The present invention relates to a process for performing Fischer-Tropsch synthesis, and in particular for ensuring satisfactory operation in the presence of a contaminant. The invention may be applicable 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 preexisting 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 gas stream may contain contaminants which have a detrimental effect on the Fischer-Tropsch catalyst. The Fischer-Tropsch catalyst may therefore, at some stage, require regeneration. 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 to avoid the need for regeneration of the catalyst would provide clear advantages.

This issue arises in particular with ammonia, which may be such a

contaminant. Even low levels of ammonia, for example less than 50 parts per thousand million (50 ppb) in synthesis gas can have a significant effect on operation of a Fischer-Tropsch reactor.

According to the present invention there is provided a process for operation of a Fischer-Tropsch synthesis reactor for generating hydrocarbons from a synthesis gas containing hydrogen and carbon monoxide, the reactor containing a Fischer- Tropsch catalyst, wherein the synthesis gas contains a contaminant, and the process comprises the following steps:

(a) continuing to pass the synthesis gas through the reactor, while increasing the hydrogen:CO ratio of the synthesis gas;

(b) ensuring the reactor does not experience a thermal runaway; and

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

The process of the invention may therefore be used if the synthesis gas stream is found to contain a contaminant, for example ammonia, so as to maintain the productivity of the reactor. Hence step (c) may be performed for a prolonged period, until the contaminant in the synthesis gas stream is no longer present, or is present at negligible quantities.

It will be appreciated that in step (a) the hydrogen:CO ratio is raised above the ratio that would give optimal productivity in the absence of the contaminant.

So, for example a process for operation of a Fischer-Tropsch synthesis reactor for generating hydrocarbons from a synthesis gas containing hydrogen and carbon monoxide, the reactor containing a Fischer-Tropsch catalyst and being operable at a temperature of 210 °C with synthesis gas having a hydrogen:CO ratio less than 2.0, wherein the synthesis gas contains a contaminant, may comprise the following steps:

(a) continuing to pass the synthesis gas through the reactor, while increasing the hydrogen:CO ratio of the synthesis gas to at least 2.0;

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

(c) holding the reactor at these operating conditions for at least one day. That the reactor is operable at a temperature of 21 CO with synthesis gas having a hydrogen:CO ratio less than 2.0 means that the reactor could operate under these conditions, although operation under those conditions may not provide optimum performance. In practice the normal operating conditions of the reactor may differ from these values, for example to improve performance, and the normal operating conditions may also vary with the age of the catalyst.

It will be appreciated that this process can be carried out without modifying the associated plant. It has been found that this process is particularly suitable for preventing poisoning of catalysts by ammonia, but it may also be effective for preventing deactivation of catalysts 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 continues under the modified conditions. The productivity of C5+ during this mode of operation may be as high as the productivity that had been obtained previously.

The process preferably includes the additional steps of:

(b1 ) adjusting the hydrogen:CO ratio in the synthesis gas stream to a range of values between 1 .8 and 2.5, while ensuring that the temperature within the reactor does not exceed 250 °C;

(b2) monitoring the reactor performance at each different hydrogen:CO ratio, and deducing the hydrogen:CO ratio that provides optimum reactor performance; and (b3) adjusting the hydrogen :CO ratio to the hydrogen:CO ratio that provides optimum reactor performance;

before performing step (c).

Preferably the process involves ensuring that the temperature within the reactor does not exceed 240 °C, in every step. This is to ensure that overheating of the reactor does not occur, and in particular that a thermal runaway is prevented.

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 to catalyse the reaction. The catalyst may be a catalyst structure, which 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.

Step (a) of the process requires an increase in the syngas H ₂ /CO ratio. A typical ratio during normal operation would be between 1 .8 and 2.0. This change should 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) of the process may also require a change of temperature. In step (b) the temperature may be 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 °C/hr.

In each case any changes may be performed continuously, or stepwise. For example when increasing a parameter, the parameter may be increased in steps, and held for at least 10 minutes at each value. When changing the hydrogen:CO ratio, the ratio may be changed continuously, or may be increased stepwise, for example increasing the ratio by 0.1 or by 0.05 at each step, and is held for at least 30 min at each value to ensure stable operation.

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 3.0, 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 190°C and 280 °C, for example 210Ό or 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 catalysts promoted with precious metals such as Pd, Pt, Ru or Re doped, for example, to 1 wt% (as a percentage of the weight of the cobalt metal) 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 that produces 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, which may be 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 reactor 55 may also be 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 catalyst, which in this example 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, being circulated at such a rate that the temperature varies by less than 10 K on passage through the reactor 55. Initially, when the catalyst is new, the Fischer-Tropsch reaction takes place at about 210°C.

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, that is to say hydrocarbons that are liquids at about 70 °C, having a chain length C8 and above. 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, that is to say hydrocarbons that are liquids at ambient conditions, primarily those of chain lengths C5 to C8.

