaust gases may be vented through a stack 39.

A mixture of carbon monoxide and hydrogen at above 800 °C emerges from the reformer 30, and is quenched to below 400 °C by passing it through a steam-raising heat exchanger 42 in the form of a thermosiphon. The heat exchanger 42 is a tube and shell heat exchanger, the hot gases passing through the tubes, and with inlet and outlet ducts communicating with the shell at the top and bottom, and communicating with a steam drum 44. The steam drum 44 is about half full of water, and so water circulates through natural convection between the heat exchanger 42 and the steam drum 44. The resulting steam from the steam drum 44 is supplied to the mixer 26 through a control valve 46.

The gas mixture, which is a form of synthesis gas, may be subjected to further cooling (not shown). It is then subjected to compression using two successive compressors 50, preferably with cooling and liquid-separation stages (not shown) after each compressor 50. The compressors 50 raise the pressure to about 2.6 MPa (26 atm) (absolute).

It will be appreciated from equation (1 ) above that the ratio of hydrogen to CO produced in this way is about 3:1 , whereas the stoichiometric requirement is about 2:1 , as is evident from equation (2). The high-pressure synthesis gas is therefore passed by a hydrogen-permeable membrane 52 to remove excess hydrogen. This hydrogen is supplied to the fuel header 34, and is the principal fuel gas.

3. Fischer-Tropsch Synthesis

The stream of high pressure carbon monoxide and hydrogen is then heated to about 200 °C in a heat exchanger 54, and then fed to a catalytic Fischer-Tropsch reactor 55, this again being a compact catalytic reactor formed from a stack of plates as described above. Referring to figure 2 there is shown a sectional view of part of the Fischer-Tropsch reactor 55. The reactor 55 consists of a stack of flat plates 12 of thickness 1 mm spaced apart so as to define channels 15 for a coolant fluid alternating with channels 17 for the Fischer-Tropsch synthesis. The coolant channels 15 are defined by sheets 14 of thickness 0.75 mm shaped into flat-topped sawtooth corrugations. The height of the corrugations (typically in the range 1 to 4 mm) is 2 mm in this example, and correspondingly thick solid edge strips 16 are provided along the sides, and the wavelength of the corrugations is 12 mm. The channels 17 for the Fischer-Tropsch synthesis are of height 5 mm (typically within a range of 2 mm to 10 mm), being defined by bars 18 of square or rectangular cross-section, 5 mm high, spaced apart by 80 mm (the spacing typically being in a range of 20 - 100 mm) and so defining straight through channels. Within each of the channels 17 for Fischer-Tropsch synthesis is a catalytic insert 20 consisting of a corrugated 50 μηι thick foil (typically of thickness in the range from 20-200 μηι) with a ceramic coating acting as a support for the catalytic material (only two such inserts 20 are shown); instead of a single foil, the insert 20 may consist of a stack of shaped foils. The reactor 55 may be made by stacking the components that define the channels 15 and 17, and then bonding them together for example by brazing or by diffusion bonding. The bonded stack is then turned through 90° so that the channels 15 and 17 are upright, and the catalytic inserts 20 are inserted into the channels 17. Referring again to Figure 1 , the reactant mixture flows through the channels 17, while a coolant flows through the other channels 15. The coolant is circulated by a pump 56 and through a heat exchanger 58. The Fischer-Tropsch reaction takes place at about 210°C, and the coolant is circulated at such a rate that the temperature varies by less than 10 K on passage through the reactor 55.

The reaction products from the Fischer-Tropsch synthesis, predominantly water and hydrocarbons such as paraffins, are cooled to about 70 °C to condense the liquids by passage through a heat exchanger 60 and fed to a separating chamber 62 in which the three phases water, hydrocarbons and tail gases separate. The aqueous phase contains water with about 1 -2% oxygenates such as ethanol and methanol which are formed by the Fischer-Tropsch synthesis. Some of the aqueous phase from the separating chamber 62 is treated by steam stripping 63 to separate the oxygenates (marked "alcohol") to leave clean water that may be discharged to waste. The separated oxygenates, which are at an oxygenate concentration of about 80%, may be stored for subsequent use, as described below. The remainder of the aqueous phase is fed as process water through the heat exchanger 58, and hence through a pressure- drop valve 64 into a stripper tank 66. In the stripper tank 66 the aqueous phase boils, typically at a pressure of about 1 .0 MPa (10 atm), the liquid phase being fed from the bottom of the stripper tank 66 into the steam drum 44, while the vapour phase, which contains steam and the bulk of the oxygenates, provides the stream 40 that is introduced into the combustion air through a control valve 68.

