A method for producing syngas using catalytic reverse water gas shift

Abstract

The present invention relates to a method for producing syngas using a catalytic reverse water gas shift (RWGS) reaction, the method at least comprising the steps of: a) providing a feed stream (10) comprising at least hydrogen (H2) and carbon dioxide (CO2); b) adding a sulphur-containing compound (15) to the feed stream provided in step a), thereby obtaining a sulphur-enriched feed stream (20); c) introducing the sulphur-enriched feed stream (20) obtained in step (b) into a RWGS reactor (2) and subjecting it to a catalytic RWGS reaction in the presence of a catalyst, thereby obtaining a syngas containing stream (30).

Description

The present invention relates to a method for producing syngas using a catalytic reverse water gas shift (RWGS) reaction. Methods for producing syngas using RWGS are known. RWGS reactions convert carbon dioxide (CO2) and hydrogen (H2) into 'syngas' , which contains at least carbon monoxide (CO) and hydrogen (H2) , and typically also water (H2O) and unconverted carbon dioxide (CO2) . The RWGS reaction is endothermic in nature; hence, it is necessary to supply sufficient thermal energy to the reactants (i.e. carbon dioxide and hydrogen) to facilitate the endothermic RWGS reaction. The RWGS reaction is in fact the backward reaction of the equilibrium of the 'water gas shift' (WGS) reaction, which is a well-known reaction to convert carbon monoxide and water to carbon dioxide and hydrogen. The RWGS reaction can proceed without the use of a catalyst, but this requires very high temperatures (e.g. 1000°C or even much higher) favoring both the kinetics and maximum achievable equilibrium conversions. If a catalyst for the RWGS reaction is used, much lower temperatures may be required for the reaction to proceed and the reaction conditions and catalyst used are to be selected such that the catalyzation of the very exothermic methanation reaction (CO2 + 4H2 -> CH4 + 2H2O) is avoided or at least minimized. The thermodynamics may drive the reaction towards methanation (rather than towards RWGS) and too low temperatures may severely lower the equilibrium conversion of RWGS itself, so finding a - 2 -process scheme, reaction conditions and a catalyst resulting in acceptable conversion of CO2 to syngas with non-methanation or very low methanation is a key challenge . Although the equilibrium conversion levels of RWGS are favourably higher at high temperatures, process integration considerations including heat recovery, recycling options, materials and corrosion limitations and technical readiness level of reactor concepts are drivers to preferably execute RWGS at a medium temperature level between 400-700°C, and preferably between 450 and 600°C. It is especially in this temperature window combined with elevated pressure levels where a-selective formation of methane and / or carbon is thermodynamically favoured. Currently, the status of developments regarding the RWGS reaction have been mostly on lab-scale or precommercial pilot or demonstration scale. There is still a lot to explore until large-scale RWGS will be a commercially attractive option. For large-scale conversion of carbon dioxide there is a need to be able to more efficiently and economically carry out the RWGS reaction. In achieving high conversion of carbon dioxide selectively to carbon monoxide, byproducts like methane and carbon formation are to be avoided. Also, the amount of energy input required for performing the endothermic RWGS reaction requires attention . As a mere example of a recently published RWGS method, WO 2022 / 263384 Al and WO 2022 / 129338 Al disclose methods for producing syngas using a catalytic RWGS reaction in the presence of a 'non-methanation promoting' - 3 -catalyst (such as cerium oxide, zirconium oxide or a combination thereof). The present inventors have found that even if a non-methanation promoting catalyst is used in a catalytic RWGS reaction, methanation may still occur, for example at the reactor walls in case nickel (Ni) as present in the material of the reactor wall is not sufficiently protected by an oxide layer. Exothermic methanation at the reactor walls may then accelerate degradation of the reactor walls through metal dusting. Furthermore, even though the term non-methanation promoting catalyst is used in WO 2022 / 129338 Al, some minor residual selectivity to methane may be present on the materials and / or catalysts mentioned, either intrinsically caused by the materials or through the presence of metal contaminants unwantedly added to the catalyst during production of the catalysts or deposited during operation. As an example, Example 3 of WO 2022 / 129338 Al reports a residual selectivity to methane of 0.1-0.3% using a mixed CeO2 / ZrO2 in an Alloy 800 reactor tube. Furthermore, the tendency of methanation and metal dusting is much higher at desired pressure levels of industrial operation. Many lab scale studies present low methanation selectivities at combined temperature and pressure levels at which methane formation from a thermodynamic point of view is limited. It is an object of the present invention to minimize one or more of the above problems. It is a further object of the present invention to provide an alternative method for producing syngas using a catalytic RWGS reaction that can be performed at lower temperatures, preferably lower than 700°C, whilst still achieving a high (> 90%) overall CO2 conversion. One or more of the above or other objects can be achieved by providing a method for producing syngas using a catalytic reverse water gas shift (RWGS) reaction, the method at least comprising the steps of: a) providing a feed stream comprising at least hydrogen (H2) and carbon dioxide (CO2) ; b) adding a sulphur-containing compound to the feed stream provided in step a), thereby obtaining a sulphur-enriched feed stream; c) introducing the sulphur-enriched feed stream obtained in step (b) into a RWGS reactor and subjecting it to a catalytic RWGS reaction in the presence of a catalyst, thereby obtaining a syngas containing stream. It has surprisingly been found according to the present invention