EA040440B1 - Catalyst for catalytic oxidative cracking of hydrogen sulfide with simultaneous production of hydrogen - Google Patents

Description

Scope of the invention

The invention relates to a catalyst suitable for a catalytic oxidation process for the recovery of sulfur from a gas stream containing H2S. The invention also relates to a process for recovering sulfur from said gas stream using a catalyst. In addition, the invention relates to the production of hydrogen associated with the process of extracting sulfur using a catalyst.

Prerequisites for the creation of the invention

The sulfur recovery units are designed to remove H2S from acid gases containing H2S from amine recovery systems and from acid water strippers to produce sulfur, a non-toxic product that can be stored and sold in liquid or solid form to various users for various industries. applications. Acid gases from amine recovery systems and acid water strippers containing variable amounts of H ₂ S are treated in a sulfur recovery unit (SRU) operating essentially on a modified Claus process to recover the bulk of the sulfur and then in a treatment section residual gas (TGT) for deep sulfur recovery. The Claus section breaks down other impurities found in acid gases, including ammonia and hydrocarbons.

The modified Claus process itself recovers about 94-96% (2 catalytic steps) or 95-98% (3 steps) of sulfur from the feedstock. Thus, when a higher sulfur recovery (SRE) is required, additional Claus tail gas treatment is necessary.

The modified Claus process involves substoichiometric combustion of an acid gas stream in a thermal reactor (thermal stage) followed by catalytic conversion in Claus reactors (catalytic stage). In the Claus section, one third of the total H2S is oxidized to SO2, which reacts with the remaining H2S to form sulfur and water according to the following reactions:

H2S + 1.5 O2 H2O + SO2 (oxidation reaction) (1)

H2S + SO2 o 1.5 S ₂ + 2 H2O (Claus reaction) (2)

H2S + 1.5 O2 θ 3 H2O + 1.5 S2 (general reaction) (3)

The goal of the process is to bring the entire reaction to almost complete completion. The H2S contained in the acid gas is burned in a Claus thermal reactor on a special burner using air (or in some specific cases oxygen-enriched air), while only one third of the total H ₂ S is oxidized to SO2, and the remaining two thirds are not involved. in reaction. The total amount of air is equal to the exact amount sufficient to oxidize one third of the total H2S and completely oxidize all hydrocarbons and ammonia contained in the feed; therefore the molar ratio of H ₂ S / O2 in the feed is about 2:1 with the possibility of ensuring the ratio of H ₂ S / SO2 in the tail gas of the Claus process is exactly 2:1 or as close as possible to this value, and this is the stoichiometric ratio for the Claus reaction , thereby increasing the degree of sulfur recovery. During the combustion of acid gases, a small part of the H ₂ S (typically 5-7%) dissociates into hydrogen and sulfur according to the following reaction:

H2S θ Hg + 0.5 S2 (dissociation or cracking reaction) (4)

According to Clark et al., Alberta Sulfur Research Ltd. (ASRL), the formation of hydrogen also occurs according to the following reaction:

H2S + Or θ 2 Hg + 2 HgO + 2 S2 (Hg formation reaction) (5)

In addition, several side reactions occur that destroy ammonia and hydrocarbons, and also form carbonyl sulfide COS and carbon disulfide CS2. In order to complete the Claus reactions, it is necessary to provide a suitable residence time at high temperature in the thermal reactor.

The Claus thermal reactor is usually followed by a waste heat boiler in which the furnace effluent is cooled to about 300°C and the heat is recovered by high pressure rising steam and in a sulfur condenser in which the process gas is cooled to the sulfur dew point by means of rising low pressure steam, and liquid sulfur is separated.

The Claus thermal stage is usually followed by two or three catalytic stages, each of which involves the use of a gas reheater to bring the gas temperature to the optimum reaction temperature, a catalytic reactor in which the Claus reaction takes place, and a sulfur condenser in which the gas is cooled and condensed. and separate liquid sulfur. The Claus reaction is an exothermic equilibrium reaction, for which low temperatures are thermodynamically favorable. The first catalytic Claus reactor is partially filled with a Claus catalyst (based on alumina) to enhance the Claus reaction and partially filled with a special high conversion catalyst (based on titanium dioxide) to enhance the hydrolysis of COS and CS ₂ . The second and third catalytic Claus reactors, if present, are typically filled with a Claus catalyst (based on alumina) to enhance the Claus reaction.

In order to achieve a sulfur recovery >99%, which is usually required for a plant for

- 1 040440 sulfur recovery, after the Claus section is usually placed the tail gas treatment section. Over the years, several different alternative processes have been proposed to improve sulfur recovery, such as the SCOT process, the RAR process, the CBA process, the CLINSULF/DEGSULF process, or the BSR Selectox process. In a traditional reduction type tail gas treatment section, the process gas from the Claus section is preheated and combined with hydrogen from an external source before being fed into the hydrogenation reactor, in which all sulfur compounds are converted to H2S over a special reduction catalyst (based on Co and Mo oxides) catalyzing both a hydrogenation reaction and a hydrolysis reaction. The reaction product is cooled in a cooling tower by circulating steam condensate. The H2S produced in the hydrogenation reactor is recovered in the amine absorber with a special aqueous amine solution and returned to the Claus section from the top of the amine recovery tower where the rich solution is stripped off.

The tail gas from the amine absorber is sent to a thermal combustion furnace to oxidize the residual H2S and other sulfur compounds such as COS and CS2 to SO2 before being released to the atmosphere through a dedicated flue.

The main disadvantages of the traditional Claus plant are the need for bulky and expensive equipment with a very low economic value of sulfur, continuous emissions of SO _x (SO2 and SO ₃ ), CO, CO2, NO _x , as well as traces of H2S to the atmosphere and continuous transfer of hydrogen from the network for process gas recovery in the TGT section.

