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Tail gas treatment of SEWGS technology

literature review on CO2 and H2S separation

FABBRI Eric    

Project supervised by: Dr H.A.J. Van Dijk

    

ECN-E--11-067 JUNE-SEPTEMBER 2011

 

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Acknowledgement/Preface  I would like to express my gratitude to my supervisor Dr H.A.J. van Dijk (Eric) for giving me the opportunity to do my training at the ECN, for the good advices, the interesting discussions and for following with interest the progression of my work. I would like to thank also Pr Vasile Hulea, Professor in catalysis at the ENSCM who gives me precious help and advices for having this internship, and more, a real passion for catalyze chem-istry. Thank you Jan Wilco Dijkstra for your help, advice and every information he gave me in proc-ess and thermodynamic chemistry. Thank you Stephane Walspurger, French compatriot, for your help and advice during three months. With a special thanks to permit me to pass my English test at the ECN. Thank to the whole team for accepting me in the good work atmosphere of BKM unit. And, of course, many thanks to all the other interns and Den Helder Richter art gallery occu-pants for their friendship and good moments shared that I will never forget.   

Abstract  This literature review is the result of an investigation of the most important way to remove sulphur for the last decades. We will discuss about Claus and Claus tail gas process possibility to solve our problem. But also solution which come from membranes, direct oxidation catalyze, or more, from acid gas removal technology, sorbent technology, liquid oxidation. Each field will be describe and explain to understand in which way it could be suitable to separate CO2 and H2S and reach our goals in matter of CO2 transport and storage condition (purest as possible). Finally the target of this work will be to propos some interesting and promising solutions in view to define research ways for future experiment and investigations.

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Contents  

List of tables 5 

List of figures 6 

List of symbols, abreviations 8 

1.  Presentation of the Institute 11 1.1  History 11 1.2  Policy and mission 11 1.3  Organization 12 1.4  Department Biomass, Coal & Environment (BKM) 13 

2.  Introduction 15 2.1  Background IGCC 17 

2.1.1  IGCC 17 2.1.2  WGS + AGR 18 2.1.3  SEWGS 19 

2.2  Project description and objectives 21 2.3  Approach of the task 22 

3.  Comparative study on three IGCC plants 23 

4.  AGR technology 27 4.1  COS Hydrolysis 27 4.2  H2S removal via solvents 27 4.3  Chemical Solvents 28 4.4  Physical Solvents 29 

4.4.1  General 29 4.4.2  The Selexol process 30 

4.5  Hybrid Solvents 31 4.6  Balance 33 

5.  Claus process 35 5.1  Generality 35 5.2  Claus Plant Configurations 37 

5.2.1  Straight-Through Configuration. 37 5.2.2  Split-Flow Configuration 38 

5.3  The Superclaus Process, the Euroclaus Process 39 5.3.1  Superclaus-99 39 5.3.2  Superclaus-99.5. 40 5.3.3  The Euroclaus process 41 

5.4  Sub-Dewpoint Processes 42 5.4.1  The Sulfreen, Hydrosulfreen, and Doxosulfreen Processes 43 

5.4.1.1  Sulfreen 43 

5.4.1.2  Hydrosulfreen 44 

5.4.1.3  Doxosulfreen 45 5.4.2  The CBA/MCRC Process 47 

5.4.2.1  The CBA Process 47 

5.4.2.2  The MCRC Process 48 5.5  SCOT / Super-SCOT processes 49 5.6  The Selectox process 50 5.7  The MODOP process 51 5.8  Balance 52 

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6.  Membranes 57 6.1  Approach 57 6.2  H2S/CO2 Separation 58 

6.2.1  Theory 58 6.2.2  Experience 59 

6.3  Balance 60 

7.  Catalytic oxidation to elemental sulphur 61 7.1  Iron catalyst 61 

7.1.1  Fe2O3 active phase supported on SiC 61 7.1.2  Iron oxide catalyst with Ce incorporated 63 

7.2  Nickel catalyst 64 7.3  Balance 67 

8.  Others 69 8.1  Liquid oxidation 69 

8.1.1  The LOCAT Process 69 8.1.2  The Stretford process 71 

8.2  WSA 72 8.3  Balance 74 

9.  RTI Canada prototype (HTDT) 76 

10.  Sorbent catalogue 78 10.1  Sud Cheme 78 10.2  Axens 79 

11.  Conclusion 80 

12.  References 83  

ECN-E--11-067

List of tables

Table 2.1  Comparing SEWGS feed and CO2 product species  22 Table 2.2  Comparing CO2 product and CO2 for transport and sequestration  22 Table 3.1  IGCC systems studied specifications [2]  25 Table 3.2  IGCC systems efficiency cost capture and emissions [2]  25 Table 3.3  Pressure and temperature at each stages of IGCC processes [2]  26 Table 4.1  Common Chemical Reagents Used in AGR Processes [2]  29 Table 4.2  Common Physical Solvents Used in AGR Processes [2]  30 Table 4.3  Common Mixed Solvents Used in AGR Processes [2]  32 Table 4.4  Selexol/Purisol process balance  34 Table 5.1  Superclaus 99-99.5/Euroclaus process balance  52 Table 5.2  Sub-dewpoint processes balance  53 Table 5.3  Sulfreen process cost  53 Table 5.4  SCOT process balance  55 Table 5.5  Selectox process balance  55 Table 5.6  MODOP process balance  56 Table 7.1  Fe2O3 on SiC balance  67 Table 7.2  Fe-Ce (2/2) balance  67 Table 7.3  NiS2 on SiC balance  68 Table 8.1  VK series catalyst characteristic [23]  72 Table 8.2  WSA process energetic data [24]  73 Table 8.3  LOCAT/Stretford process balance  74 Table 8.4  WSA process balance  75 Table 9.1  Data RTI desulphurization process[26]  77 

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List of figures

Figure 2.1  Schematic overview of an IGCC  18 Figure 2.2  Schematic WGS + AGR treatment  18 Figure 2.3  Schematic SEWGS treatment  19 Figure 2.4  Power production with CO2 capture by SEWGS. Power is generated by gas

turbines (GT) and steam turbines (ST), heat is recovered for steam generation (HRSG). Steam is needed for auto thermal reforming and water gas shift (ATR/WGS), and for CO2 removal by SEWGS.  20 

Figure 2.5  SEWGS principle. Upper reactor: adsorption and reaction at high pressure, lower reactor: desorption at low pressure  20 

Figure 2.6  Structure of Hydrotalcite compound  21 Figure 2.7  Place of project in SEWGS process  21 Figure 4.1  Flow Diagram for a Conventional AGR Unit  28 Figure 4.2  H2S and CO2 removal with a Selexol process  31 Figure 4.3  Equilibrium Solubility Data on H2S and CO2 in Various Solvents  33 Figure 5.1  H2S conversion and sulphur composition according to temperature [8]  36 Figure 5.2  Claus plant schemas Straight-Through Configuration [8]  38 Figure 5.3  Claus plant schemas Split-Flow Configuration Source[8]  38 Figure 5.4  Superclaus-99 process schema [8]  40 Figure 5.5  Superclaus-99 process schema balance  40 Figure 5.6  Superclaus-99.5 process schema  41 Figure 5.7  Superclaus-99.5 process schema balance  41 Figure 5.8  Euroclaus process schema [7] [8]  42 Figure 5.9  Euroclaus process schema balance  42 Figure 5.10  Sulfreen process schema [7] [8]  44 Figure 5.11  Hydrosulfreen process schema [7] [8]  45 Figure 5.12  Hydrosulfreen process schema balance  45 Figure 5.13  Doxosulfreen process schema [7] [8]  46 Figure 5.14  Doxosulfreen process schema balance  46 Figure 5.15  Four-Reactor CBA process schema [7] [8]  47 Figure 5.16  Three-Reactor CBA process schema [7] [8]  48 Figure 5.17  MCRC process schema [11]  49 Figure 5.18  Scot process schema balance  49 Figure 5.19  Scot process schema  50 Figure 5.20  MODOP process schema balance  52 Figure 5.21  Sulphur phase diagram equilibrium  54 Figure 6.1  Membrane reactor configuration for CO2 AND H2 separation with H2S removal  57 Figure 6.2  Hydrogen production from gasification system with membrane reactor  57 Figure 6.3  H2S comportment on mixed ionic/electronic conducting membrane  59 Figure 6.4  H2S, CO2 and He permeation on dual phase membrane 850°C 1.03 bar [14]  60 Figure 7.1  Fe2O3/SiC; Selectivity, conversion at different operating conditions [15]  62 Figure 7.2  Selectivity conversion according to different active phase on SiC 240°C [15]  63 Figure 7.3  Selectivity conversion according to different iron oxide and Ce incorporated

catalyst [17]  64 Figure 7.4  H2S conversion, S-selectivity and sulphur deposition as a function of time on a