that by on purpose actively adding a sulphur-containing compound to the feed stream, the occurrence of methanation (methane formation) at e.g. the reactor walls can be prevented or even avoided. A further advantage according to the present invention is that even though the RWGS reaction is performed at relatively low temperatures (such as below 700°C), a desirable overall conversion of CO2 may be achieved, whilst minimizing carbon formation and / or methanation (methane formation) even though the latter two are thermodynamically favoured. An important advantage of the present invention is that less expensive materials need to be used for e.g. the reactors, heaters and heat exchangers in view of the lower temperatures being used which alleviates materials problems related to the nature of the gas stream (e.g. metal dusting, methanation, etc.). In step a) of the method according to the present invention a feed stream is provided comprising at least hydrogen (H2) and carbon dioxide (CO2) . The person skilled in the art will readily understand that the feed stream is not particularly limited and may come from various sources. Typically, the feed stream comprises 30-80 vol.% H2, preferably 40-70 vol.% H2, and typically 20-70 vol.% CO2, preferably 30-60 vol.% CO2. Other components such as CH4, CO, H2O, C2+, 0=2+, N2, Ar, O2, and nitrogen compounds (such as NOX, NH3) , which are sometimes intrinsically present in the feed stream may be present. Also, the feed stream may contain small amounts of sorbent (such as amines, KOH, MeOH, glycols, etc.), e.g. as used in DAC [Direct Air Capture]) or other CO2 removal units . Generally, the feed stream has a temperature of 5-700°C and, preferably above 20°C. If needed, the feed stream provided in step a) may be heated (e.g. by indirect heat exchange) to a temperature of 200-700°C, preferably 450-600°C, before it enters the RWGS reactor in step c). The feed stream typically has a pressure in the range of from 1 to 200 bara. Preferably, the pressure is from 5 to 70 bara. Preferably, the feed stream has a hydrogen to carbon dioxide (H2 / CO2) volume ratio of below 3.0, preferably below 2.0, more preferably below 1.75. The H2 / CO2 volume ratio of hydrogen to carbon dioxide is typically adjusted to the process line-up including the number of reactors, optional intermediate product removal and optional CO2 recycling ratio such that the required hydrogen to carbon monoxide ratio in the eventual product stream is obtained, e.g. matching the desired ratio of the next syngas conversion step. In step b) of the method according to the present invention, a sulphur-containing compound is added to the feed stream provided in step a), thereby obtaining a sulphur-enriched feed stream. As mentioned above, the sulphur-containing compound is on purpose actively added to the feed stream, in addition to any sulphur-containing component already intrinsically being present in the feed stream. The sulphur-containing compound to be added in step b) is not particularly limited. However, preferably, the sulphur-containing compound added in step b) is H2S. Further it is preferred that the sulphur-containing compound is added in step b) in an amount of from 50 to 2000 ppbv, preferably between 100 and 500 ppbv, based on the feed stream. In step c) of the method according to the present invention, the sulphur-enriched feed stream obtained in step (b) is introduced into a RWGS reactor and subjected to a catalytic RWGS reaction in the presence of a catalyst, thereby obtaining a syngas containing stream. As the person skilled in the art is familiar with RWGS reactors and conditions of catalytic RWGS reactions, this is not discussed here in detail. Typical temperatures of the catalytic RWGS reaction in the RWGS reactor are from 400 to 700°C, preferably from 450 to 600°C, more preferably from 475 to 575°C. The person skilled in the art will understand that the temperature may vary over the reactor (e.g. higher at the inlet than at the outlet, in particular for an adiabatic process). Preferably, the temperature of the catalytic RWGS reaction in step c) is kept below 700°C, preferably below 600°C. As, the RWGS reaction is endothermic, heating needs to be provided to the reactor. This heating may come from any source, e.g. indirectly via heating by molten salt circulating around the individual tubes of a multitubular reactor wherein the circulating molten salt itself is heated by electrical heating, preferably in counter-current mode, or directly via the feed stream in the case of an adiabatic process. Typical pressures as used in the RWGS reactor are 1200 bara, preferably above 20 bara and preferably below 70 bara. Further, typical gas hourly space velocities (GHSV) are 500-100,000 h-1, preferably above 3,000 h-1 and preferably below 10,000 h-1. In the RWGS reactor a catalytic RWGS reaction takes place and this requires the presence of a catalyst. Typically, the RWGS reactor contains a catalyst bed. As the person skilled in the art is familiar with suitable RWGS beds and catalysts, this is not discussed here in detail. Preferably, the catalyst bed comprises a catalyst that is suitable for performing a RWGS reaction below 700°C. Further it is preferred that the catalyst does not promote methanation under the used conditions. Preferred examples of suitable 'non-methanation promoting' catalysts comprise at least cerium oxide, zirconium oxide, or a combination thereof. Hence, according to a preferred embodiment of the present invention, the catalyst as used in step c) comprises cerium oxide, zirconium oxide or a combination thereof. The catalyst may contain further components in addition to the cerium oxide and / or zirconium oxide. According to an especially preferred embodiment of the present invention, the catalyst as used in step c) comprises monoclinic zirconia (m-ZrO2) . It has been surprisingly found according to the present invention that the catalytic activity of monoclinic zirconia is insensitive against the presence of sulphur-containing compounds such as H2S (i.e. a low S uptake), thus maintaining equally good catalytic RWGS activity whereas also the extremely low selectivity to methanation is preserved or even further suppressed. Further it is preferred that the catalyst comprises at least 60 wt. % m-ZrO2, based on the total weight of the zirconia (in the catalyst), preferably at least 70 wt.