In some installations in locations where hydrogen is not available, such as gas fields, a reducing gas mixture is generated in a reducing gas generator by substoichiometric combustion of the fuel gas. The main disadvantage of this alternative configuration is that the equipment is even more bulky than the traditional Claus plant. This is due to a 10-15% higher process gas consumption due to high amounts of inert substances (mainly nitrogen from air and water and carbon dioxide from the combustion process) coming from the once-through combustion of fuel gas.

As mentioned in Clark, Catalysis Communications 5 (2004) 743-747, the industry has long been trying to solve the problem of extracting H ₂ from H ₂ S. Clark solves this problem by partially oxidizing H 2 S over alumina catalysts. It is believed that the key factor in this process is the stimulation of the reaction of H ₂ S and O ₂ with the formation of hydrogen, water and sulfur at a controlled temperature using an external furnace. The task of reducing emissions into the atmosphere has not been solved.

Several alternative processes have been proposed in previous years focusing on thermal or catalytic partial oxidation of H2S.

US Pat. a catalyst for the partial oxidation reaction of H2S to form sulfur and water using air or enriched air or pure oxygen with an H2S/O2 feed ratio of about 2:1, followed by a cooling zone and a sulfur condenser. The main purpose of the first patent is to desulfurize the gas stream, and the main purpose of the second patent is to offer an alternative solution to the traditional Claus thermal reactor. Both patents are based on the catalytic reaction of the partial oxidation of hydrogen sulfide with oxygen to form sulfur and water.

U.S. Patent No. 7,560,088 to the Conoco Phillips Company describes a process for removing sulfur from a gas stream containing H2S using a compact system comprising a partial catalytic oxidation zone in a flameless reactor with a short contact time followed by a temperature controlled zone, first a Claus catalytic reaction zone, a second temperature control zone, a first sulfur outlet, and a first off-gas outlet. The main purpose of this patent is to propose an alternative solution to the traditional Claus plant based on the partial catalytic oxidation of hydrogen sulfide to form sulfur and water.

US Pat. No. 4,481,181 to GA Technologies Inc. describes a process for removing sulfur and recovering hydrogen from a gas stream containing H ₂ S, which couples the thermal partial oxidation of H ₂ S to form sulfur and water and the thermal dissociation of H ₂ S into hydrogen and sulfur in one reaction zone, preceded by a feed heating section and followed by a cooling zone and a sulfur condenser, using pure oxygen and a significant proportion of nitrogen at a ratio of H ₂ S / O2 in the feed in the range from 10:1 to 25:1. The main purpose of this patent is the implementation of thermal decomposition of hydrogen sulfide by its partial oxidation and dissociation into sulfur and hydrogen.

In patent WO2010/036941 issued to Chevron U.S.A. Inc. and Drexel University, describes a process for conducting thermal dissociation of H2S at temperatures below 1600° C. using, in one embodiment, H and SH radicals over a suitable plasma catalyst.

In addition, Siirtec-Nigi Italian Patent No. 1203898 describes a process called HCR based on the use of a traditional Claus thermal reactor with a slightly higher H ₂ S / O ₂ ratio in the feed to maintain the H ₂ S / SO ratio. ₂ in the rest

- 2 040440 fine gas from the Claus process at a level well above 2:1. The main purpose of this process is to increase the production of hydrogen in the thermal reactor and prevent the transfer of hydrogen to the TGT section. In addition, emissions from the sulfur recovery plant cannot be avoided in said process.

It is evident from the above description that several attempts have been made in the past to provide an effective alternative to the traditional Claus installation. In particular, some processes proposed over the years are based on thermal or catalytic partial oxidation of H ₂ S, and some other processes are oriented towards thermal or catalytic cracking of H2S. None of the proposed processes is expected or attempted to implement the conversion of H2S to hydrogen and sulfur over a suitable catalyst capable of simultaneously favoring both reactions.

WO2012/154041 describes a process for producing hydrogen from a gas stream containing H2S, including catalytic oxidative cracking of the gas stream to form H2 and S ₂ . The purpose of the invention described herein is to solve the problem of gas emissions to the atmosphere while obtaining a valuable hydrogen export stream.

A problem with gas streams containing H2S, which are regularly fed to sulfur recovery plants, is the simultaneous presence of methane and/or heavier hydrocarbons. In particular, methane tends to be converted to sulfur compounds such as CS2 or COS, which is undesirable.

Another problem with gas streams containing H2S, which are regularly fed to sulfur plants, is the co-presence of ammonia. Ammonia is usually converted in the thermal stage of the Claus plant. However, the thermal conversion of ammonia is associated with the risk of formation of solid salts such as ammonium sulfides or sulfites as a result of incomplete combustion. These salts clog the coldest sections of the Claus plant, in particular in the sulfur condensers. Proper combustion of ammonia requires a homogeneous mixture of ammonia and air, as well as a high flame temperature. However, the formation of nitrogen oxides promotes the oxidation of sulfur dioxide, SO2, to sulfur trioxide, SO2. The Claus catalyst then becomes sulfonated, and the cold areas of the plant undergo noticeable corrosion.

WO2014/073966 proposes a catalyst that exhibits activity and selectivity in the oxidative cracking of H2S, in particular in the presence of NH ₃ and/or CH ₄ and more generally carbonaceous compounds. The catalyst contains iron and molybdenum on a support containing aluminum, such as aluminum oxide.

However, reducing the tendency to caking at high temperatures is still a problem in the field of catalytic oxidative cracking of hydrogen sulfide. In another aspect, it is desirable to improve the stability of the catalyst, for example, by reducing the vapor pressure of the active phases at high temperatures during the catalytic oxidative cracking of hydrogen sulfide. In one additional aspect, to meet increasingly stringent market requirements for end-of-life (EOR) conditions of the catalyst, higher stability is particularly desirable. In this regard, it is especially effective to reduce the concentration of ammonia at the outlet of the reactor.

If one or more of these problems can be solved, the catalyst will become more stable at high temperatures than described in the prior art. In some aspects of the invention, this improvement improves the conversion of H ₂ S during the reaction. In addition, higher stability at high temperatures can also increase the life of the catalyst.