NiS2/SiC catalyst at 100°C [18]  65 Figure 7.5  Sulphur deposal schemas in a catalyst pore (SiC) [19]  66 Figure 8.1  LO-CAT process schema [21]  70 Figure 8.2  WSA process schema  72 

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Figure 9.1  RTI Canadian institute reactor [26]  76   

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List of symbols, abreviations

 ACI American Combustion, Inc. AGE acid gas enrichment AGI acid gas injection AGR acid gas removal bar SI unit of pressure (1 bar = 14.5 psi) BFW boiler feedwater BSR Beavon Stretford Reactor Btu British thermal units (1,055 joules) °C temperature, degrees Centigrade (Celsius) CCT Clean Coal Technology (DOE) CFB circulating fluidized-bed CGCU cold gas cleanup CO carbon monoxide CO2 carbon dioxide COPE Claus Oxygen-based Process Expansion COS carbonyl sulfide DEA diethanolamine DGA diglycolamine DIPA diisopropanol amine DSRP Direct Sulphur Recovery Process °F degrees Fahrenheit (°F = 1.8 x °C + 32) g gram GT gas turbine GTI Gas Technology Institute HAPS Hazardous Air Pollutants HCN hydrogen cyanide HGCU hot gas cleanup HP high pressure (steam) HRSG heat recovery steam generator HRU heat recovery unit HSS heat stable salts IGCC integrated gasification combined cycle lb(s) pound(s) (454 grams) LLB Lurgi Lentjes Babcock LNG liquefied natural gas LP low pressure (steam) lt long ton (2,240 pounds, 1,016 kg) LTS low temperature separation MDEA methyldiethanolamine MEA monoethanolamine MMEA methyl monoethanolamine mt metric ton (2,204.6 pounds, 1,000 kg) NSPS New Source Performance Standards Oxyburner oxygen burner OxyClaus oxygen-blown Claus PCD particulates control device ppm parts per million ppmv parts per million volume ppmw parts per million by weight

ECN-E--11-067

PSDF Power Systems Development Facility psi(a) pounds per square inch (absolute) scf standard cubic feet (1 atm, 60°F) SCR selective catalytic reduction SR sulphur recovery SRU sulphur recovery unit st short ton (2,000 pounds, 907 kilograms) SWS sour water stripper t ton TGT tail gas treating UOP Universal Oil Products WHB waste heat boiler

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ECN-E--11-067

1. Presentation of the Institute

1.1 History The Energy Research Centre of the Netherlands (ECN) originated in 1955 as Reactor Centre of the Netherlands (RCN). RCN started its research on the development of nuclear energy to be used for peaceful applications. At the time, nuclear energy was a promising new way to obtain clean and inexpensive energy for generations to come. This was in contrast to exploited coal that was considered to be expen-sive, dirty, laborious, and dangerous. During the first years, RCN collaborated intensively with the Norwegians at the research reactor in Kjeller, Norway. Then the Netherlands decided to build their own reactor, The High Flux Re-actor (HFR), which is located far from high population concentrations and close to cooling-water in the dunes near the village of Petten, where the first nuclear fission reaction was carried out in 1962. The HFR was built as a research reactor to study the effects of radiation and was not intended to be used as a nuclear power plant, which means that no electricity would be produced. Later on also specialty materials were being made for medical applications. In 1975, resistance rose against the use of nuclear energy and more research into other fields of energy generation needed to be done, including wind energy, solar energy, tidal energy, terres-trial heat and biomass. RCN was appointed to be the institute to largely implement this research and consequently, in 1976, the name was changed to Energy Research Centre of the Nether-lands. From then on, more research was conducted towards non-nuclear energy sources and in 1988 the nuclear research department of ECN was split off into a subsidiary. Together with KEMA (Partnership of the Netherlands towards the testing of Electrical Materials) the Nuclear Research and Consultancy Group (NRG) was started up.   

1.2 Policy and mission ECN aims, with its energy research, at a durable, safe, efficient, reliable, and environmentally friendly energy supply. This is achieved in cooperation with the government, industry, and uni-versities, both in the Netherlands and abroad. ECN performs a lot of research for the knowledge and information facilities of the government, necessary for the preparation and evaluation of policies and to construct policy objectives in areas concerning energy, environment and techno-logical innovation. As an associate with industry, ECN develops and implements new products, processes and technologies, which are of importance for the transition towards durable energy housekeeping. ECN closely works together with universities and research institutes all over the world and performs a bridging function towards implementation by carrying out technological research The research policy at ECN, the so-called Trias Energetica, is divided into three parts, which holds the efficient use of energy, generation and deployment of renewable energy sources, and clean conversion of fossil fuels. This will result in minimizing the demand for energy produc-tion, facilitating the transition between a post-oil economy and using fossil fuels without harm-ing the environment. Thus, this policy contains all the elements required for the transition to-

ECN-E--11-067 11

wards a sustainable energy management and allows ECN to operate at the cutting edge of Euro-pean efforts in the development of new technologies that meet tomorrow’s energy demand.

1.3 Organization ECN employs around 600 people and another 300 are working at NRG.  

  Research at ECN is organized into nine departments : Energy Efficiency in Industry ; Policy Studies ; Engineering and Services ; Hydrogen and Clean Fossil Fuels ; Solar Energy ; Wind Energy ; Biomass, Coal, and Environmental Research ; Energy in the Built Environment ; and Intelligent Energy Grids ; as depicted in the lower part of the organization chart. ECN has its own department for Quality, Safety, and Environment (KVM) that together with the Board of Directors and four other supporting departments comprises the Supervisory Board. This department aims at regulatory, supporting, and policy preparatory tasks concerning the fields of quality, safety, and environment. In this regard, ECN has been certified in accordance with the ISO 9001 quality standard. Furthermore, ECN also meets the ISO 14001 environmental standard, which states that the research performed at ECN limits the influence on the environ-ment. The granted Pollution of Surface Waters act is directly linked to answering the ISO 14001 standard. In cooperation with the Achmea insurance company, ECN looks for the optimum working cir-cumstances for all employees. Moreover, all accidents, incidents, and near misses are reported and registered to allow measures to be taken and to improve safety. To maintain these high standards, ECN works with a management system. On the basis of this system, every employee can work safely, healthy, and environmentally friendly.

ECN-E--11-067

  

1.4 Department Biomass, Coal & Environment (BKM) The BKM department is split into two different parts:

• Biomass and Coal contribute to a cleaner, less wasteful and more sustainable use of these two energy resources.

• Environment assesses the impact of human activities on air and soil quality and provides policy support.

ECN-E--11-067 13

  This department is structured into six research groups – Heat and Power, CO2 capture technol-ogy, Air Quality and Climate Change, Transportation fuels and Chemicals, Environmental Risk Assessment, Syngas and SNG, as shown above. The group CO2 capture technology develops new processes and new technology in order to re-duce CO2 emissions. It is essential in today’s society especially for producing chemicals, fuels and of course energy. The CO2 capture is an essential and compulsory step in creation of clean sustainable energy.