%, more preferably at least 80 wt.%, even more preferably at least 90 wt.%, yet even more preferably at least 95 wt.%. The catalyst may even consist of (i.e. comprise 100 wt.% of) m-ZrO2, based on the total weight of the zirconia (in the catalyst). The amount of m-ZrO2 can be determined using X-ray diffraction (XRD), more specifically powder X-ray diffraction in reflection mode with Rietveld refinement analysis to quantify the fractions of crystalline zirconia phases. Also, it is preferred that the catalyst comprises at least 60 wt.% m-ZrO2, based on the total weight of the catalyst, preferably at least 70 wt.%, more preferably at least 80 wt.%, even more preferably at least 90 wt.%, yet even more preferably at least 95 wt.%. The catalyst may even consist of (i.e. comprise 100 wt.% of) m-ZrO2, based on the total weight of the catalyst. The shape of the RWGS catalyst particles is not particularly limited and can vary widely. For example, the catalyst particles can have the shape of a ring, sphere, tablet, pellet, etc. Preferably, the catalyst particles have a ring-shape. The size and shape of the catalyst particles can be chosen such as to achieve a - 9 -compromise between parameters such as volumetric activity, pressure drop and heat transfer. Although several types of reactors may be used, the RWGS reactor preferably comprises a multi-tubular reactor heated by molten salt circulating around the tubes of the multi-tubular reactor. In this embodiment, the molten salt provides for the heat required for the endothermic reaction as taking place in the multi-tubular reactor. Preferably, the molten salt is circulating in countercurrent mode around the tubes of the multi-tubular reactor (when compared to the fluid flow in the tubes of the reactor). The circulating molten salt is preferably heated from outside the reactor. Preferably, each of the tubes of the multi-tubular reactor comprises a catalyst. As a result of the RWGS reaction in step c), a syngas containing stream is obtained, at least comprising hydrogen (H2) and carbon monoxide (CO). Typically, the syngas containing stream also contains water (H2O) and unconverted carbon dioxide (CO2) . Typically, the amounts of components in the syngas containing stream as obtained in step c) are around thermodynamic equilibrium concentrations of the RWGS reaction. Generally, the syngas containing stream has a hydrogen to carbon monoxide (H2 / CO) volume ratio in the range of 1.5 to 10, preferably below 5.0. One of the advantages of the present invention is that the used RWGS reaction results in low methanation (methane formation). Preferably, the syngas containing stream comprises at most 1.0 vol. % methane (CH4) , preferably at most 0.1 vol.% methane, most preferably at most 0.01 vol.% methane. Preferably, the temperature of the syngas containing stream obtained in step c) (at the outlet of the RWGS reactor) is kept below 700°C, preferably below 650°C, more preferably below 600°C and typically above 450°C. The syngas containing stream obtained in step c) may be used as such as a product stream or further processed (e.g. subjected to CO2 removal and / or sulphur removal). If desired, multiple RGWS reactors may be used in series, if desired with intermediate heating and / or product separation steps. According to an especially preferred embodiment of the present invention, the method further comprises the steps of : d) cooling the syngas containing stream obtained in step c) thereby obtaining a cooled syngas stream; e) separating the cooled syngas stream obtained in step d) in a gas / liquid separator thereby obtaining a water-enriched stream and a water-depleted syngas stream; f) adding a sulphur-containing compound to the water-depleted syngas stream obtained in step e), thereby obtaining a sulphur-enriched syngas stream; g) introducing the sulphur-enriched syngas stream obtained in step f) into a second RWGS reactor and subjecting it to a second catalytic RWGS reaction, thereby obtaining a second syngas containing stream. In step d) of the method according to the present invention, the syngas containing stream obtained in step c) is cooled, thereby obtaining a cooled syngas stream. This cooling may be performed in various ways and in multiple steps. Preferably the cooling of the syngas containing stream obtained in step c) comprises indirect cooling in a heat exchanger against the feed stream provided in step a). Typically, the cooled syngas stream has a temperature of 20-80°C and, preferably below 60°C. in step e) of the method according to the present invention, the cooled syngas stream obtained in step d) is separated in a gas / liquid separator thereby obtaining a water-enriched stream and a water-depleted syngas stream. Typically, the amounts of components in the water-depleted syngas stream are around thermodynamic equilibrium concentrations. Typically, the water-depleted syngas stream comprises at most 5 vol.% H2O, preferably at most 1 vol.% H2O, more preferably at most 0.1 vol.% H2O, even more preferably no H2O at all. If desired, the water-depleted syngas stream as obtained in step e) may be heated before the adding of a sulphur-containing compound in step f). In step f) of the method according to the present invention, a sulphur-containing compound is added to the water-depleted syngas stream obtained in step e), thereby obtaining a sulphur-enriched syngas stream with a sulphur-content optimized for step (g). As mentioned above in relation to adding a sulphur-containing compound to the feed stream, the sulphur-containing compound is on purpose actively added to the water-depleted syngas stream obtained in step e), in addition to any sulphur-containing component already intrinsically being present. Again, the sulphur-containing compound to be added in step f) is not particularly limited. However, preferably, the sulphur-containing compound added in step f) is H2S. Further it is preferred that the sulphur-containing compound is added in step f) in an amount of from 50 