As an additional background to the creation of the invention, reference is made to the publication of Jiratova et al. Chinese Journal of Catalysis, volume 37, pages 258-267. This relates to the hydrodesulfurization activity of NiMO catalysts supported on mechanochemically produced Al-CE mixed oxides. In this document, the active metal is NiMo ₆ .

Statement of the Invention

In order to best achieve one or more of the above objectives, in one aspect of the invention, there is provided a catalyst composition suitable for catalytic oxidative cracking of a gas stream containing H ₂ S, wherein the catalyst composition contains at least one active metal selected from the group consisting of Fe , Co, Ni and combinations thereof, on a carrier containing cerium dioxide and alumina, and preferably provided that the active metal does not contain NiMo ₆ . This condition is preferred for the catalyst product as such, and optionally for the catalyst used in the plant in the hydrogen production process, as well as for the catalyst obtained by the catalyst production process described herein. The catalyst contained in the plant used in the hydrogen production process or obtained in accordance with the catalyst production process contains NiMo ₆ in some embodiments and is different from NiMo ₆ in some other embodiments.

The invention also relates to a process for the production of hydrogen from a gas stream containing H2S, including catalytic oxidative cracking of the gas stream with the formation of H ₂ and S ₂ , the use of a catalyst containing at least one active metal selected from the group consisting of

- 3 040440 from iron, cobalt, nickel and combinations thereof, wherein said active metal is deposited on a carrier containing cerium dioxide and alumina.

In another aspect, the catalyst compositions of the invention are suitable for catalytic oxidative cracking of a gas stream containing H2S in the case of the simultaneous presence of NH ₃ and/or CH4 and/or carbonaceous compounds. In yet another aspect, the H ₂ S-containing stream may further contain compounds selected from the group consisting of COS, RSH, HCN, benzene, toluene, ethylbenzene, and xylene.

In another aspect, the invention relates to a process for preparing a catalyst as defined above or as used in said process comprising providing an aqueous solution of nickel precursors selected from the group consisting of nickel tetracarbonyl, nickel nitrates, nickel bromides, nickel chlorides, nickel fluorides, nickel phosphates , nickel sulfates, nickel acetylacetonates, nickel acetates, nickel fumarates, nickel gluconates, nickel citrates, nickel benzoates, nickel maleates, nickel oxalates, nickel oleates, nickel stearates, nickel ammonium complexes, iron tetracarbonyl, iron pentacarbonyl, iron nonacarbonyl, iron nitrates , ferrous bromides, ferrous chlorides, ferrous fluorides, ferrous phosphates, ferrous sulfates, ferrous acetylacetonates, ferrous acetates, ferrous fumarates, ferrous gluconates, ferrous citrates, ferrous benzoates, ferrous maleates, ferrous oxalates, ferrous oleates, ferrous stearates, ferrous stearates, stearates iron, complexes of iron ammonium, tetracarbonyl to cobalt, cobalt pentacarbonyl, cobalt nonacarbonyl, cobalt nitrate, cobalt bromide, cobalt chloride, cobalt fluoride, cobalt phosphate, cobalt sulfate, cobalt acetylacetonate, cobalt acetate, cobalt fumarate, cobalt gluconate, cobalt citrate, cobalt benzoate, cobalt oleates, cobalt stearates and cobalt-ammonium complexes.

In another additional aspect, the invention provides a process for producing hydrogen from a gas stream containing H2S and optionally NH ₃ and/or CH4, comprising catalytic oxidative cracking of the gas stream to form H2 and S2 using a catalyst as defined above.

The invention also relates to a plant suitable for conducting catalytic oxidative cracking of a gas stream containing H ₂ S, and the specified plant contains an inlet for an acid gas stream containing H ₂ S, an inlet for an oxygen-containing stream and a catalytic oxidative cracking reaction zone containing catalytic material , suitable for partial oxidation and cracking of H ₂ S, and the catalytic material contains one or more catalysts containing at least one active metal selected from the group consisting of Fe, Co, Ni and combinations thereof supported on a carrier containing cerium dioxide and aluminum oxide; and the catalyst preferably corresponds to the above definition.

Brief description of graphic materials

The drawing shows schematically an example of an installation according to the invention.

Detailed description of the invention

The invention is broadly based on the recognition of a reasonable combination of at least one catalytically active metal selected from the group consisting of nickel (Ni), iron (Fe), cobalt (Co) and mixtures thereof, supported on a carrier containing cerium dioxide and oxide aluminum. The catalyst composition preferably contains Ni, more preferably the active metal is Ni, i. e. the only active metal was nickel in any suitable form such as sulfide.

The carrier containing cerium dioxide and alumina is believed to reduce the tendency to caking. Surprisingly, the compositions of the invention have been found to reduce the vapor pressure of an active phase containing an active metal, preferably in the form of its sulfide, under the desired reaction conditions. Without being bound by any theory, it is believed that by reducing the sintering tendency and/or the vapor pressure of the active phase, thermal stability improves. In turn, it is believed that higher thermal stability improves catalytic efficiency because higher conversions can be achieved. In addition, with higher thermal stability, the service life of the catalysts also increases.

In another aspect, the catalysts of the invention are considered advantageous in terms of end-of-cycle conditions. Compared to previously described catalysts operating for a similar period of time, using the catalysts of the invention can reduce the concentration of ammonia at the outlet of the reactor. Or in an alternative embodiment that uses the same concentration of ammonia obtained using the previously described catalysts, the new catalysts remain stable for a longer period of time.

It should be understood that the term supported catalyst refers to a catalyst composition containing a catalytically active part (i.e. particles that are either active or converted into an active phase in situ) and a catalytically inactive part, and the catalytically inactive part (carrier) essentially makes up the majority of the catalyst. This distinguishes a supported catalyst from a bulk catalyst, in which the catalytically inactive portion is generally negligible. Thus, in the supported catalyst, the catalytically inactive

- 4 040440 part is usually more than 50 wt.% of the catalytic composition. The carrier preferably constitutes more than 60 wt.%, more preferably more than 70 wt.% and most preferably more than 80 wt.% of the entire catalyst composition. In some embodiments, the implementation of the carrier is not more than 90 wt.% of the entire composition, more preferably not more than 87 wt.%, most preferably not more than 83 wt.%.