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2. Introduction

With the fast development of countries like India and China, CO2 emissions become a priority to handle. The increasing amount of CO2 released in the atmosphere due to human activity is said to have an impact on global climate through the so called greenhouse effect. Basically, each process that includes transformation or conversion of fossil fuels produces CO2. Nowadays, this CO2 is mainly vented in the atmosphere without further treatment, but legisla-tions around the world may be modified in the upcoming years. A tax about CO2 emissions may be applied to the countries that have the highest emissions, related to a certain amount of CO2 they are allowed to produce. With the instability related to oil supply and the increase of its price, many countries turn to coal as a major primary energy resource. Known reserves of coal exceed a century, which makes it a good candidate for emerging countries energy supply. As fossil fuels fired power plants account for more than a third of the total CO2 worldwide emissions, they are obvious targets for imple-mentation of CO2 capture technologies. Technologies to separate CO2 are widely used, but have not been tested in power plants on full industrial scale for capture yet. The combination of empty fossil fuel fields, saline aquifers pro-vides enough storage space to have a significant impact on emissions. The reliability of those sites has been tested, like in the Sleipner gas field in the North Sea, for their capacity to store CO2 safely and indefinitely. The state of the art on CO2 capture and storage shows that methods are approved on a small scale. The challenge is now to extend it to the full scale and reduce the cost of their use, to make them competitive. Actually, in Carbon Capture, transport and Storage (CCS), the largest cost come from the capture part where research and development still need to be done. Target industries for CCS technology are high-purity CO2 sources for CO2 industrial applica-tions, cement, iron and steel, refineries and biofuel conversion industries. CCS is a cost-effective solution for CO2 capture thanks to relatively low cost due to high purity and high con-centration of CO2 in the stream. And fact that additional costs are limited to equipment and compressor power requirements. In addition transport and storage are relatively low cost for these plants because they are normally in the vicinity of industrial complexes or coastal loca-tions. The removal of CO2 in power plants can be done in several ways, at different steps of the proc-ess. One option is to use end-of-pipe technologies consisting of removing CO2 from the flue gas of the power plant using a scrubber, e.g. solvent or amine solution. This solution has a large poten-tial in the industry because it is considered the easiest and cheapest way to implement carbon capture unit in existing pulverized coal fired power plant. Nowadays, nearly all coal power plants in the world are indeed using heat provided by coal combustion to generate electricity. In this plant design, post combustion CO2 capture technology can be implemented in combination with other technologies for NOx or SOx gases. However the overall efficiency is significantly decreased by several percent-points compared to a plant without CO2 capture. Decreasing the absolute efficiency penalty involved by CO2 capture is the main topic in the development of new solvents and scrubbing systems. Alternatively, solid sorbents are considered for post com-bustion CO2 capture.

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Another solution, known as oxyfuel combustion, is the use of nitrogen-free, pure oxygen gas for combustion of fossil fuels. On the one hand, nitrogen-free flue gas allows extracting CO2 from a concentrated gas stream. On the other hand, oxygen separation still remains a costly process. It has been estimated that the oxyfuel capture would result in an efficiency penalty of 9 to 10% points. Finally a promising route to reduce greenhouse gas emissions during power production is pre-combustion decarbonisation. Here, the CO2 is captured prior to combustion of the fuel by transferring the energy content of the fuel to hydrogen.  

  In pre-combustion capture, coal is first gasified to yield syngas, which is further converted to H2 and CO2 via the water gas shift reactors.

CO + H2O ↔ CO2 + H2 Water-gas shift Here again, solvent scrubbing system as well as membrane reactors may be interesting options for capture. Alternatively, there is investigations on the Sorption-Enhanced Reaction Process (SERP), in which CO2 is adsorbed by an adsorbent during the catalytic production of hydrogen. The sorption-enhanced water-gas shift (SEWGS) process, a so-called 2nd generation cap-ture technology, has been developed as a solution which integrates two processes. CO2 is ad-sorbed during the WGS reaction, thanks to a sorbent mixed with the catalyst in the reactor. Both materials are placed in a series of reactors as pellets. Since the adsorption and catalytic step are performed at high pressure (typically 30 bars) and the desorption is carried out at low pressure, a pressure swing adsorption unit needs to be used. The continuous production of hydrogen is ensured by multiplying the number of reactors.  

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  In such a process, high pressures facilitate the adsorption of gases on solid sorbents. The SEWGS process results thus in relatively small volume units when compared to post combus-tion techniques. The penalty of the capture on the whole process is between 8 and 10 %points, less than for oxyfuel. The development of Sorption Enhanced Water shift Reaction (SEWGS) for the production of energy from coal requires an efficient post treatment for CO2 and H2S removed in view of stor-age (CO2) and reuse (H2S transformation in elemental sulphur). That is the purpose of this re-port.  

2.1 Background IGCC

2.1.1 IGCC An integrated gasification combined cycle (IGCC) is a technology that turns coal into synthesis gas (syngas) which is then used in a gas-turbine – steam turbine combined cycle to produce electricity. Impurities need to be removed from the coal gas before it is combusted to ensure low emissions of sulphur dioxide, particulates and mercury. Excess heat from the primary combus-tion and generation is then passed to a steam cycle, similarly to a combined cycle gas turbine. This then also results in improved efficiency compared to conventional pulverized coal. The gasification process can produce syngas from coal, heavy petroleum residues and biomass. The process used by IGCC plants can be broken down into five steps:

1. The coal, NG or biomass is gasified to produce a synthetic gas (syngas) 2. Contaminants (sulphur, HCl, HCN, Hg) are removed from the syngas, then electricity is

generated using a combined cycle, consisting of the following three steps: 3. A gas turbine-generator burns the clean syngas 4. Heat from the exhaust and heat from the gas turbine are used to create steam 5. The steam is used to power a steam turbine-generator, also some of the steam can be

used (in AGR unit for example)

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http://en.wikipedia.org/wiki/Coalhttp://en.wikipedia.org/wiki/Syngashttp://en.wikipedia.org/wiki/Steam_cyclehttp://en.wikipedia.org/wiki/Combined_cycle_gas_turbinehttp://en.wikipedia.org/wiki/Syngashttp://en.wikipedia.org/wiki/Biomass

 Figure 2.1 Schematic overview of an IGCC

The plant is called integrated because its syngas is produced in a gasification unit in the plant which has been optimized for the plant's combined cycle. In this example the syngas produced is used as fuel in a gas turbine which produces electrical power. To improve the overall process efficiency heat is recovered from both the gasification process and also the gas turbine exhaust in 'Waste Heat Boilers' producing steam. This steam is then used in steam turbines to produce additional electrical power, see Figure 2.1. Research continues … The concern about climate change and the potential for carbon dioxide sequestration potentially makes IGCC one of the most environmentally responsible technologies available today. The challenge lies in to have higher thermal efficiency and to hold the cost down. At the moment, two major processes are studied to remove CO2 and H2S.  

2.1.2 WGS + AGR

 Figure 2.2 Schematic WGS + AGR treatment

The WGS reaction is an important industrial reaction. It is often used in conjunction with steam reforming of methane or other hydrocarbons, which is important for the production of high pu-

ECN-E--11-067

http://en.wikipedia.org/wiki/Gasificationhttp://en.wikipedia.org/wiki/Steam_reforminghttp://en.wikipedia.org/wiki/Steam_reforminghttp://en.wikipedia.org/wiki/Hydrocarbon

rity hydrogen. The water-gas shift reaction was discovered by Italian physicist Felice Fontana in 1780. Selection of Technology - In the cases with CO2 separation and capture, the gasifier product must be converted to hydrogen-rich syngas. The first step is to convert most of the syngas car-bon monoxide (CO) to hydrogen and CO2 by reacting the CO with water over a bed of catalyst. The H2O:CO molar ratio in the shift reaction, shown below, is adjusted to approximately 2:1 by the addition of steam to the syngas stream thus promoting a high conversion of CO. In the cases without CO2 separation and capture, CO shift convertors are not required.

Water Gas Shift: CO + H2O ↔ CO2 + H2 The reaction is slightly exothermic, yielding 41.1 kJ (10 kcal) per mole

The CO shift converter can be located either upstream of the acid gas removal step (sour gas shift) or immediately downstream (sweet gas shift). If the CO converter is located downstream of the acid gas removal, then the metallurgy of the unit is less stringent but additional equipment must be added to the process. Products from the gasifier are humidified with steam or a wet quanch. If the CO converter is located downstream of the acid gas removal, then the gasifier product would first have to be cooled and the free water separated and treated. Then additional steam would have to be generated and re-injected into the CO converter feed to meet the re-quired water-to-gas ratio. If the CO converter is located upstream of the acid gas removal step, no additional equipment is required. This is because the CO converter promotes carbonyl sulfide (COS) hydrolysis without a separate catalyst bed.

2.1.3 SEWGS

 Figure 2.3 Schematic SEWGS treatment

The Sorption Enhanced Water Gas Shift process (SEWGS) is a technology that combines the water-gas shift reaction with CO2 capture at high temperature in a single unit [1], see Figure 2.3. The feed to the SEWGS unit is syngas, which is produced by reforming and high temperature shift. The products are a CO2 and H2S stream at low pressure and high temperature a H2-rich stream at high pressure and high temperature. The CO2 can be compressed and transported to a suitable storage location, e.g. geological formations thanks to a post treatment. In a reactor CO is converted with H2O to CO2 and H2 (water-gas shift reaction). Simultaneously, CO2 is re-moved from the gas by sorption on a promoted hydrotalcite-like material. The sorbent is re-generated by purging counter currently with steam at low pressure. Since adsorption of CO2 is thermodynamically favoured at high pressure and desorption at low pressure, parallel reactors are operated in pressure swing cycles.