to 2000 ppbv, preferably from 100 to 500 ppbv. In step g) of the method according to the present invention, the sulphur-enriched syngas stream obtained in step f) is introduced into a second RWGS reactor and subjected to a second catalytic RWGS reaction, thereby obtaining a second syngas containing stream. The second syngas containing stream may be used as such as a product stream or further processed. The person skilled in the art will understand that further catalytic RGWS reactors may be present (and that a sulphur-containing compound may be added to obtain a sulphur-enriched syngas stream for introducing in said further RWGS reactors). According to a preferred embodiment of the present invention, the syngas containing stream obtained in step (c) and / or the second syngas containing stream obtained in step (g) is subjected to a sulphur-removal step, before being subjected to a subsequent syngas conversion reaction such as methanol synthesis, Fischer-Tropsch synthesis or the like. The sulphur-removal step can for example be done using absorption. In a further aspect, the present invention provides an RWGS catalyst, at least comprising: - monoclinic zirconia (m-ZrO2) . Preferably, the catalyst comprises at least 60 wt. % m-ZrO2, based on the total weight of the zirconia (in the catalyst), preferably at least 70 wt. %, more preferably at least 80 wt.%, even more preferably at least 90 wt.%, yet even more preferably at least 95 wt.%. As also mentioned above, the catalyst may even consist of (i.e. comprise 100 wt.% of) m-ZrO2, based on the total weight of the zirconia (in the catalyst). According to a further preferred embodiment, the catalyst comprises at least 60 wt.% m-ZrO2, based on the total weight of the catalyst, preferably at least 70 wt.%, more preferably at least 80 wt.%, even more preferably at least 90 wt.%, yet even more preferably at least 95 wt.%. Again, the catalyst may even consist of (i.e. comprise 100 wt.% of) m-ZrO2, based on the total weight of the catalyst. The catalyst is also substantially free of metals from Groups 8, 9, and 10 of the Periodic Table of Elements. As used herein, the term "substantially free of metals" is intended to denote an amount of metals that is less than 0.01wt.%. Hereinafter the present invention will be further illustrated by the following non-limiting drawings. Herein shows: Fig. 1 schematically a first embodiment of a method for producing syngas using a catalytic RWGS reaction according to the present invention; Fig. 2 schematically examples of different reactor types that can be used for the RWGS reactor as used according to the present invention; Fig. 3 schematically a second embodiment of a method for producing syngas using a catalytic RWGS reaction according to the present invention; Fig. 4 shows the CO2 conversion and CH4 selectivity for monoclinic zirconium oxide as a function of time in the absence of H2S; and Fig. 5 shows the CO2 conversion and CH4 selectivity for monoclinic zirconium oxide as a function of time in the presence of 100 ppbv H2S . For the purpose of this description, same reference numbers refer to same or similar components. The flow scheme of Figure 1, generally referred to with reference number 1, comprises a RWGS reactor 2, heat exchangers 3-5, a first gas / liquid separator 6 (in the form of a H2O knock-out drum) and a CO2 removal unit 8. In the embodiment of Fig. 1, the RWGS reactor 2 comprises a catalyst bed and is provided with external heating 7 (e.g. in the form of electrical heating or molten salt heater). The heat exchangers 3, 4 and 5 may be integrated with the external heating 7. During use of the embodiment of Fig. 1, a feed stream 10 is provided, which comprises at least hydrogen (H2) and carbon dioxide (CO2) . A sulphur-containing compound (i.c. H2S) is added (via stream 15) to the feed stream 10 to obtain a sulphur-enriched feed stream 20. In the embodiment of Fig. 1, the sulphur-enriched feed stream 20 is heated in the first heat exchanger 3 and in a second heat exchanger 4. This second heat exchanger 4 may form part of the (overhead of the) RWGS reactor 2. The person skilled in the art will readily understand that the sulphur-containing compound may also be added after the heating in the heat exchangers 3 and 4 . The heated, sulphur-enriched feed stream 20 is introduced into the RWGS reactor 2 and subjected to a catalytic RWGS reaction, thereby obtaining a syngas containing stream, which is removed as stream 30 from the RWGS reactor 2. Then, the syngas containing stream 30 is cooled in the first heat exchanger 3 by indirect heat exchange against the sulphur-enriched feed stream 20 (or against the feed stream 10), thereby obtaining a first cooled syngas stream 40. The first cooled syngas stream 40 is further cooled in the third heat exchanger 5, thereby obtaining a second cooled syngas stream 50. Subsequently, the second cooled syngas stream 50 is separated in the gas / liquid separator 6 thereby obtaining a water-enriched stream 60 and a water-depleted syngas stream 70. The water-depleted syngas stream 70 is then separated in the CO2 removal unit 8, thereby obtaining a CO2-enriched stream 80 and a CO2-depleted syngas stream 90. Stream 90 can be further processed or used as a product stream. The CO2-enriched stream 80 may be combined with the feed stream 10. Fig. 2 shows schematically non-limiting examples of different reactor types that can be used for the RWGS reactor in the apparatus 1 according to the present invention . The reactor of Fig. 2a) comprises a multi-tubular reactor heated by a molten salt circulating around the tubes of the multi-tubular reactor. Preferably, the molten salt flow inside the shell of the multi-tubular reactor is counter-currently when compared to the flow of the gas inside the tubes. As shown, the molten salt may be heated by separate external heating, preferably an eheater. If molten salt is used for two or more reactors, then there may be a common circuit for the molten salt. The reactor of Fig. 2b) comprises a single catalyst bed, whilst the reactor of Fig. 2c) comprises a single catalyst bed provided with external heating. In Fig. 1 the reactor of the type shown in Fig 2c) is