The catalytically active part of the catalytic composition contains at least one element 4 period group VIIIB. In particular, the active metal is selected from the group consisting of Fe, Co and Ni. The metal is typically present in the form of particles dispersed on the carrier, in an amount of at least 1 wt.% and usually up to 50 wt.% of the catalyst composition. The support contains alumina, ie. (Al2O ₃ ). The alumina may be, for example, alpha, theta or gamma alumina. In addition, the carrier contains cerium dioxide (CeO2).

The support typically contains cerium dioxide and alumina in a weight ratio of cerium (i.e.

mass of cerium dioxide mass of cerium dioxide + mass of aluminum oxide

X 100% )

2 to 50%, preferably in a ratio of 5 to 40%, more preferably 10 to 30%.

The catalytically active metal in the form of its sulfide is present in the catalyst composition, typically in the range of 1 to 40% by weight, preferably 3 to 35% by weight, more preferably 5 to 30% by weight, most preferably 8 to 25% by weight. wt.%, for example 17 wt.% compared to the total weight of the catalyst composition when the catalyst is present in powder form, or compared to the total weight of the catalyst layer when the catalyst contains a mechanical support.

In some embodiments, a non-sulfided catalyst may be loaded into an appropriate reaction chamber and the active metal may be sulfided by keeping the catalyst in the gas stream to be treated for a certain amount of time, preferably 3 to 12 hours, to sulphurize the active metal and activate the catalyst.

The catalyst may consist essentially of a catalyst composition, ie. carrier containing aluminum oxide and cerium dioxide, as well as the active metal contained therein. In such a case, the catalyst will essentially be in an appropriate form, such as powder, spheres or granules. In addition to the catalytic composition containing a carrier or support and active metals, the catalyst may also contain a mechanical support structure, i.e. substrate.

It should be understood that such a support does not form part of the catalyst composition as defined above, but is in addition to it. The support may be any structure known in the art as a support for catalysts. In one embodiment of the present invention, the support may be in the form of microspheres, pellets, spheres, honeycomb monolith, or open cell foams. The support may be formed from alumina, silica, silica, titanium dioxide, mixtures thereof, or any other suitable material available in the art of catalytic substrates.

If the catalyst contains a support, it is usually applied to the catalyst composition consisting of alumina, cerium dioxide and at least one active metal, as defined above.

In a preferred embodiment, the catalytically active metal is in the form of its sulfide. For example, nickel is preferably in the form of nickel sulfide, iron is preferably in the form of iron sulfide, and cobalt is preferably in the form of cobalt sulfide.

The catalyst composition of the invention can be obtained by a method known to a person skilled in the art. Reference is made to, for example, Catalyst Handbook, M.V. Twigg (Ed.), Wolfe Publishing Ltd, 1989, and Structured Catalysts And Reactors, A. Cybulski and J.A. Moulijn (Eds.), Marcel Dekker Inc., 1998 - chapter 21 (Transformation of a structured carrier into structured catalyst), pp. 599-615.

In the most suitable manner, an aqueous solution of the precursor is prepared and this solution is dispersed onto the carrier material. Examples of nickel-containing precursors are inorganic and organic nickel salts, nickel chelates, nickel clusters, nickel hydroxides and oxyhydroxides, and nickel organometallic complexes. Examples of such compounds are nickel tetracarbonyl, nickel nitrates, nickel bromides, nickel chlorides, nickel fluorides, nickel phosphates, nickel sulfates, nickel acetylacetonates, nickel acetates, nickel fumarates, nickel gluconates, nickel citrates, nickel benzoates, nickel maleates, nickel oxalates, oleates. nickel, nickel stearates, nickel-ammonium complexes, and the like. In one aspect, the catalyst precursors contain at least one metal selected from the group consisting of Ni(0), Ni(I), Ni(II), Ni(III) and Ni(IV). Ni(II) is preferably used. Examples of iron-containing precursors are inorganic and organic iron salts, iron chelates, iron clusters, iron hydroxides and oxyhydroxides, and iron organometallic complexes. Examples of such compounds are iron tetracarbonyl, iron pentacarbonyl, iron nonacarbonyl, iron nitrates, iron bromides, iron chlorides, iron fluorides, iron phosphates, iron sulfates, iron acetylacetonates, iron acetates, fumarates.

- 5 040440 iron, ferrous gluconates, ferrous citrates, ferrous benzoates, ferrous maleates, ferrous oxalates, ferrous oleates, ferrous stearates, ferrous stearates, ferrous stearates, iron ammonium complexes and the like. In one aspect, the catalyst precursors contain at least one metal selected from the group consisting of Fe(0), Fe(I), Fe(II), Fe(III), Fe(IV), Fe(V), Fe( VI) and Fe(VII). It is preferable to use Fe(II) and/or Fe(III). Examples of cobalt-containing precursors are inorganic and organic cobalt salts, cobalt chelates, cobalt clusters, cobalt hydroxides and oxyhydroxides, and cobalt organometallic complexes. Examples of such compounds are cobalt tetracarbonyl, cobalt pentacarbonyl, cobalt nonacarbonyl, cobalt nitrates, cobalt bromides, cobalt chlorides, cobalt fluorides, cobalt phosphates, cobalt sulfates, cobalt acetylacetonates, cobalt acetates, cobalt fumarates, cobalt gluconates, cobalt benzoates, cobalt citrates. cobalt, cobalt oxalates, cobalt oleates, cobalt stearates and cobalt complexes, and the like. In one aspect, the catalyst precursors contain at least one metal selected from the group consisting of Co(0), Co(I), Co(II), Co(III), Co(IV) and Co(V). Preferably, Co(II) is used. The catalyst precursors may additionally contain organic ligands or anions such as acetate, citrate, EDTA (ethylenediaminetetraacetate) or NTA (nitrilotriacetate).