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http://en.wikipedia.org/wiki/Hydrogenhttp://en.wikipedia.org/wiki/Felice_Fontanahttp://en.wikipedia.org/wiki/Exothermichttp://en.wikipedia.org/wiki/Joule#Kilojoulehttp://en.wikipedia.org/wiki/Caloriehttp://en.wikipedia.org/wiki/Mole_(unit)

Figure 2.4 Power production with CO2 capture by SEWGS. Power is generated by gas tur-bines (GT) and steam turbines (ST), heat is recovered for steam generation (HRSG). Steam is needed for auto thermal reforming and water gas shift (ATR/WGS), and for CO2 removal by SEWGS.

 

 

 Figure 2.5 SEWGS principle. Upper reactor: adsorption and reaction at high pressure,

lower reactor: desorption at low pressure

Sorbent material, hydrotalcites, can be used in the sorption enhanced water gas shift based cap-ture system. The experiments at ECN with a SEWGS single column proofed good long term stability of the hydrotalcite sorbents for CO2 capture. In 2007 combined WGS and CO2 capture experiments in the single column unit (column is filled with mixture of WGS catalyst and CO2 sorbent) also confirmed the technical feasibility of the SEWGS process for pre combustion CO2 capture. Since 2008, continuous SEWGS experiments are being performed on the multi column SEWGS test rig at ECN. In our case, the SEWGS technology is also tested for application in coal gasification combined cycles (IGCC). In the EOS LT CAPTECH consortium, ECN and KEMA have shown that the sorbents of choice (promoted hydrotalcites) are capable of capturing CO2 in presence of hydro-gen sulphide. H2S also adsorbed on the material and was also desorbed in the regeneration step of the cycle. Multiple cycles were performed and no deactivation of the sorbent was observed. Accordingly, the sour-SEWGS represents a simultaneous desulphurization and decarbonisation of the syngas, producing i) a CO2 and H2S stream at low pressure and high temperature and ii) a H2 rich stream at high pressure and high temperature. Description of selected sorbent for SEWGS - Hydrotalcite (HT) is a promising basic solid starting material used as a stable and efficient CO2 sorbent under WGS conditions. The adsorp-tion capacity and stability of an HT based material is tremendously improved by modifying the material by promoting it with potassium carbonate. Although this promotion is frequently used, the origin and mechanism of the promoting effect of potassium or other carbonates remains un-clear.

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http://caesar.ecn.nl/index.php?eID=tx_cms_showpic&file=uploads/pics/power-production-with-co2.gif&width=800m&height=600m&bodyTag=

 Figure 2.6 Structure of Hydrotalcite compound

Hydrotalcites are double layered hydroxide compounds made of metallic oxides like MgO or Al2O3, see Figure 2.6. Their layered structure gives many opportunities to insert material in the structure, making them good candidates for sorbents in the SEWGS process. Hydrotalcites are said to work good as sorbents thanks to a mixed oxide structure, despite their decomposition behaviour which leads to an amorphous solid oxide solution around 350°C. The lack of homogeneity (interlayer spacing) generates access to active sites that are responsible for sorption. But, this material also has some limitations: at working temperature of 400°C, those materials slowly decompose and rearrange in each working cycle. As a consequence, the sorbent loses some capacity which cannot be recovered (kinetic limitation due to the slow regeneration). Fur-thermore, for high a Mg-content hydrotalcite some adsorbed CO2 molecules remain perma-nently on the sorbent, even during steam flow CO2 evacuation step. This is due to the formation of magnesium carbonate, which decarbonation is a priori slower than CO2 desorption on K-O-Al centres.  

2.2 Project description and objectives As shown in Figure 2.7 we need to find a tail gas treatment for the separation of H2S and CO2 resulting from a sour-SEWGS unit. In the best way this unit need to recover CO2 with high pu-rity for transport and sulphur with a commercial value. Of course with the lowest costs (func-tioning operating and installation costs) and best energy efficiency. This treatment unit needs to separate with very high efficiency CO2 and H2S, cool, dry, and compress CO2 for storage and transport. As develop in Table 2.1 and Table 2.2 the SEWGS product is our unit inlet and conditions for CO2 storage needs to be reach in the unit outlet for CO2 gas.   

 Figure 2.7 Place of project in SEWGS process

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Table 2.1  Comparing SEWGS feed and CO2 product species   SEWGS feed  SEWGS product CO  6  %  0  % CO2  24  %  35  % H2  35  % 

3. Comparative study on three IGCC plants

The goal of this chapter is to gather physical and process information on present IGCC plants together. The list, describe in Table 3.1, includes six IGCC cases utilizing General Electric En-ergy (GEE), ConocoPhillips (CoP), and Shell gasifiers each with and without CO2 capture as well as and two case of NGCC (natural gas cycle combined) with and without CO2 capture [2]. Distinguishing Characteristics GEE A key advantage of the GEE coal gasification technology is the extensive operating experience at full commercial scale. Furthermore, Tampa Electric is an IGCC power generation facility, operated by conventional electric utility staff, and is environmentally one of the cleanest coal-fired power plants in the world. The GEE gasifier also operates at the highest pressure of the three gasifiers in this study, 5.6 MPa (815 psia) compared to 4.2 MPa (615 psia) for CoP and Shell. Entrained-flow gasifiers have fundamental environmental advantages over fluidized-bed and moving-bed gasifiers. They produce no hydrocarbon liquids, and the only solid waste is an inert slag. The relatively high H2/CO ratio and CO2 content of GEE gasification fuel gas helps achieve low nitrogen oxide (NOx) and CO emissions in even the higher-temperature advanced combustion turbines. The key disadvantages of the GEE coal gasification technology are the limited refractory life, the relatively high oxygen requirements and high waste heat recovery duty (synthesis gas cooler design). As with the other entrained-flow slagging gasifiers, the GEE process has this disadvantage due to its high operating temperature. The disadvantage is magnified in the single stage, slurry feed design. The quench design significantly reduces the capital cost of syngas cooling, while inno-vative heat integration maintains good overall thermal efficiency although lower than the syn-thesis gas cooler design. Another disadvantage of the GEE process is the limited ability to eco-nomically handle low-rank coals relative to moving-bed and fluidized-bed gasifiers or to en-trained-flow gasifiers with dry feed. For slurry fed entrained gasifiers using low-rank coals, de-velopers of two-stage slurry fed gasifiers claim advantages over single-stage slurry fed.   Distinguishing Characteristics CoP A key advantage of the CoP coal gasification technology is the current operating experience with subbituminous coal at full commercial scale at the Plaquemine plant and bituminous coal at the Wabash plant. The two-stage operation improves the efficiency, reduces oxygen require-ments, and enables more effective operation on slurry feeds relative to a single stage gasifier. The fire-tube SGC used by E-Gas has a lower capital cost than a water-tube design, an added advantage for the CoP technology at this time. However, this experience may spur other devel-opers to try fire-tube designs. However, the two-stage operation results in a quenched syngas that is higher in CH4 content than other gasifiers. This becomes a disadvantage in CO2 capture cases since the CH4 passes through the SGS reactors without change, and is also not separated by the AGR thus limiting the amount of carbon that can be captured.

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Distinguishing Characteristics Shell  The key advantage of the Shell coal gasification technology is its lack of feed coal limitations. One of the major achievements of the Shell development program has been the successful gasi-fication of a wide variety of coals ranging from anthracite to brown coal. The dry pulverized feed system developed by Shell uses all coal types with essentially no operating and design modifications (provided the drying pulverizers are appropriately sized). The dry fed Shell gasi-fier also has the advantage of lower oxygen requirement than comparable slurry fed entrained flow gasifiers. The dry feed entrained-flow gasifiers also have minor environmental advantages over the slurry feed entrained-flow gasifiers. They produce a higher H2S/CO2 ratio acid gas, which improves sulphur recovery and lessens some of the gray water processing and the fixedsalts blowdown problems associated with slurry feeding. A disadvantage of the Shell coal gasification technology is the high waste heat recovery duty (synthesis gas cooler). As with the other slagging gasifiers, the Shell process has this disadvan-tage due to its high operating temperature. The ability to feed dry solids minimizes the oxygen requirement and makes the Shell gasifier somewhat more efficient than entrained flow gasifiers employing slurry feed systems. The penalty paid for this increase in efficiency is a coal feed system that is more costly and operationally more complex. Demonstration of the reliability and safety of the dry coal feeding system was essential for the successful development of the Shell technology. The high operating temperature required by all entrained-flow slagging processes can result in relatively high capital and maintenance costs. However, the Shell gasifier employs a cooled re-fractory, which requires fewer change outs than an uncooled refractory. Life of a water wall is determined by metallurgy and temperature and can provide a significant O&M cost benefit over refractory lined gasifiers.