used. Further, the reactor of Fig. 2d comprises 3 catalyst beds with intermediate external heating between the beds. Generally, if any of the reactors of 2b)-d) is used, then preheating (as in heat exchangers 4) is preferred. Fig. 3 schematically shows a second embodiment of a method for producing syngas using a catalytic RWGS reaction according to the present invention. In the embodiment of Fig. 3, the apparatus further comprises a second RGWS reactor 9 (of the multi-tubular type as shown in Fig 2a), a fourth heat exchanger 13, a fifth heat exchanger 14 and external heating 17. As can be seen in Fig. 3, the water-depleted syngas stream 70 is, instead of being sent to the CO2 removal unit 8 as in Fig. 1, after heating in the fourth heat exchanger 13 and after adding a sulphur-containing compound 15, introduced as sulphur-enriched syngas stream 75 into the second RWGS reactor 9 and subjected to a second catalytic RWGS reaction, thereby obtaining a second syngas containing stream 100. The second syngas containing stream 100 is cooled in the fourth heat exchanger 13 against the water-depleted syngas stream 70 to form stream 110. If desired, and similar to the embodiment of Fig. 1, stream 110 may be sent to a CO2 removal unit (not shown). Examples Example 1 This Example 1 provides a comparison of relative activities and selectivities of CeO2, CeZrO2, t-ZrO2, m-ZrO2 and Si-stabilized m-ZrO2 tested on small scale in the absence of feeding H2S. Zirconium oxide (Zr02) comprising samples. - Sample Al: For the preparation of Sample Al a mixed cerium / zirconium oxide powder (Actalys product series) was obtained from Solvay. The powder contained 69 wt. % CeO2 balanced with ZrO2 (and traces of HfO2) and had a BET surface area of 108 m2 / gram (the BET surface was determined through three-point N2 adsorption at P / Po values of 0.08, 0.14 and 0.20 at 77K after outgassing for 4 hours at 400°C in a flow of nitrogen using a Micromeritics Tristar 3030 analyzer). The powder was pressed at a pressure of 2000 kg / cm2 to pellets. The sample Al was a solid solution of CeO2 and ZrO2. Due to peak broadening the crystallographic phase composition analysis revealed the sole presence of either a cubic or tetragonal phase, or a combination thereof. It was established that no m-ZrO2 was present. - Sample A2: Sample A2 was obtained from Norpro Saint Gobain. The sample A2 had commercial code SZ61152 and was obtained as pellets. The sample contained 97 wt. % zirconia and 3 wt. % SiO2. The crystallographic phase distribution was determined through powder X-ray Diffraction (XRD) and Rietveld refinement analysis. Zirconium oxide was present only in the tetragonal phase. The BET surface area of the sample A2 was 140 m2 / g (the same method for determining the BET surface as mentioned above was used). - Sample A3: Sample A3 was obtained from Norpro Saint Gobain. The sample had commercial code SZ3*163 and was obtained as pellets. The sample was a 99%+ pure zirconium oxide. As should be noted, pure zirconium is substantially free (i.e., <0.01wt.%) of metals from Groups 8, 9, and 10 of the Periodic Table of Elements. The crystallographic phase composition of the zirconium oxide present was determined (again by XRD) to be 98 wt. % monoclinic zirconia and 2 wt. % tetragonal zirconium oxide. The surface area of the sample was 53 m2 / g. - Sample A4 Sample A4 was obtained from Norpro Saint Gobain. The sample had commercial code SZ3*164 and was obtained as pellets. The sample was a 99wt.%+ pure zirconium oxide. The crystallographic phase composition of the zirconium oxide present was determined (by XRD) to be 95 wt. % monoclinic zirconium oxide and 5 wt. % tetragonal zirconium oxide. The surface area of the sample was 92 m2 / g . - Sample A5 Sample A5 was obtained from Norpro Saint Gobain. The sample had commercial code SZ3*107 and was obtained as pellets. The sample contained 95 wt. % zirconium oxide and 5 wt. % silicium oxide. The crystallographic phase composition of the zirconium oxide present was determined (by XRD) to be 95 wt. % monoclinic zirconium oxide and 5 wt. % tetragonal zirconium oxide. The surface area of the sample was 108 m2 / g. Performance testing of samples A1-A5 in RWGS The catalytic activity and selectivity in catalytic RWGS of samples A1-A5 were tested in a small-scale catalyst screening unit. The catalyst screening unit was a so-called Flowrence testing unit available from Avantium. The unit consisted of a gas feed section providing a common feed stream of H2, CO2 and Ar to a glass chip that equally divided the flow to 16 reactors in parallel. The reactors consisted of a metal reactor block with 16 reactors in a metal block. The reactors hold an inner reactor tube in which the catalyst was positioned. The innertube was a non-porous alfa-alumina tube with an internal diameter of 2 mm. In the reactor a purge flow of nitrogen was added to flush the annular space between innertube and metal reactor wall and dilute the product gas downstream the inner tube. The diluted product gas stream was directed to a backpressure control system and subsequently to stream-selection valves to direct the individual product streams to an online GC. In the GO, the concentrations of Ar, CO, CO2, CH3OH, CH4 and N2 were determined. The CO2 conversion was determined based on the ratio of CO2 and Ar as compared to the feed gas. The feed gas composition was measured using the composition of a reactor in which no catalyst was positioned. No detectable CO2 conversion was measured in this empty reactor in the temperature range of 200-530°C. The product selectivity was determined based on the distribution of the carbon products detected. No other products than carbon monoxide, methane and trace levels of CH3OH were detected. The catalyst loading of each sample was 100 mg. The gas feed gas composition was 31 vol.% CO2, 61 vol.% H2 and 8 vol.