In the most preferred embodiment, the support is first calcined and then impregnated with the precursor solution. The calcination is preferably performed at a temperature above 700°C, more preferably at least 750°C, even more preferably in the range of 750-1100°C, most preferably in the range of 900-1100°C. The duration of the calcination process at the desired temperature is preferably in the range of 10 to 30 hours, more preferably in the range of 13 to 30 hours. The calcination process is preferably carried out under isothermal conditions. The calcination is preferably carried out in the presence of oxygen, more preferably in air. After calcination, an active metal catalyst is applied to the calcined carrier as described above. For example, a wet impregnation method or catalytic metal deposition can be used. Without being limited by any theory, it is believed that the calcination described above can stabilize the structure of the catalyst. Thus, the resulting catalyst is particularly suitable for high temperatures, for example 1100° C., used in catalytic oxidative cracking. In this embodiment, it is important to first carry out the calcination step to obtain a stabilized carrier, and then deposit the catalytic metal thereon. It is believed that if the procedure is reversed, the catalytic structure may change such that the catalytic metal is not available for the catalytic reaction.

After precipitation of the active phase, another calcination step is preferably carried out. In some aspects, this additional calcination step may be carried out under the same conditions and in accordance with the same procedures as mentioned above for the first calcination step.

The invention further relates to a process for producing hydrogen from a gas stream containing H2S, comprising catalytic oxidative cracking of the gas stream to form H2 and _S2 using a catalyst as defined above.

It should be emphasized that catalytic oxidative cracking in accordance with this aspect of the invention is a fundamentally different process, different from both the thermal stage and the catalytic stage of the existing Claus process. In accordance with the above reaction equations (1)-(5), the Claus processes are aimed at bringing the above reaction (3) to almost complete completion. The present invention is based on the understanding of providing a process based on side reactions (4) and (5) and facilitating these reactions to produce hydrogen and sulfur from an H2S containing gas stream.

The invention process is also fundamentally different from the recent proposals by Clark et al. References to the publications of these authors are based on the theory of direct oxidation of H2S with the formation of hydrogen, water and sulfur. With the resulting conversion, not only can the formation of SO2 be avoided, but the process of H2S conversion and the simultaneous production of sulfur and H2 can also be improved.

In the present invention, a catalytic oxidative cracking (COC) step replaces the thermal Claus process step and/or both the thermal Claus process step and the catalytic Claus process steps. Thus, the process of the invention favors H ₂ S dissociation and partial oxidation instead of complete oxidation and the Claus reaction. However, the addition of a thermal Claus step and/or a catalytic Claus step after the COC step is not excluded.

Catalytic oxidative cracking is carried out in one or more reaction zones, preferably located in the same reaction chamber. In the text of this document, the term chamber may refer to one or more reaction zones. The reaction chamber is defined as the volume of the reactor with an optional catalyst bed. There is only one type of catalyst in one reaction chamber. As a rule, the reaction chamber is essentially cylindrical, and the reactant flow moves in the axial direction. If the reaction chamber contains a catalyst bed, one or more reactions may take place in the axial direction of the gas flow. In an embodiment in which more than one reaction occurs, the reaction conversion profile for one reaction may be different from that for another reaction. In other words, one reaction can proceed, for example, advantageously

- 6 040440 vein at the beginning of the catalyst layer, and another reaction, for example, can proceed along the entire length of the catalyst layer.

The invention provides a person skilled in the art with an explanation of how to carry out the above reactions (4) and (5). Due to the fact that the gas stream is subjected to catalytic oxidative cracking, the person skilled in the art clearly understands how to carry out this process.

Typically, the catalyst will be conventionally provided in a catalyst bed over which a stream of gas to be purified is passed. The choice of layer types and their volumes is directly within the skill of the ordinary person skilled in the art.

The reaction zone or zones of catalytic oxidative cracking provide oxygen. The oxygen is preferably provided as a gas that is more oxygenated than air. The oxygen source is preferably an oxygen-containing gas stream containing at least 40% oxygen by volume, preferably at least 60% oxygen by volume. More preferably, such oxygen is provided in the form of substantially pure oxygen, namely 90 to 99% oxygen by volume or substantially 100% depending on the available source.

The use of an oxygen-enriched gas and preferably pure oxygen is associated not only with the optimization of the catalytic oxidative cracking process, but also with such advantages as the elimination of excessively bulky equipment that would be required given the presence of large volumes of inert gas (nitrogen). Moreover, in the context of the objective of the invention, which in addition to sulfur recovery and emission reduction is to produce hydrogen, it will be beneficial to reduce and preferably prevent the presence of nitrogen in the tail gas of the process.

The amount of oxygen supplied to the reactor is chosen such that the ratio of H2S/O2 in the feed is above the typical value of about 2:1. The H2S/O2 ratio of the feed should preferably be in the range of 2:1-6:1 or preferably above 2:1 and up to 6:1, more preferably in the range of 3:1-5:1, even more preferably in the range of 3.5 :1-4.5:1, even more preferably in the range of 3.9:1-4.7:1, most preferably in the range of 4:1-4.5:1.

In a preferred embodiment of a catalytic oxidative cracking process based on a H2S/O2 ratio of 4:1 to 4.5:1, most preferably 4.1:1 to 4.5:1, the preferred reaction temperatures for simultaneous cracking and partial oxidation H ₂ S are in the range of 700-1300°C, preferably in the range of 800-1300°C, more preferably in the range of 950-1250°C. Most preferably, a temperature in the range of 1050-1100°C is used.

In one embodiment, the molar ratio of NH3/O2 in the feed is from 0.1 to 1.5. Preferably the molar ratio of NH3/O2 in the feed is 0.1-1.2, more preferably 0.11, most preferably 0.3-1.

In one embodiment, the feed (containing H2S acid gas and oxygen containing gas) fed to the reaction zone or zones of catalytic oxidative cracking is preheated to increase the reaction temperature and to increase hydrogen production and reduce SO2 formation.