Net output IGCC from 517 to 640 MW NGCC from 560 to 482 MW

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Table 3.1  IGCC systems studied specifications [2]

  

Table 3.2  IGCC systems efficiency cost capture and emissions [2]    Efficiency  Total  Plant 

Cost ($/kW) CO2  Emissions (lb/MWh)* 

CO2 Capture 

SO2  Emis‐sions (lb/MWh)* 

Shell Gasifier 

41.1 %  1,977  149    0.07 

Shell Gasifier + CO2 capture 

32 %  2,668  1,4  90 %  0.08 

CoP gasifier  39.3 %  1,733  189    0.09 CoP gasifier + CO2 capture 

31.7 %  2,431  1,5  88 %  0.07 

GEE gasifier  38.2 %  1,813  154    0.09 GEE gasifier + CO2 capture 

32.5 %  2,390  1,5  90%  0.07 

NGCC  50.8 %  0.645  783    / NGCC  +  CO2 capture 

43.7 %  1,172  85.8  90 %  / 

* Value is based on gross output 

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  IGCC COMMON PROCESS AREAS  

 

Table 3.3 Pressure and temperature at each stages of IGCC processes [2]  T(°C )/ P(bar)  1  2  2’  3  4  5  6  7 GEE  Without 

CO2  cap‐ture 

60°C 72b 

593°C 55b 

  41°C 51b 

44°C 49b 

237°C 49b 

601°C 1b 

130°C 1b 

CO2  cap‐ture 

60°C 72b 

593°C 55b 

270°C 53b 

39°C 50b 

37°C 47b 

196°C 31b 

566°C 1b 

132°C 1b 

CoP  Without CO2  cap‐ture 

60°C 58b 

204°C 38b 

  39°C 34b 

37°C 34b 

196°C 34b 

599°C 1b 

132°C 1b 

CO2  cap‐ture 

60°C 58b 

204°C 38b 

236°C 35b 

33°C 33b 

37°C 32b 

196°C 31b 

566°C 1b 

132°C 1b 

Shell  Without CO2  cap‐ture 

101°C 1b 

890°C 42b 

  35°C 36b 

51°C 35b 

196°C 31b 

596°C 1b 

132°C 1b 

CO2  cap‐ture 

101°C 1b 

890°C 42b 

In  the same time as cooling 

35°C 33b 

49°C 32b 

196°C 30b 

566°C 1b 

132°C 1b 

 As Table 3.3 shows it, the main difference between the processes are located at the gasifier. That is the reason of temperature and pressure variation at in and output of gasifier. The CoP gasifier have the lowest functioning temperature and pressure. GEE gasifier obtain the cheapest kW price, instead of the highest functioning pressure. The Shell gasifier, as it concern, function with the highest temperature. Downstream the gasifier the temperatures and pressures are similar for the three processes. We must notice the differences pressure and temperature conditions between with and without a CO2 removal stage. In fact temperature and especially pressure are lower after CO2 removal, which is easily understandable because a part of the flow has been removed. Differences

• AGR GEE: Selexol unit CoP: MDEA process Shell: Sulfinol process For CO2 and H2S removal all IGCC use a dual Selexol unit

• Coal pre-treatment Shell process needs pre-drying coal treatment (from 15°C to 101°C with 1b pressure)

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4. AGR technology Gasification of coal to generate power produces a syngas that must be treated prior to further utilization. A portion of the treatment consists of acid gas removal (AGR) and sulphur recovery. The environmental target for IGCC is usually 0.02 kg SO2/MW which requires that the total sulphur content of the syngas be reduced to less than 30 ppmv. This includes all sulphur species, but in particular the total of COS and H2S, thereby resulting in stack gas emissions of less than 4 ppmv SO2.

4.1 COS Hydrolysis The use of COS hydrolysis pretreatment in the feed to the acid gas removal process provides a means to reduce the COS concentration [3]. This method was first commercially proven at the Buggenum plant, and was also used at both the Tampa Electric and Wabash River IGCC pro-jects. Several catalyst manufacturers including Haldor Topsoe and Porocel offer a catalyst that promotes the COS hydrolysis reaction. The non-carbon capture COS hydrolysis reactor designs are based on information from Porocel. In cases with carbon capture, the hydrolysis reactors re-duce COS to H2S. The COS hydrolysis is slightly exothermic equimolar reaction. The reaction is represented by:

COS + H2O ↔ CO2 + H2S ΔHr= -31 kJ/mol Since the reaction is exothermic, higher conversion is achieved at lower temperatures. However, at lower temperatures the reaction kinetics are slower. Based on the feed gas for this evaluation, Porocel recommended a temperature of 177 to 204°C. Since the exit gas COS concentration is critical to the amount of H2S that must be removed with the AGR process, a retention time of 50-75 seconds was used to achieve 99.5 % COS conversion. The Porocel activated alumina-based catalyst, designated as Hydrocel 640 catalyst, promotes the COS hydrolysis reaction without promoting reaction of H2S and CO to form COS and H2. Although the reaction is exo-thermic, the heat of reaction is dissipated among the large amount of non-reacting components. Therefore, the reaction is essentially isothermal. The product gas, now containing less than 4 ppmv of COS, is cooled prior to entering the mercury removal process and the AGR.

4.2 H2S removal via solvents H2S removal generally consists of absorption by a regenerable solvent. The most commonly used technique is based on countercurrent contact with the solvent. Acid-gas-rich solution from the absorber is stripped of its acid gas in a regenerator, usually by application of heat. The re-generated lean solution is then cooled and recirculated to the top of the absorber, completing the cycle. As follow is a general and simplified diagram of the AGR process in Figure 4.1. There are well over 30 AGR processes in common commercial use throughout the oil, chemical, and natural gas industries. However, in a 2002 report by SFA Pacific a list of 42 operating and planned gasifiers shows that only six AGR processes are represented: Rectisol, Sulfinol, me-thyldiethanolamine (MDEA), Selexol, aqueous di-isoproponal (ADIP) amine and FLEXSORB [3]. These processes can be separated into three general types: chemical reagents, physical solvents, and hybrid solvents.   

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 Figure 4.1 Flow Diagram for a Conventional AGR Unit

 

4.3 Chemical Solvents Frequently used for acid gas removal, chemical solvents are more suitable than physical or hy-brid solvents for applications at lower operating pressures. The chemical nature of acid gas ab-sorption makes solution loading and circulation less dependent on the acid gas partial pressure. Because the solution is aqueous, co-absorption of hydrocarbons is minimal. In a conventional amine unit, the chemical solvent reacts exothermically with the acid gas constituents. They form a weak chemical bond that can be broken, releasing the acid gas and regenerating the solvent for reuse. In recent years MDEA, a tertiary amine, has acquired a large share of the gas-treating market. Compared with primary and secondary amines, MDEA has superior capabilities for selectively removing H2S in the presence of CO2, is resistant to degradation by organic sulphur compounds, has a low tendency for corrosion, has a relatively low circulation rate, and consumes less en-ergy. Commercially available are several MDEA-based solvents that are formulated for high H2S selectivity [3]. Chemical reagents are used to remove the acid gases by a reversible chemical reaction of the acid gases with an aqueous solution of various alkanolamines or alkaline salts in water. In Table 4.1 commonly used chemical reagents are listed along with the principal licensors. The process consists of an absorber and regenerator, which are connected by a circulation of the chemical reagent aqueous solution. The absorber contacts the lean solution with the main gas stream (at pressure) to remove the acid gases by absorption/ reaction with the chemical solution. The acid-gas-rich solution is reduced to low pressure and heated in the stripper to reverse the reactions and strip the acid gas. The acid-gas-lean solution leaves the bottom of the regenerator stripper and is cooled, pumped to the required pressure and recirculated back to the absorber.

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For some amines, a filter and a separate reclaiming section (not shown) are needed to remove undesirable reaction byproducts.