% Ar. The flow rate per reactor was 17 Nml / min. The reactor pressure was 30 barg. The samples were heated to 300°C in a flow of nitrogen. At 300°C, the feed gas with above mentioned flow, composition and pressure was applied. The temperature of the reactor was raised to 530°C and the reactors were equilibrated for 100 hrs. CO2 conversion and product selectivity were determined for each reactor. The thermodynamic equilibrium conversion of CO2 via the RWGS reaction (in absence of methanation) at this condition was calculated using HSC Chemistry 7.1 and was 44% . As some catalysts were showing near-thermodynamic equilibrium conversion, the temperature was decreased to 470°C using the same composition, flow and pressure. At this temperature, the equilibrium CO2 conversion was 38%. The CO2 conversion was determined between 115-120 hours time-on-stream. For some samples, only small amounts of methane were detected next to CO as main product. The CO2 conversion and selectivity to unwanted methane formation at both conditions are summarized in Tables 1 and 2 below. Conversion and selectivity levels are presented on a molar basis. Table 1. CO2 conversion and selectivity to CH4 at 530°C. Sample Monoclinic / Tetragonal fraction [wt. %] CO2 conversion [%] at 530° CH 4 Selectivity [%] at 530°C Al not applicable 43 0.08 A2 0 / 100 12 0.18 A3 98 / 2 43 0.003 A4 95 / 5 43 0.004 A5 95 / 5 38 0.07 10 Table 2. CO2 conversion and selectivity to CH4 at 470°C. Sample Monoclinic / Tetragonal fraction [wt. %] CO2 conversion [%] at 470° CH 4 Selectivity [%] at 470°C Al not applicable 21 0.02 A2 0 / 100 4 0.10 A3 98 / 2 27 0.000 A4 95 / 5 27 0.006 A5 95 / 5 18 0.04 15 Tables 1 and 2 above clearly show superior CO2 conversion of samples A3 and A4 containing >95 wt. % of monoclinic zirconium oxide over sample A2 containing zirconium oxide in the pure tetragonal phase. Another important aspect is the selectivity to methane, which is an unwanted product in the RWGS reaction. Samples A3 and A4, both containing at least 95 - 21 -wt. % of zirconia in monoclinic phase, outperformed the other samples in terms of avoiding CH4 make, with sample A3 showing the lowest or no CH4 formation at 530 and 470°C, respectively. Albeit good in terms of activity and selectivity and clearly outperforming sample A2, sample Al based on the mixed cerium / zirconium oxide was lower in performance than samples A3 and A4. It has to be noted here that this performance data was determined in the absence of cofeeding H2S . Further samples, closely resembling Al, A3 and A4 were subjected to tests in the absence and presence of H2S in further testing at somewhat larger scale in Example 2. Example 2 This Example 2 provides a comparison of various samples with and without H2S dosing. - Sample Bl. Mixed cerium / zirconium oxides For the preparation of catalyst Sample Bl a mixed cerium / zirconium oxide powder was obtained from Solvay (Actalys) . The powder contained 69 wt. % CeO2 balanced with ZrO2 (and traces of HfO2) and had a BET surface area of 108 m2 / gram (the same method for determining the BET surface as mentioned above was used). The powder was pressed at a pressure of 2000 kg / cm2, crushed into a sieve fraction of 40-80 mesh and calcined in a static oven in air at 700°C for 4 hours. After equlibration at ambient conditions the water pore volume of the sample was determined before and after drying and was 0.20 ml / g and 0.25 ml / g, respectively. This catalyst sample Bl was partially used for further performance testing. Prior to testing, the sample was sieved to 60-80 mesh. Another fraction of the sample was used for impregnation of metal salts to obtain metal-doped cerium / zirconium mixed oxide samples. - Samples B2, B3 and B4. Metal-doped cerium / zirconium mixed oxide samples Catalyst samples B2, B3 and B4 were prepared by impregnation of aqueous metal solutions on portions of catalyst Bl. For each sample, 5 gram of sample Bl was used. For each sample a 1 ml metal solution was prepared by weighing the amount of metal salts corresponding to the desired metal weight loading on the final catalyst and adding demineralized water to a volume of 1.0 ml. This solution was homogenized and added to the sieve fraction. The impregnated sieve fractions were equilibrated on a roller bench for 2 hours. Subsequently, the samples were dried using a fluid bed dryer (Sherwood Tornado Model 501) at 65°C for 15 minutes. The dried samples were thereafter calcined in an oven (Carbolite GPC 1300 (13 / 65)) at 700 °C for 4 hours. The obtained samples were further crushed and sieved to 60-80 mesh prior to testing. Sample B2 contained 0.1 wt. % Ni on the mixed cerium / zirconium oxide and was prepared using nickel nitrate as metal salt. Sample B3 contained 0.1 wt. % Rh and 0.1 wt. % Pt and was prepared by using RhCla and H2PtCle as metal salts. Sample B4 contained 0.01 wt. % Rh and 0.01 wt. % Pt and was prepared by using RhCla and H2PtCle as metal salts. - Sample B5. Monoclinic zirconium oxide Monoclinic ZrO2 was obtained from Norpro Saint Gobain (commercial grade SZ3*164) in pellet form. The pellets were crushed and sieved to 60-80 mesh before use. Performance testing of catalyst samples B1-B5 in RWGS The catalytic activity of catalyst samples B1-B5 was tested in a so-called microflow testing unit. This unit consisted of a gaseous feed preparation section using thermal mass flow controllers to feed a gaseous stream of hydrogen, carbon dioxide and optionally a diluted feed of H2S in N2 to a tubular fixed bed reactor made of alumina. The length of the reactor was 50 cm and the internal diameter was 2.9 mm. The mesh fractions of the catalyst samples were placed in the isothermal zone of the reactor by positioning the catalyst bed between 2 solid rods of alfa-alumina of a length of 15 cm and 2 mm diameter. A catalyst weight was chosen such to fill 3.3 ml of catalyst in the reactor. The reactor is heated by an external furnace. The pressure in the reactors was controlled using a back pressure regulator at a pressure of 30 barg. The product stream was directed to an online GC to measure the gas composition. A sequence of gas feed compositions and temperature conditions was applied. First a feed of 1.9 Nl / hr CO2 and 1.65 Nl / hr H2 was fed to the reactor. The pressure was controlled at a level of 30 barg. The reactor was heated in three steps at 400°C, 470 and 500°C and equilibrated at each temperature for 48 hours. Then the temperature was set at 470°C and after 4-8 hours the product gas composition