In one embodiment of the present invention, the H ₂ S-containing acid gas and the oxygen-containing gas are mixed in a static mixer just prior to being fed into the catalytic bed of the catalytic oxidative cracking reaction zone or zones.

In one embodiment, the hydrogen concentration in the product at the outlet of the reaction chamber (after quenching) is at least 3% vol, preferably at least 5% vol, most preferably at least 7% vol.

It should be noted that the reaction is preferably carried out in autothermal mode. This refers to the fact that, although the process is preferably adiabatic, in fact, heat exchange occurs during it, since the oxidation reaction is exothermic and the cracking reaction is endothermic, and the heat generated by the exothermic reaction is used in the endothermic reaction.

In all cases, it is believed that the process of the invention favors reactions (4) and (5) rather than reactions (1) and (2) due to a reasonable choice of catalyst, which leads to a decrease in the conversion of H ₂ S, but, on the other hand, to the production of a much higher amount of H ₂ and a significant reduction in the formation of SO ₂ . As the H ₂ S conversion decreases, the rate of recirculation of the acid gases from the H ₂ S containing gas source (eg, amine regenerator) to the reaction chamber increases as compared to a conventional Claus plant.

The catalytic oxidative cracking process of the invention is designed to lower the temperature so as to maintain the desired reaction equilibrium. As a result, hydrogen yield is increased and SO2 generation is minimized, which in turn helps to minimize hydrogen consumption in the tail gas treatment section to reduce SO2 to H2S.

Preferably, the H2S-containing acid gas and the oxygen-containing gas are separately fed into the reaction zone, and these gases are mixed before being fed into the catalytic bed.

The gaseous product from the reaction chamber is preferably rapidly cooled to avoid recombination of H ₂ and S ₂ with the formation of H ₂ S, namely: the reverse reaction (4) occurs, due to which

- 7 040440 the process becomes sub-optimal in terms of overall conversion. This quenching is preferably carried out substantially instantaneously. The quench is preferably carried out to a temperature below 950°C, preferably in the range of 600-850°C. The residence time in the quench zone should preferably be as short as possible, typically 10 to 300 ms, preferably 10 to 100 ms, more preferably 10 to 50 ms.

The quench zone (which may be one of the reaction chamber zones) is preferably followed by a waste heat boiler and a sulfur condenser to cool the process gas and recover liquid sulfur. The latter operation is preferably carried out by rising high or medium pressure steam in the waste heat boiler and low or medium pressure steam in the sulfur condenser.

In one embodiment, quenching of the gaseous product at the outlet of the reaction chamber occurs by mixing with water at the end of the reaction chamber, and the mixing of gas with water is carried out using a water spray in a suitable mixing chamber immediately below the catalyst bed. In the most preferred embodiment, quenching (ie quenching) is performed in the first part of the dual zone waste heat boiler. In this zone with short pipes, the gas typically reaches a temperature in the range of about 600-700°C, and in the second zone with conventional pipes it reaches a temperature of 300-350°C.

Although the formation of SO2 is substantially reduced by the process of the invention, some SO2 will inevitably be formed in any case. To remove such SO2 in the hydrogen production process and plant of the invention, a tail gas treatment section (TGT) is preferably located after the catalytic oxidative cracking step. The TGT section contains, for example, an absorber. Section TGT contains, for example, the hydrogenation reactor, for example, located before (relative to the gas flow) absorber. For example, a portion (eg, about 10-15%) of the hydrogen produced is used to reduce residual SO2 to H2S in the hydrogenation reactor. Due to the much higher hydrogen content and much lower SO2 content compared to the conventional Claus plant, the reduction step in the tail gas treatment section can be carried out without any hydrogen transfer.

The tail gas is preferably preheated and fed into the hydrogenation reactor. In this case, SO ₂ as well as other residual sulfur compounds such as COS and CS ₂ are converted to H ₂ S, which is then removed. Such removal of H2S from the tail gas after the hydrogenation reactor can be carried out in a conventional manner, for example by scrubbing the gas with a lean amine solution in a absorber such as a TGT absorber.

In one embodiment, the catalytic oxidative cracking step in the hydrogen production process and plant of the invention is followed by a single catalytic Claus process step using a gas reheater, a catalytic Claus reactor, and a sulfur condenser (in that order in the direction of gas flow) to convert most of the SO2 to sulfur and thereby minimizing H2 consumption for SO2 recovery in the tail gas treatment (TGT) section. In the plant, the catalytic stage of the Claus process, for example, is carried out after (relative to gas flow) the COC zone and before (relative to gas flow) the TGT section.

In one embodiment, the hydrogen stream obtained from the TGT sink is sent to the frontier of the plant (eg, end users), such as a plant selected from the group consisting of hydrotreaters, hydrocrackers, or hydrodesulfurization units. It should be noted that the composition of the hydrogen-rich stream from the top of the TGT absorber may differ depending on variables such as SRU feedstock quality (e.g. composition of the H2S containing gas stream fed to the COC reaction zone), plant configuration and operating conditions, and may contain trace amounts or percentages of H2O, N2, CO, CO2, H2S, COS and CS2.

In a preferred embodiment, the hydrogen stream coming from the TGT absorber (eg, the gas stream from the top of the TGT absorber) is further purified in a hydrogen purification section (eg, a pressure swing absorption unit). It should be noted that the composition of the hydrogen-rich stream from the top of the TGT absorber prior to purification may differ depending on variables such as SRU feedstock quality (e.g. composition of the H ₂ S containing gas stream fed to the COC reaction), plant configuration and operating conditions. , and may contain residual amounts or percentages of H ₂ O, N ₂ , CO, CO ₂ , H ₂ S, COS and CS ₂ .

The purified hydrogen is sent to the boundary of the plant or, for example, to one or more units selected from the group consisting of hydrotreating units, hydrocracking units or hydrodesulfurization units.

The hydrogen production process optionally includes the step of providing a gas stream containing H2S. The gas stream is obtained, for example, from an amine recovery system and/or from an acid water stripper. The method optionally includes the step of regenerating the amine-based absorbent in the amine regeneration system, for example by heating, thereby providing the regenerated absorbent and a gas stream containing H ₂ S, the gas stream containing H ₂ S being fed into the COC reaction.