Table 4.1 Common Chemical Reagents Used in AGR Processes [2]

  

 Typically, the absorber temperature is 25 to 50°C for amine processes, and the regeneration temperature is the boiling point of the solutions, generally 100 to 130°C. The liquid circulation rates can vary widely, depending on the amount of acid gas being captured. However, the most suitable processes are those that will dissolve 14.8 to 740 m3 acid gas per m3 of solution circu-lated. Steam consumption can vary widely also: 70 to 148 kg per m3 of liquid is typical, with 80 to 92 being a typical “good” value. AGR with MDEA solvent, uses 88 kg of steam per m3 of liquid. The steam conditions are 4.5 bar and 151°C. The major advantage of these systems is the ability to remove acid gas to low levels at low to moderate H2S partial pressures.  

4.4 Physical Solvents

4.4.1 General Physical solvents involve absorption of acid gases into certain organic solvents that have a high solubility for acid gases [3], [6]. As the name implies, physical solvents involve only the physi-cal solution of acid gas – the acid gas loading in the solvent is proportional to the acid gas par-tial pressure (Henry’s Law). Physical solvent absorbers are usually operated at lower tempera-tures than is the case for chemical solvents. The solution step occurs at high pressure and at or below ambient temperature while the regeneration step (dissolution) occurs by pressure letdown and indirect stripping with low-pressure 4.5 bar steam. It is generally accepted that physical sol-vents become increasingly economical, and eventually superior to amine capture, as the partial pressure of acid gas in the syngas increases. The physical solvents are regenerated by multistage flashing to low pressures. Because the solu-bility of acid gases increases as the temperature decreases, absorption is generally carried out at lower temperatures, and refrigeration is often required.

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Most physical solvents are capable of removing organic sulphur compounds. Exhibiting higher solubility of H2S than CO2, they can be designed for selective H2S or total acid gas removal. In applications where CO2 capture is desired the CO2 is flashed off at various pressures, which re-duces the compression work and parasitic power load associated with sequestration. Physical solvents co-absorb heavy hydrocarbons from the feed stream. Since heavy hydrocarbons cannot be recovered by flash regeneration, they are stripped along with the acid gas during heated re-generation. These hydrocarbon losses result in a loss of valuable product and may lead to CO2 contamination. Several physical solvents that use anhydrous organic solvents have been com-mercialized, see Table 4.2. They include the Selexol process, which uses dimethyl ether of polyethylene glycol as a solvent; Rectisol, with methanol as the solvent; Purisol, which uses N-methyl-2-pyrrolidone (NMP) as a solvent; and the propylene-carbonate process.

Table 4.2 Common Physical Solvents Used in AGR Processes [2]

  

4.4.2 The Selexol process The Selexol process was patented by the Allied Chemical Corp. and has been used since the late 1960’s [5], [6]. The process was sold to Norton in 1982 and then bought by Union Carbide in 1990 (Dravo Corp, 1976). Dow Chemical Co. acquired the gas processing expertise, including the Selexol process, from Union Carbide in 2001. The process is offered for license by several engineering companies of which the most experienced in the process is probably UOP. There are more than 55 Selexol plants worldwide, treating natural and synthesis gases. The Selexol solvent is a mixture of dimethyl ethers of polyethylene glycol, and has the formula-tion CH3(CH2CH2O)nCH3, where n is between 3 and 9. There are other process suppliers using the same solvent as the Selexol process. The Selexol solvent is chemically and thermally stable, and has a low vapour pressure that limits its losses to the treated gas. The solvent has a high solubility for CO2, H2S, and COS. It also has an appreciable selectivity for H2S over CO2. The Selexol process can be configured in various ways, depending on the requirements for the level of H2S / CO2 selectivity, the depth of sulphur removal, the need for bulk CO2 removal, and whether the gas needs to be dehydrated.

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Where selective H2S removal is required, together with deep CO2 removal, two absorption and regeneration columns may be required—essentially a two-stage process. Such a process layout is illustrated in Figure 4.2. H2S is selectively removed in the first column by a lean solvent that has been deeply stripped with steam, while CO2 is removed, from the now H2S-free gas, in the second absorber. The second-stage solvent can be regenerated with air or nitrogen if very deep CO2 removal is required. If only bulk CO2 removal is required, then the flashed gas, containing the bulk of the CO2, can be vented and the second regenerator duty can be substantially lowered or be totally eliminated. For selective removal of H2S, without the need for bulk CO2 removal, an absorber, a flash vessel and a steam stripper (regenerator) may be all that is required. The flashed gas is re-compressed back into the absorption column, while the partially stripped solvent flows to the steam stripper.  

 Figure 4.2 H2S and CO2 removal with a Selexol process

For relatively high levels of H2S and COS removal, the Selexol process uses refrigeration to cool the lean solvent to -6°C, -3°C. This allows for up to 99% of the COS to be removed, but at the expense of a high solvent circulation rate and appreciable CO2 co-absorption. This also re-sults in a low H2S content of the acid gas going to sulphur recovery and high overall unit costs. A lower cost configuration, one that results in a high H2S-content acid gas (low CO2 coabsorp-tion), allows much of the COS to bypass the absorber. The relatively poor selectivity between COS and CO2 (about 2.3) is the cause of this. A COS hydrolysis unit may be required, if both a high H2S-content acid gas and a high level of COS removal are to be achieved. However, the process can be configured to give both deep removal of sulphur compounds and selective H2S over CO2 removal if several absorption, flash, and regeneration stages and a chilling system is used. Such process configurations tend to be complex and costly [6].  

4.5 Hybrid Solvents Hybrid solvents combine the high treated-gas purity offered by chemical solvents with the flash regeneration and lower energy requirements of physical solvents. Some examples of hybrid sol-vents are Sulfinol, Flexsorb PS, and Ucarsol LE. Sulfinol is a mixture of sulfolane (a physical solvent), diisopropanolamine (DIPA) or MDEA (chemical solvent), and water. DIPA is used when total acid gas removal is specified, while MDEA provides for selective removal of H2S.

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Flexsorb PS is a mixture of a hindered amine and an organic solvent. Physically similar to Sulfinol, Flexsorb PS is very stable and resistant to chemical degradation. High treated-gas pu-rity, with less than 50 ppmv of CO2 and 4 ppmv of H2S, can be achieved. Both Ucarsol LE-701, for selective removal, and LE-702, for total acid gas removal, are formulated to remove mercap-tans from feed gas. Mixed chemical and physical solvents combine the features of both systems. The mixed solvent allows the solution to absorb an appreciable amount of gas at high pressure. The amine portion is effective as a reagent to remove the acid gas to low levels when high purity is desired. Mixed solvent processes generally operate at absorber temperatures similar to those of the amine-type chemical solvents and do not require refrigeration. They also retain some advantages of the lower steam requirements typical of the physical solvents. Common mixed chemical and physi-cal solvent processes, along with their licensors, are listed in Table 4.3. The key advantage of mixed solvent processes is their apparent ability to remove H2S and, in some cases, COS to meet very stringent purified gas specifications. Figure 4.3 shows reported equilibrium solubility data for H2S and CO2 in various representative solvents. The solubility is expressed as standard cubic feet of gas per gallon liquid per atmos-phere gas partial pressure. The figure illustrates the relative solubilities of CO2 and H2S in dif-ferent solvents and at various temperature. It shows an order of magnitude higher solubility of H2S over CO2 for all solvent mentioned in the index at a given temperature. This gives rise to the selective absorption of H2S in physical solvents. It also illustrates that the acid gas solubility in physical solvents increases with lower solvent temperatures.

Table 4.3 Common Mixed Solvents Used in AGR Processes [2]

  

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 Figure 4.3 Equilibrium Solubility Data on H2S and CO2 in Various Solvents

The ability of a process to selectively absorb H2S may be further enhanced by utilizing the rela-tive absorption rates of H2S and CO2. Thus, some processes, besides using equilibrium solubil-ity differences, will use absorption rate differences between the two acid gases to achieve se-lectivity. This is particularly true of the amine processes where the CO2 and H2S absorption rates are very different.  

4.6 Balance First of all, CO2 recovery from the chemical solvents is achieved by the application of heat whereas for physical solvents CO2 can be stripped by reducing the pressure without the applica-tion of heat [7]. For H2S recovery however, both type of sorbents require steam stripping. Moreover chemicals solvents are corrosive contrary to physical ones, for these reasons, ener-getic, low capital expenditure and lower operational expenditure we will choose preferably a physical solvent process for our project. Two candidate might be suitable; the Selexol and Purisol process. Both of them need the same equipment and function in the same way. Let it be noticed that, as put above, our case requires selective removal of H2S, without bulk CO2 removal, an absorber, a flash vessel and a steam stripper (regenerator) may be all that is required. Notice too that for the same CO2 solubility the H2S solubility is 8.82 for Selexol and 10.2 for Purisol at 25°C. But to increase again selec-tivity, limit solvent loss, and achieve deep H2S removal, chilling to -6°C, -3°C might have to be required.