was measured. The calculated CO2 conversion, CO yield and CH4 yield of samples B1-B5 are listed in Table 3 below. Subsequently, while maintaining a temperature of 470°C and 30 barg pressure, 0.17 Nl / hr of a diluted H2S in N2 stream was added extra to the feed to provide a gas feed comprising 230 ppbv H2S . After at least 2 hours equilibration, the product gas composition was analyzed to assess the effect of H2S dosing on CO2 conversion and product selectivity. Only CO and CH4 were detected as products; no other hydrocarbons or oxygenate compounds were detected. The calculated CO2 conversion, CO yield and CH4 yield of samples B1-B5 in the presence of H2S are listed in Table 4 below. Conversion and yield are provided on a molar basis, i.e. the yield is the product of molar conversion and selectivity to the corresponding products. Table 3. Performance of samples B1-B5 without H2S dosing. Sample Catalyst amount [g] CO2 conversion CO yield CH4 yield Bl 0.55 13.9 13.89 0.01 B2 0.53 13.1 12.93 0.17 B3 0.55 22.9 10 . 85 12.05 B4 0.55 16.7 15.03 1. 67 B5 0.39 20.6 20.60 0.001 Table 4. Performance of samples B1-B5 while dosing 230 ppbv H2S Sample Catalyst amount [g] CO2 conversion CO yield CH4 yield Bl 0.55 13.5 13.50 0.004 B2 0.53 12.9 12.85 0.05 B3 0.55 23.8 16.87 6.93 B4 0.55 16.7 16.07 0.63 B5 0.39 20.3 20.30 0.000 15 Clearly, whilst comparing Tables 3 and 4, H2S had no significant effect on the CO2 conversions of samples Bl, B2 and B5. On the non-metal promoted samples Bl and B5 CO2 conversions were only marginally reduced, which was in line with the 5% increase in flow rate due to dilution with H2S / N2 feed. Sample B5 clearly outperformed Bl in terms of CO2 conversion. Also, an already extremely low methane make was further suppressed to below the detection limit in the presence of H2S. Sample B2, a low Ni-promoted CeZrO2, did not show benefits over the non-metal promoted sample Bl. In the absence of H2S some methane was formed which was only partially suppressed by H2S. Samples B3 and B4, with varying amounts of Pt / Rh on CeZrO2, did show a higher CO2 conversion as compared to Bl. However, significant amounts of unwanted CH4 were formed. The presence of H2S only partially reduced the CH4 make. In summary, the purely oxidic (e.g., substantially free of metals from Groups 8, 9, and 10 of the Periodic Table of Elements) samples Bl and B5 can be considered sulphur-tolerant with extremely low selectivity to methane and thus extremely good selectivity to CO which was even further improved by the addition of H2S. Monoclinic zirconia outperformed the mixed CeZrO2 in terms of activity. Example 3 This Example 3 provides information on the long duration stability testing of monoclinic zirconium oxide catalyst samples. - Sample C. Monoclinic zirconium oxide Monoclinic ZrO2 was obtained from Norpro Saint Gobain (commercial grade SZ3*164) in pellet form. Pellets with a size between 3-8 mm used. This zirconium oxide contained 98 wt. % monoclinic zirconium oxide and 2 wt. % tetragonal zirconium oxide. The catalytic activity of catalyst samples was tested in a so-called bench scale testing unit. This unit consisted of a gaseous feed preparation section using thermal mass flow controllers to feed a gaseous stream of hydrogen and H2S in N2 and a Cori-flow to feed a gaseous stream of CO2 to a tubular fixed bed reactor made of Alloy601. The length of the reactor was 107 cm and the internal diameter was 20 mm. The catalyst sample was placed in the isothermal zone of the reactor by positioning the catalyst in between a bed of inert alfa-alumina rings (8x8 mm, ID 2 mm), 3 cm below and 20 cm on top of the catalyst. This was placed on an alfa-alumina rod with a length of 13 cm and 2 mm diameter supported on a cold finger of 20 cm with an outside diameter of 12.7 mm. A thermowell with an outside diameter of 6.4 mm was inserted from the top of the reactor through the catalyst bed into the alpha-alumina rod. A catalyst weight of 66 grams was used to fill 20 cm of catalyst bed height in the reactor. The reactor was heated by an external furnace. The pressure in the reactors was controlled using an equilibar pressure regulator at a pressure of 24 barg. The product stream was directed to an online GC to measure the gas composition. A sequence of gas feed compositions and temperature conditions was applied. - Example 3A. Operation without H2S dosing. A feed of 200 Nl / hr CO2, 200 Nl / hr H2 and 20 Nl / hr N2 was fed to the reactor. The pressure was controlled at a level of 24 barg. The reactor was heated to 520°C; after 125 hours the selectivity to CH4 started to increase. After 255 hours time-on-stream, the temperature was temporarily decreased to 470°C for a period of 50 hrs and after that the temperature was increased to 520°C. Figure 4 shows the CO2 conversion and CH4 selectivity as a function of time in the absence of H2S. Open symbols in Figure 4 represent datapoints at 520° and closed symbols those measured at 470°C. Clearly, methane formation continued to increase and after 550 hr the test was stopped. - Example 3B. Operation with H2S dosing. A new reactor with fresh catalyst was reloaded. Now a feed of 200 Nl / hr CO2, 200 Nl / hr H2 and 4 Nl / hr 5ppm H2S in N2 was fed to the reactor. Accordingly, 100 ppb H2S in the combined feed was fed to the reactor. The pressure was controlled at a level of 24 barg. The reactor was heated to 520°C and kept stable for ~1000 hour. During this time, methane make was close to the detection limit and remained at extremely low levels for 1000 hrs. The CO2 conversion level was close to level as when dosing no H2S . Figure 5 shows the CO2 conversion and CH4 selectivity as a function of time in the presence of 100 ppb H2S. Clearly, the dosing of H2S had no negative effect on the CO2 conversion, whilst it effectively suppressed unwanted methane make on catalyst and / or reactor walls. Discussion As can be seen from the above Examples, the method according to the present invention allows for an effective way of producing syngas using a catalytic RWGS reaction . The person skilled in the art will readily understand that many modifications may be made without departing from the scope of the invention.