The hydrocarbon phase from the separating chamber 62 is the longer-chain hydrocarbon product. The vapour and gas phase from the separating chamber 62 is fed through two successive cooling heat exchangers 70, the second of which cools the vapours to ambient temperature. Any liquids that condense on passage through the first heat exchanger 70 are fed back into the separating chamber 62. The output from the second heat exchanger 70 is fed into a phase separating chamber 72, where the water and light hydrocarbon product liquid separate.

The remaining vapour phase, which is at the same pressure as the Fischer- Tropsch reactor 55, is then passed through a heat exchanger 74 to a throttle valve 76 followed by a phase separating vessel 78. As the gas passes through the throttle valve 76 it expands into a lower pressure region adiabatically, with no significant heat input from the surroundings. Consequently, in accordance with the Joule Thomson effect, the gas is cooled considerably. The liquids that emerge from the phase separating pressure 78 contain water and light hydrocarbon product. The gases that emerge from the phase separating vessel 78, which are the tail gases from the Fischer-Tropsch process, are passed back through the heat exchanger 74 to cool the in-flowing gases and, optionally, through a hydrogen permeable membrane (not shown). Part of the tail gas may be fed back into the synthesis gas stream upstream of the first compressor 50. At least part of the tail gas is fed into the fuel header 34, to ensure that there is no excessive build-up of methane in the Fischer-Tropsch reactor 55.

The fuel header 34 not only provides the fuel for the burner 32, but also supplies fuel via a fuel compressor 80 to a gas turbine 82. Indeed compressed fuel gas may also be supplied to other equipment (not shown) that does not form part of the plant 10. The gas turbine 82 may be arranged to provide electrical power for operating the plant 10. As indicated by a broken line in the figure, in this example the electrical power generated by the gas turbine 82 is used to power the compressors 50. Alternatively the gas turbine 82 may be coupled directly to drive the compressors 50.

4. Regeneration

The above description is of the normal operation of the plant 10. A commercial plant may include several steam/methane reformers 30 operating in parallel, and may also include several Fischer-Tropsch reactors 55 operating in parallel.

During operation of the plant 10 the catalysts on the catalyst inserts 20 in the Fischer-Tropsch reactor 55 may become deactivated. This may for example be due to poisoning by ammonia, but deactivation can also occur in other ways, for example merely due to the length of time that the catalyst has been in use; or because the catalyst has been exposed to significant levels of water vapour, which accelerates the oxidation of active cobalt; or as a consequence of a rapid increase in CO conversion, with a consequential decrease in space velocity.

The catalysts in the Fischer-Tropsch reactor 55 or other similar reactor can be regenerated in situ, without removing the catalyst inserts 20 from the reactor 55, and without disconnecting the reactor 55 from the plant 10. A suitable regeneration process is as follows:

Step (1 ): The pressure within the reactor 55 is gradually decreased, in a slow and controlled manner, from the normal operating pressure (of about 2.5 MPa (gauge)) down to 0.5 MPa (gauge). This is achieved by gradually decreasing the degree of compression provided by the compressors 50. Preferably the pressure is decreased stepwise, by between 0.05 and 0.2 MPa, for example 0.1 MPa, on each step, and then held for at least 10 minutes before the next step.

Throughout this process the syngas flow and the temperature can remain at their normal operating values. The temperature of the reactor 55 is continuously monitored, and if any rapid changes are observed, then the pressure is held steady (or gradually increased) until stable operation is again observed.

Step (2): The syngas H ₂ /CO ratio is then gradually increased, either continuously or stepwise, to achieve a ratio of at least 2.0, more preferably above 3.0, for example about 3.5 or even higher. The reactor 55 typically operates with a H ₂ /CO ratio between 1 .9 and 2.0 and therefore the ratio must be increased by at least 0.1 , although the regeneration will be faster and may also be more effective if the increase is by 1 .0 or more. In the case of the reactor 55 this would arise from increasing the ratio to between 2.9 and 3.0. This can for example be achieved by gradually bypassing the membrane separator 52, and, if necessary, increasing the proportion of steam and changing the reaction conditions in the reformer 30 so as to achieve a higher ratio.

Step (3): The normal operating temperature of the reactor 55 is about 225 °C, and the temperature is controlled by the fluid flowing through the channels 15. Under normal operating conditions this requires a coolant; but a source of a hot fluid (not shown) is also available (for example using heat from the steam drum 44 or from an electrical heater), so the reactor temperature can instead be raised, for example during start-up. By providing a combination of the hot fluid and the coolant, a desired reactor temperature can therefore be obtained.