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Table 4.4 Selexol/Purisol process balance Selexol/Purisol process + -

Conversion 100-99.5% (chilling) Selectivity Large CO2 absorption without chill-

ing; but with and a flash regenerator connected to absorber, the selectiv-

ity will be considerably increase Feed Wet/Dry 1% H2S

Solvent Selexol : none

Lost with chilling Purisol : 3.63 kg/h

Pre-treatments Need compressor unit before ab-sorber

Post-treatments No need drying unit for CO2

Drying unit for H2S follows by Claus unit

Regeneration way Steam stripper for H2S recover 3t of steam per produced ton sulphur

Costs 4000-5000 Kw/h Coolers, boiler, pumps, compressors,

steam Others Need a Claus unit for H2S gas treat-

ment  On the whole we realize that such processes are not suitable for our project, in particular due to energetic functioning cost and large CO2 removal. In addition a such process need a Claus unit to treat the H2S/CO2 (20% CO2 estimated) in outlet.  

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5. Claus process

5.1 Generality More than 100 years ago, the Claus sulphur recovery process was introduced as a means of dis-posing of H2S generated in the recovery of ammonium sulfate from gasworks ammoniacal liq-uor [3], [7], [8]. Later, it was extended to process H2S from the LeBlanc soda ash process. To-day, a modification of the original process is the basis of hundreds of sulphur recovery units worldwide, with individual capacities up to 2000 t/d, or more. They process H2S from a variety of sources; some comes from the chemical industry, but a large majority is produced in the processing of fossil hydrocarbon fuels and coal. In the original process, the Claus reaction was carried out in a single step over a catalyst. Be-cause the heat of reaction was dissipated by radiation alone, sulphur yields of 80–90% were possible only if the H2S space velocity was low. Processing capacity was not significantly in-creased by installing cooling coils within the catalyst bed, nor by recycling cooled waste gas through the reactor. It was not until the 1930s that IG Farben industry in Germany introduced a revolutionary development which gave improved efficiency at much higher space velocities and also allowed waste heat to be recovered as steam. This consisted of separately burning one-third of the H2S in a flame with the stoichiometric amount of air to convert it completely to sulphur dioxide, cooling the hot gases in a waste heat boiler, and introducing them, with the remaining two-thirds of the hydrogen sulfide, to the catalytic reactor at 200–300°C.

H2S + 3/2O2 → SO2+H2O Combustion 2H2S + SO2 ↔ 3/n Sn + 2 H2O Thermodynamic equilibrium

This was later modified so that all the H2S, with the stoichiometric amount of air, was passed through the flame and waste heat boiler, followed by a sulphur condenser. Because a major pro-portion of the sulphur was formed in the flame zone and removed in a condenser, the loading on the catalytic stage was reduced, and the conversion efficiency was correspondingly raised. In this form, the process was able, in a single stage, to convert H2S in concentrations of ca. 15% or more to sulphur, with a conversion efficiency of 92–94% at space velocities of 250–300 m3 H2S per h m3 catalyst. This version has become known as the modified Claus process. The efficiency is raised by adding one, two, or three further catalytic stages, in which the gases leaving the sulphur condenser of the previous stage are reheated and passed through another catalytic reactor, followed by another sulphur condenser. The sulphur condensers again allow the thermodynamically limited H2S-SO2 reaction to proceed. Plants operating on this principle are usually referred to simply as Claus plants. It is interesting to note that the original Claus process (under the name “direct oxidation”) has become of interest again as a method of dealing with hydrogen sulphide streams that are too weak for processing by the modified process. De-spite the enormous accumulated experience in design and operation of Claus plants, the process is inherently difficult to maintain at full efficiency. Under completely stable conditions, the process can operate consistently at its design efficiency, but any perturbation may lead to a quite substantial loss of efficiency. The problem is partly due to the fact that the chemical reactions involved do not go to completion. As environmental and safety regulations have become more demanding, it has become neces-sary to recover sulphur at high efficiency from gas streams which, formerly, were considered too lean or too impure for the process, and would have been incinerated and discharged to the atmosphere. The German TA Luft demands sulphur recovery rates of 97 % for recovery units with a sulphur capacity ≤ 20 t/d, 98 % for units with capacity 20–50 t/d, and 99.5 % for units of capacity ≥ 50 t/d.

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As the chemical reactions involved in the Claus process do not go to completion, these high sul-phur recovery rates are impossible to reach without the aid of special tail gas units.

Figure 5.1 H2S conversion and sulphur composition according to temperature [8]

The first real study of the thermodynamics of the H2S–O2–S–H2O equilibrium was performed in 1953 by Gamson and Elkins [9], who produced a curve expressing the temperature dependence of the calculated equilibrium conversion of a stoichiometric mixture of pure H2S and air to ele-mental sulphur (dashed line in right hand figure of Figure 5.1). At that time, only the species S2, S6, and S8 were known to exist, but since then S3, S4, S5, and S7 have been shown to be significant in certain temperature ranges (Figure 5.1 left), leading to re-vised curves for the temperature dependency of the conversion (dotted and solid lines in Figure 5.1) [10]. The left-hand region of these curves represents the equilibrium between H2S and O2. As the partial pressure of each component is proportional to its molar concentration in the gas mixture, for a given mass of sulphur, the higher the molecular mass of a species, the lower its molar concentration and, therefore, the lower its partial pressure (1 mol S8 = 2 mol S4 = 4 mol S2). As Figure 5.1 shows, the proportion of S8 falls with increasing temperature in fa-vour of S6, S7, S5, and, especially, S2, and therefore the partial pressure of sulphur increases very rapidly with temperature, promoting the reverse reaction. This effect tails off above ca. 550°C, by which point the majority of the sulphur is S2, and the equilibrium sulphur conversion begins to rise again, because a significant proportion of the H2S is beginning to decompose into its elements. This is the main mechanism by which sulphur is produced in the sulphur furnace. When the hot gas from the burner is cooled, the sulphur con-version at first drops to the minimum represented on graphic, as the dissociated hydrogen and sulphur recombine, and then rises along the right portion of the curve, as the oxidation of hy-drogen sulphide progresses. However, as the temperature is lowered, the reaction becomes pro-gressively slower until, below about 350°C, it is too slow to be of practical use. At this tempera-ture, the best theoretical yield (assuming the system has reached equilibrium) is only about 80–85 %. The function of the catalyst in a Claus plant is to extend the lower end of the temperature range over which the reaction takes place at a satisfactory rate to the lowest possible tempera-ture so as to attain the highest possible conversion. The catalyst has no influence on the ultimate position of the equilibrium, only on the rate at which it is attained. Since the position of the

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equilibrium is a function of the partial pressures of the individual components of the gas, if the partial pressure of any one of them is disturbed, the values of the rest will adjust so as to bring the value of thermodynamic equilibrium back to K. It is not possible to increase conversion efficiency by using excess oxygen, because although a great proportion of the H2S reacts, it forms SO2 rather than sulphur. Conversion is, however, favoured, if the hydrogen sulphide is highly concentrated, and it is further increased if the nitrogen content of the combustion air is reduced or eliminated altogether. Because water is one of the products, a high moisture loading in the feed gas has a detrimental effect on conversion efficiency. A fortunate charac-teristic of the reaction system is that one of the products, sulphur, is much less volatile than any of the other constituents, and can therefore be removed by condensation. This reduces its partial pressure considerably, so that if the remaining gases are reheated and passed over a catalyst, they react again to produce more sulphur. Furthermore, since the sulphur loading of the reaction gases is reduced at each stage, the succeeding catalytic stage can be run at progressively lower temperature without risk of sulphur deposition on the catalyst. This is the principle on which the design of practical Claus plants is based.  

5.2 Claus Plant Configurations Figure 5.2 and Figure 5.3 show the two main process configurations which are in use for proc-essing acid gases in Claus plants with 3 stages.  