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The plant of the invention is, for example, a sulfur recovery unit (SRU), for example, having an inlet for a gas stream containing H2S connected to an inlet of amine recovery systems and/or an acid water stripper. The plant of the invention optionally comprises amine recovery systems and/or an acid water stripper having an outlet for a gas stream containing H ₂ S connected to an inlet of a catalytic oxidative cracking reaction zone. In some embodiments, the plant comprises a COC plant comprising a catalyst bed containing a COC catalyst (having the catalyst composition described in the hydrogen production process) and a quench unit located downstream of the reaction chamber, and preferably a static mixer located upstream of the COC reaction chamber. The static mixer has, for example, one or more inlets for a gas stream containing H2S and one or more inlets for an oxygen-containing gas.

The drawing shows schematically an example of an installation according to the invention.

The COC stage (3) or COC section (3) includes a COC unit (4) comprising a COC reaction chamber (42), the chamber containing a catalyst bed with a COC catalyst having the composition described herein. The COC plant further comprises a waste heat boiler (6) for cooling, such as quenching, the reaction product and rising steam; the boiler (6) is, for example, a shell and tube heat exchanger. The COC unit (4) further comprises a preferred static mixer (5) located upstream of the COC chamber for mixing acid gas with oxygen-containing gas (41), the acid gas being, for example, acid gas (1) from one or more amine recovery units. and/or acid gas (2) from one or more acid water strippers. Stage (3) COC additionally contains a (first) sulfur condenser (7), for example, a heat exchanger having a sulfur outlet (8) and a gas outlet (14), which is in fluid communication with the inlet of the TGT section, optionally (as shown ) through the optional catalytic stage (12) of the Claus process, in particular with the reheater (9) of this stage. The reheated gas is fed into the first Claus catalytic reactor (10), the outlet of which is connected to the second sulfur condenser (11), which also has a sulfur outlet (8) and a gas outlet (15), which is fed to the TGT section (13).

As illustrated, the preferred TGT section contains an absorber (25) whose inlet is in fluid communication with said gas outlet (14) of the COC section (3) to remove H ₂ S from the gas stream from the COC section or from the optional Claus section. The TGT section further preferably comprises a regenerator (27) for regenerating the rich absorbent liquid from the absorber. More preferably, the TGT section (13) comprises an optional TGT preheater (16) receiving gas from an optional second sulfur condenser (11) or from a first sulfur condenser (7) and has an outlet connected to a preferred hydrogenation reactor (18), which optionally contains a hydrogenation catalyst and optionally has no input for external H2 feed. The hydrogenation reactor has a gas outlet connected to an optional TGT waste heat recovery boiler (19) for cooling the gas by heat exchange, for example with boiler make-up water, which has an outlet connected to the cooling tower (20). In the cooling tower (20), the gas is cooled by circulating condensate (21). The cooling tower has a gas outlet at the top which is connected to an absorber inlet (25) in which, for example, a depleted amine solution (26) is circulated in the gas stream to absorb the remaining H ₂ S. The absorber (25) has a gas outlet (34) in the upper part, and the gas stream (34) is enriched in hydrogen. The gas stream is optionally further purified in a purification unit (43), such as a pressure swing adsorption unit, to obtain a purified commercial hydrogen stream (44). The absorber (25) has an outlet (28) for the absorbent liquid in the lower part, connected via a pump (29) and a heat exchanger (30) to an amine regeneration unit (27), in which the solution is regenerated, for example, by heating. The amine recovery plant (27) is, for example, a column capable of stripping the amine solution and has, for example, a gas outlet (33) at the top, in which gas is fed to, for example, the COC stage (3), and an outlet for regenerated or depleted amine absorbent supplied to the absorber (25), for example, through a pump (31), through a heat exchanger (30) for heat exchange with a saturated amine solution and through a cooler (32). The cooling tower (20) additionally has an outlet at the bottom for liquid, for example cooling water, which is partially recirculated, preferably through a pump (22) and cooler (23), and which is partially (24) blown and directed to the installation boundary.

The regenerator (27), for example, is used separately and in addition to the amine regeneration unit, which is the source of acid gas (1), or, for example, the regenerator (27) processes absorbent from additional absorbents other than absorber (25), or, for example, absorber (25) is used not only for the gas flow from the quench cooler (20), but also for any main source of acid gas.

The common source or external source of acid gas in the hydrogen production process and plant of the invention is, for example, a process in which acid gas is generated, such as a natural gas treatment process or a crude oil refining process, or is, for example

- 9 040440 measures, an installation in which acid gas is produced, such as an oil refinery, a natural gas processing plant, a gasification plant or a gas synthesis plant. The process or plant that generates acid gas includes, for example, removing sulfur-containing components from the stream, such as natural gas and flue gas, for example, using absorption and desorption, and, for example, stripping is carried out using an amine recovery unit providing acid gas (1).

The hydrogen production process is preferably carried out in the plant described herein. Preferences with respect to the catalyst composition and the catalyst production method apply equally to the catalyst contained in the plant catalyst bed and used in the hydrogen production process.

Examples

The invention will now be illustrated with reference to the following non-limiting examples.

Example 1. Obtaining a catalyst.

NiS-based catalysts with a nominal sulfide load in the range from 9 to 25 wt.% were obtained by the method of impregnation according to the moisture capacity of the support containing aluminum oxide and cerium dioxide.

The procedure for preparing cerium oxide supported on alumina used an aqueous solution of a suitable precursor salt in the form of cerium nitrate and Al ₂ O ₃ in powder form. The Al ₂ O ₃ support was pretreated at 600°C (10°C/min) for 2 h.

To prepare various samples, an aqueous solution of a suitable cerium salt at an appropriate concentration was used to impregnate 10 g of support. This solution was placed on a plate for half an hour with stirring and heating to 100°C. After that, the samples were dried and calcined in air at a temperature of 1100°C for a period of time ranging from 15 to 30 hours.