So, as the next stage, the temperature is gradually raised to 230 °C, for example being ramped continuously up at 0.2 °C/hr. If the reactor 55 had been operating at 230 °C initially, and this temperature had been maintained throughout the preceding regeneration steps, then there is no requirement to raise the temperature.

In a modification to the regeneration process, the temperature at this step is gradually raised to 225 ^q C. Again it is clear that if the reactor 55 had been operating at 225 °C initially, and this temperature had been maintained throughout the preceding regeneration steps, then no temperature rise is required. Step (4): The reactor 55 is then held at the resulting low-pressure, high syngas ratio and high temperature state, that is to say 0.5 MPa, ratio of 3.5, and temperature 230 ° or 225 °C, for a prolonged period. Preferably the reactor 55 is held in this state for between 60 and 72 hours, but a longer period of say 80 or 90 hours would be equally acceptable. This step regenerates and reactivates the catalyst in the inserts 20.

Step (5): The temperature of the reactor 55 is then gradually lowered to 210 °C at a rate of 0.5°C/min. Step (6): The syngas H ₂ /CO ratio is then gradually decreased, either continuously or stepwise, to achieve an operating ratio of, for example, between 1 .9 and 2.0. This can for example be achieved by gradually passing more of the synthesis gas through the membrane separator 52. Step (7): The pressure in the reactor 55 is then gradually increased back to its normal operating value, by gradually bringing the compressors 50 back into operation. For example the pressure may be increased either continuously, or stepwise. In one example the pressure is increased in steps of 0.1 MPa, and the reactor 55 is held at each successive pressure for at least 10 minutes before the next pressure increase. The temperature is monitored, and if any rapid change occurs, then the pressure increase is no longer increased, or may even be decreased, or the temperature decreased, until a substantially steady condition is again attained.

Step (8): The final step is to return the temperature gradually back up to the normal operating temperature, which as described above may be 225 °C. By way of example the temperature might be ramped at 0.2°C/min from 210 ^< Ό up to 220 °C; and then ramped more slowly, at 0.1 Ό/min from 220 °C up to 225 ^< €.

The reactor 55 is then back in its normal operating conditions, but with the catalyst reactivated and regenerated.

It will be appreciated that this regeneration process can be carried out in situ, and without modifying the associated plant. It has been found that this process is particularly suitable for regenerating catalysts that have been poisoned by ammonia, but it may also be satisfactory for regenerating catalysts that have been deactivated in other ways, for example catalysts that have been deactivated over time under normal operation conditions, catalysts that have been deactivated by water, which accelerates the oxidation of active cobalt, and catalysts deactivated as a consequence of a rapid increase in CO conversion.

As regards catalysts that are poisoned by ammonia, it is understood that the deactivation occurs as a consequence of competitive adsorption between ammonia and the synthesis reactants. There are believed to be two types of active sites on the catalyst: weakly acidic sites and strongly acidic sites. At weakly acidic sites, ammonia adsorbs only weakly and therefore can be removed under normal Fischer-Tropsch operating conditions merely by ensuring that ammonia is completely absent from the gas feed. This would lead to partial catalyst recovery. At strongly acidic sites, however, ammonia adsorbs strongly and is not removed under normal Fischer-Tropsch operating conditions. Removal of ammonia from these sites is enhanced by use of higher temperatures, and by increasing the hydrogen concentration, which can provide almost complete catalyst regeneration.

As mentioned above, a suitable catalyst for the Fischer-Tropsch synthesis comprises metallic cobalt promoted with for example ruthenium, on a support such as alumina. The cobalt is in the form of discrete crystals on the surface of the alumina. The acidity decreases with increasing loading of ruthenium. The most highly acidic sites (Bronsted acid sites) are believed to be located at the interface between the alumina and the cobalt. The hydrogen-rich atmosphere is expected to enhance ammonia desorption because reduction of any oxidised cobalt would reduce the acidity of the catalyst. Since acidity of the catalyst is critical for its deactivation by ammonia, acidity can be carefully tuned during catalyst preparation to achieve catalysts that show

considerable stability against NH ₃ -poisoning. This could be achieved for instance in one or a combination of the following ways:

Use precipitation of cobalt nitrate with ammonia as preparation route - Pre-treating the support (e.g. alumina) with dilute NH ₃ or alkaline solution prior to Co-Ru addition

Use Co/Ru precursors with stronger basic character (e.g. NH ₃ -containing precursor)

Addition of metal promoters with stronger basic character.