Claims

1. A method for producing syngas using a catalytic reverse water gas shift (RWGS) reaction, the method at least comprising the steps of:a) providing a feed stream (10) comprising at least hydrogen (H2) and carbon dioxide (CO2) ;b) adding a sulphur-containing compound (15) to the feed stream provided in step a), thereby obtaining a sulphur-enriched feed stream (20) wherein the sulphur-containing compound (15) is added in an amount of from 50 to 2000 ppbv, preferably between 100 and 500 ppbv, based on the feed stream (10);c) introducing the sulphur-enriched feed stream (20) obtained in step (b) into a RWGS reactor (2) and subjecting it to a catalytic RWGS reaction in the presence of a catalyst, thereby obtaining a syngas containing stream (30), wherein the catalyst comprises at least 60wt.% monoclinic zirconia (m-ZrO2) , based on the total weight of zirconia, and wherein the catalyst is substantially free of metals of Groups 8, 9, and 10 of the Periodic Table of Elements.

2. The method according to claim 1, wherein the sulphur-containing compound (15) added in step b) is H2S.

3. The method according to any one of the preceding claims, wherein the catalyst as used in step c) further comprises cerium oxide.

4. The method according to claim 3, wherein the catalyst comprises at least 70 wt.%, more preferably at least 80 wt.%, even more preferably at least 90 wt.%, yet even more preferably at least 95 wt.%.- 29 -5. The method according to claim 1, wherein the catalyst comprises at least 60 wt. % m-ZrO2, based on the total weight of the catalyst, preferably at least 70 wt.%, more preferably at least 80 wt.%, even more preferably at least 90 wt.%, yet even more preferably at least 95 wt.%.

6. The method according to any one of the preceding claims, further comprising the steps of:d) cooling the syngas containing stream (30) obtained in step c) thereby obtaining a cooled syngas stream (40,50); e) separating the cooled syngas stream (40,50) obtained in step d) in a gas / liquid separator (6) thereby obtaining a water-enriched stream (60) and a water-depleted syngas stream (70);f) adding a sulphur-containing compound (15) to the water-depleted syngas stream (70) obtained in step e) , thereby obtaining a sulphur-enriched syngas stream (75); g) introducing the sulphur-enriched syngas stream (75) obtained in step f) into a second RWGS reactor (9) and subjecting it to a second catalytic RWGS reaction, thereby obtaining a second syngas containing stream (100) .

7. The method according to any one of the preceding claims, wherein the syngas containing stream (30) obtained in step (c) and / or the second syngas containing stream (100) obtained in step (g) is subjected to a sulphur-removal step.

8. An RWGS catalyst, at least comprising:- monoclinic zirconia (m-ZrO2) , wherein the catalyst is substantially free of metals from Groups 8, 9, and 10 of the Periodic Table of Elements.

9. The RWGS catalyst according to claim 8, wherein the catalyst comprises at least 60 wt.% m-ZrO2, based on the total weight of the zirconia, preferably at least 70- 30 -wt.%, more preferably at least 80 wt. %, even more preferably at least 90 wt.%, yet even more preferably at least 95 wt. % .

10. The RWGS catalyst according to claim 8 or 9, wherein 5       the catalyst comprises at least 60 wt.% m-ZrO2, based onthe total weight of the catalyst, preferably at least 70 wt.%, more preferably at least 80 wt.%, even more preferably at least 90 wt.%, yet even more preferably at least 95 wt. % .10

Images