For nickel impregnation of the cerium/alumina support, an aqueous solution of nickel nitrate hexahydrate with a molar concentration in the range of 1-3 M was used. In addition, in this case, the solution was applied to the plate for half an hour while heating to 100°C. After that, the samples were dried and calcined in air at a temperature of 1100°C for 15 h. After calcination, the catalysts were reduced to the desired granulometry and sulfided. For such a procedure, a gas stream enriched with H2S at a concentration of 5 to 50 vol.% must be fed into the reactor during heating with a linear temperature change from 2 to 40°C/min to a temperature of 800-1200°C for 1-8 hours .

In particular, the sulfiding treatment was carried out for laboratory studies in a quartz reactor containing a catalyst to be sulfonated. In particular, the activation step was carried out by supplying a gaseous stream containing N2 and H2S at a concentration of 20% by volume, raising the temperature to 1000°C at a heating rate of 20°C/min for 2 hours under isothermal conditions.

Example 2. Laboratory study of the catalyst.

The laboratory process involves the reaction of H2S with oxygen in a substoichiometric ratio with the simultaneous production of elemental sulfur and hydrogen. Some of the H2S reacts with oxygen in an exothermic reaction that releases heat from the unreacted H ₂ S for a violent endothermic reaction.

In particular, for the reaction test, a sulfonated catalyst weighing ~ 3 g is loaded into the reactor.

The reactor is heated from ambient temperature to 1000°C at a rate of 20°/min in a stream containing only nitrogen. After reaching this temperature, the reactor is heated to the reaction temperature (1050-1060°C) in the presence of a feed stream containing: H2S: 65% O ₂ : 17%, NH ₃ : 6% and N2, as well as trace amounts of methane and CO2. The contact time varies between 0.5 and 2.2 s.

After testing the catalyst, the reactor is cooled with a stream containing H2S (20 vol.%) from the reaction temperature to 500°C, and then cooled with a nitrogen stream to ambient temperature.

The tested catalyst compositions and the results obtained are shown in the table below.

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Claims (11)

catalytic compositions; reaction conditions; laboratory test results No. A Composition 17% NiS/ 5% CeO ₂ /A1 ₂ Oz Temp. (°C) 1050 H ₂ S conversion (%) 41 Yield H ₂ (%) 4 Selectivity SO ₂ (%) 1.2 Conc. SO ₂ outlet (ppm) 1700 Conc. NH ₃ outlet (ppm) <50 B C 17% NiS/10% CeO ₂ /A1 ₂ Oz 17% NiS/ 25% CeO ₂ /A1 ₂ Oz 1020 1050 41 37.5 3 1.7 0.25 0.05 750 120 <50 <50 D 9% NiS/ 15% CeO ₂ /A1 ₂ Oz 1050 37.2 3.4 0 0 <75 E 17% NiS/15% CeO ₂ /A1 ₂ Oz 1070 44 7 0 0 <50 F 25% NiS/15% CeO ₂ /A1 ₂ Oz 1050 39 3 0 0 <75 In some of these laboratory experiments, a powder containing a catalyst is loaded into a reactor and a layer of ceramic fiber is placed behind the catalyst bed. Even after a long time at high temperatures, the ceramic fiber layer did not change color. This indicates that the metal within the catalyst composition generally has little ability to go into the gas phase at high temperature, since the discoloration of the ceramic fiber layer is usually attributed to metal migration from the catalyst surface into the gas phase. Conversely, in similar tests using a different catalyst (containing iron sulfide and molybdenum supported on a support containing aluminum), the color of the ceramic fiber layer changed from white to black. At the same time, the color of the catalyst powder changed from black to grey. These phenomena can be explained if we take into account the low stability of the active phases on the carrier of aluminum oxide at high temperatures and under more unfavorable end-of-cycle conditions, due to which metal sulfides migrate into the ceramic fiber layer. As mentioned above, this phenomenon was not observed in the case of the catalysts of the invention. In one aspect, this demonstrated a higher level of stabilization of ceria-alumina promoted nickel sulfide than other catalysts. Thus, these experiments show a higher thermal stability of the catalyst in accordance with the compositions of the invention. CLAIM 1. A method for producing hydrogen from a gas stream containing H2S, including catalytic oxidative cracking of the gas stream with the formation of H2 and S2, the use of a catalyst containing at least one active metal selected from the group consisting of iron, cobalt, nickel, and combinations thereof, moreover, the specified active metal is deposited on a carrier containing cerium dioxide and alumina. 2. The method of claim 1 wherein the catalyst composition contains nickel. 3. Process according to claim 1 or 2, wherein the catalytic oxidative cracking is carried out with a molar ratio of H2S/O2 in the feedstock greater than 2:1, preferably in the range of 2:1 to 6:1. 4. The method according to claim 3, in which the ratio is in the range from 3:1 to 5:1, preferably from 3.5:1 to 4.5:1. 5. A process according to any one of the preceding claims, wherein the catalytic oxidative cracking is carried out using an oxygen-containing gas stream containing at least 40 vol% oxygen, preferably at least 60 vol% oxygen. 6. The method of claim 5, wherein the oxygen-containing gas stream is oxygen having a purity of 90-100% by volume. 7. Process according to any one of the preceding claims, wherein the catalytic oxidative cracking is carried out at a temperature in the range of 700 to 1300°C, preferably in the range of 950 to 1250°C. 8. A process according to any one of the preceding claims, comprising the additional step of hydrogenating the formed SO2 and optionally also COS and/or CS2. 9. A process according to any one of the preceding claims, wherein the catalytic oxidative cracking gas is quenched substantially without delay to a temperature below 950°C. 10. A process according to any one of the preceding claims, wherein the H2S-containing acid gas and the oxygen-containing gas are mixed in a static mixer prior to being fed into the first catalyst bed in the catalytic oxidative cracking reaction zone or zones. 11. Process according to any one of the preceding claims, wherein the molar ratio of NH3/O2 in the feedstock is between 0.1 and 1.5. -