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. Deactivation

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. It may therefore subsequently be necessary to replace the catalysts, or to subject them to a regeneration process. However, particularly if ammonia is present in the synthesis gas, it may be appropriate to modify the operation of the plant 10.

5. Modified Operation

If an operator recognises that a contaminant such as ammonia is present in the synthesis gas, he may be able to modify the operation of the Fischer-Tropsch reactor 55 to compensate for the effect of the contaminant. It has been found that even a low level of ammonia in the synthesis gas stream can have a significant detrimental effect. For example the presence of as little as 30 ppb of ammonia in the synthesis gas has been observed to decrease the CO conversion from 25.0% to 19.7%, and decrease the C5+ productivity (i.e. the productivity of hydrocarbons with at least five carbon atoms) from 0.66 to 0.5 g/g _cat .hr. For example the present invention may be applied in a Fischer-Tropsch reactor 55 for which the normal operating condition is at 210 °C with a hydrogen:CO ratio of 1 .86, and a gas pressure of 2.6 MPa (absolute). In the absence of contaminants it is known that an increase in the syngas ratio (i.e. the hydrogen:CO ratio) will increase conversion, but will decrease the C5+ selectivity, and may lead to a thermal runaway. The process of the invention involves continuing to pass the synthesis gas through the reactor 55, while gradually increasing the hydrogen:CO ratio of the synthesis gas. This adjustment of the syngas ratio has been found to suppress the detrimental effect of ammonia on Fischer-Tropsch performance at a constant operating temperature. Surprisingly, this adjustment can be carried out in such a way as to avoid any negative impact on C5+ selectivity, so that the C5+ selectivity does not significantly decrease. Any effect due to change of operating temperature is less significant.

The increase in the hydrogen :CO ratio may be carried out stepwise, for example introducing additional hydrogen into the synthesis gas so as to increase the syngas ratio to 2.0, and then to 2.1 , and then to 2.2, and then to 2.3, and then to 2.4. Each time the syngas ratio is changed, the conditions are monitored to ensure that the maximum temperature in the reactor 55 does not exceed 240 °C, and if the temperature approaches that value the syngas ratio is slightly decreased so the temperature drops, before the syngas ratio is gradually increased again. The reactor 55 is operated at each different syngas ratio for at least 1 hour, ensuring stable operation conditions, and the performance of the reactor 55 is monitored for example by analysing the composition of the outlet gas. This analysis may for example utilise a gas chromatograph.

The maximum temperature that is set depends on the age of the catalyst and the normal operating temperature of the reactor 55. In this example the reactor 55 is normally operating at 210°C, which is appropriate when the catalyst is new. In this case the maximum temperature, as the syngas ratio is varied, is set at 240 °C. Aged catalyst is less prone to thermal runaway than new catalyst. So, later in the life of the catalyst, the normal operating temperature of the reactor 55 may be higher, for example 240 °C, and the maximum temperature during the changes in syngas ratio may be set at a higher value, such as 250 °C. The occurrence of a thermal runaway may indicate that the reactor 55 is producing a significantly increased proportion of methane, and so is achieving a lower C5+ selectivity.

From the performance of the reactor 55 with different syngas ratios, the variation of C5+ selectivity with the syngas ratio can be observed, and the optimum syngas ratio deduced. The proportion of hydrogen would then be gradually changed, in a similar stepwise fashion or continuously, until that optimum syngas ratio is reached. This syngas ratio would then be held for as long as the contaminant is present in the synthesis gas.

There may be no change in operating temperature at steady-state.

Alternatively, the operating temperature may be slightly increased, for example from 210 ° to 220 °C, as such a temperature increase tends to compensate for deactivation of the catalyst.

For example, in an experimental test, the presence of 30 ppb (parts per billion, i.e. parts per 10 ⁹ ) of ammonia in the synthesis gas has been observed to decrease the CO conversion from 25.02% to 19.67%, and decrease the C5+ productivity from 0.66 to 0.50 g/g _cat .hr. This was with a reactor for which the optimum operating conditions without the presence of ammonia required a syngas ratio of 1 .86. Increasing the ratio to 2.0 led to an improved performance, with the CO conversion increasing from 19.67 to 22.82%, and the productivity increasing from 0.50 to 0.56 g/g _ca t-hr, and with the selectivity being unaffected. However increasing the syngas ratio to above 2.1 was found to lead to the expected decrease in selectivity, and increase in conversion, so the overall effect was detrimental. So, in this case, in the presence of this level of ammonia contamination, the performance of the reactor 55 can be optimised by operating at a syngas ratio of about 2.0.

It may therefore be appropriate to perform further tests, for example to monitor performance of the reactor at syngas ratios between 1 .9 and 2.1 , in order to find a more accurate indication of the optimum syngas ratio for this level of contamination.

The mechanism by which a Fischer-Tropsch catalyst is deactivated by ammonia is believed to be competitive adsorption. Increasing the syngas ratio increases the proportion of hydrogen present in the gas phase, and this additional hydrogen shifts the equilibrium from adsorption of ammonia to desorption. Hence less ammonia is adsorbed.