5.2.1 Straight-Through Configuration. All of the acid gas is passed through the reaction furnace and waste heat exchanger, see Figure 5.2. A significant amount of sulphur is produced in the furnace, and removed in the waste heat exchanger and in the condenser before the gas enters the first catalytic converter. The amount of sulphur condensed in the waste heat exchanger depends on the pressure of the steam produced in the waste heat exchanger, and on the gas temperature at the outlet of the waste heat exchanger. The straight-through configuration is used for feed gases containing H2S at ca. 40 mol% or more. When the H2S content of the feed gas is below ca. 40 mol%, the flame tempera-ture in a straight-through Claus plant is too low to destroy satisfactorily any carbon or ammonia compounds present in the acid gas. The sulphur produced may be discoloured and there may be a risk of plugging the sulphur lines in the “cold” part of the Claus plant. However, with acid gas and/or air preheaters upstream of the combustion furnace, or with oxygen instead of air for the partial oxidation of the H2S, acid gases with H2S levels of 20 mol% and even lower can be proc-essed.

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 Figure 5.2 Claus plant schemas Straight-Through Configuration [8]

 

5.2.2 Split-Flow Configuration In a split flow configuration, see Figure 5.3, acid gases containing H2S at ca. 20–40 mol% can be processed while still utilizing air in the burner. Part of the acid gas is fed directly to the first catalytic converter, and the remainder is burned in the combustion furnace. In theory, two-thirds of the acid gas may be bypassed this way, and in that extreme case essentially no sulphur is pro-duced in the furnace, because one-third of the H2S is burned completely to sulphur dioxide. A disadvantage is that component that are destroyed in the furnace of the straight-through Claus configuration are now bypassed toward the catalytic stages.  

 Figure 5.3 Claus plant schemas Split-Flow Configuration Source[8]

 

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5.3 The Superclaus Process, the Euroclaus Process The Dutch companies Comprimo and VEG-Gasinstituut, in cooperation with the University of Utrecht, have developed a new version of the Claus process, called the Superclaus process, which increases the sulphur recovery rate of a Claus plant to > 99 %. Two versions have been developed: Superclaus-99 and Superclaus-99.5. The Superclaus process produces elemental sulphur from H2S gas, and is characterized by the use of a new selective oxidation catalyst in the last reactor stage. The process has the following features: 1) Overall sulphur recovery up to 99.5% without tail gas treatment 2) Application in new as well as existing Claus plants 3) Low investment and utility costs 4) Continuous catalytic process without water condensation 5) Use of a new catalyst for direct selective oxidation of H2S to sulphur The new catalyst has the unique ability to convert H2S with excess oxygen into elemental sul-phur at > 90% efficiency. Moreover, high water concentrations in the process gas hardly influ-ence the conversion [7] [8].  

5.3.1 Superclaus-99 The conventional Claus process employs a series of steps consisting of:

(1) sulphur conversion (2) cooling, sulphur condensation and removal (3) process gas reheating before the next conversion step.

Since the Claus equilibrium is involved in each conversion step, a close approximation to equi-librium is reached in the final stage. The standard Claus process has several limitations which are due to:

• failure to reach equilibrium; • increasing water content of the process gas with decreasing concentration of hydrogen

sulfide and sulphur dioxide; • formation of COS and CS2 in the flame, which requires hydrolysis in the first Claus re-

actor; • and the demand for exact control of the air : acid gas ratio.

The Superclaus process incorporates two new concepts, see Figure 5.4. These are a less sensi-tive and more flexible air: acid gas control, and a new catalyst for selective oxidation of H2S di-rectly to sulphur, rather than to SO2. The acid gas feed is burned in the thermal stage in such a manner that the tail gas leaving the second Claus reactor stage contains 0.8–1.5 mol% H2S. The burning step is operated under oxygen-deficient conditions so that the H2S/SO2 ratio in the fur-nace is much higher than the conventional value of 2:1. As the process gas passes through the various stages, the presence of surplus H2S suppresses the concentration of SO2 in the gas. As a result, the second Claus reactor outlet may contain concentrations of 0.8 mol% H2S and 0.05 mol% SO2. In the Superclaus stage, the gas leaving the second-stage Claus sulphur con-denser is reheated and mixed with surplus air before it enters the Superclaus reactor. Here, an essentially non equilibrium oxidation reaction occurs:

H2S + ½ O2 → 1/n Sn + H2O selective oxidation S + O2 → SO2 sulfur oxidation, side reaction

The new Superclaus catalyst is the key to this oxidation step. Over 85% of the H2S is oxidized directly to sulphur, and very little SO2is formed. The catalyst consists of a carrier on which

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metal oxides are supported such as Fe, Fe/Cr catalyst supported on alumina or silica. The catalyst is chemically and thermally stable and has good mechanical strength. It has a specific surface area < 20 m2/g, compared with the Claus catalyst value of 300–325 m2/g. The Super-claus catalytic stage is followed by a final sulphur condenser and coalescer vessel, and then the incinerator/stack [7] [8]. The main features of the Superclaus-99 process are represented in Figure 5.5.

Figure 5.4 Superclaus-99 process schema [8]

Figure 5.5 Superclaus-99 process schema balance

5.3.2 Superclaus-99.5. The Superclaus-99.5 process consists of a thermal stage, two or three catalytic Claus stages, a hydrogenation stage, and a selective H2S oxidation stage, see Figure 5.6. The thermal stage and the catalytic Claus reactors are operated conventionally, with an H2S:SO2 ratio of 2:1. The com-bustion air to the burner is controlled by a conventional H2S/SO2 tail gas analyser. As only H2S is oxidized to elemental sulphur, other sulphur components, e.g., SO2, COS, and CS2, must be converted first to H2S in the hydrogenation reactor. The gas from the hydrogenation reactor is cooled to the optimum inlet temperature of the selective oxidation reactor. The control of the H2S:SO2 ratio of 2:1 is less critical than for standard Claus operation. Fluctuations in the feed gas composition to the unit are absorbed by the excess air supplied to the selective oxidation

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stage. The total recovery with two or three catalytic Claus reactors is 99.2–99.7 %, depending on the oxidation efficiency [7] [8]. The main features of the Superclaus process are represented in Figure 5.7.

 Figure 5.6 Superclaus-99.5 process schema

 Figure 5.7 Superclaus-99.5 process schema balance

 

5.3.3 The Euroclaus process The Euroclaus process is an extension of the Superclaus process, see Figure 5.8. It exhibits en-hanced efficiency of the selective oxidation section and in particular a reduction of the SO2 level in the Claus product gas. The product gas from the Claus section contains some SO2, which is not converted in the Superclaus reactor. The amount of SO2 in Claus tail gas is usually in the range of 500 – 1200 vol ppm wet. This amount depends on the catalyst activity of the preceding Claus reactor and the amount of H2S in the process gas (ratio H2S to SO2) [8]. The main features of the Euroclaus process are represented in Figure 5.9.  

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 Figure 5.8 Euroclaus process schema [7] [8]

 

 Figure 5.9 Euroclaus process schema balance

 

5.4 Sub-Dewpoint Processes Sulphur recovery of the modified Claus process is limited by the operating temperatures in the catalytic reactor stages. The lower the operating temperatures, the higher the recovery rates. However, operating the Claus reactors at lower temperatures, i.e., below the sulphur dewpoint, involves adsorption of sulphur on the catalyst, which has to be regenerated after a certain opera-tion time. So-called sub-dewpoint processes have been developed, all of which utilising the more favourable equilibrium at lower temperatures. These processes can be simply described as an extension of the conventional modified Claus process. All are cyclic processes, adsorbing the sulphur produced on the catalyst. The catalyst is regenerated before it becomes significantly de-activated. In the beginning, around 1970, the maximum achievable sulphur recovery rate with the modified Claus plus sub-dewpoint process, was ca. 98.7–99.2 %; in the mid-1990’s, sulphur recovery rates of 99.0–99.9% became possible. The best known sub-dewpoint processes are:

1) The Sulfreen processes 2) The CBA processes 3) The MCRC processes

Common to them all is the ability of the catalyst, not only to adsorb the H2S and SO2 from the process gas and act as catalyst for the Claus reaction, but also to have an affinity for the sulphur produced. Product sulphur is adsorbed on the walls of the pores in the catalyst. This reduces the

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partial pressure of sulphur vapour and allows additional H2S and SO2 to react in accordance with the reaction equilibrium. They are all cyclic processes and a regeneration step is required. Regeneration can be achieved with heated feed or steam.

5.4.1 The Sulfreen, Hydrosulfreen, and Doxosulfreen Processes Claus plants, as a rule consisting of one thermal and two or three catalytic stages, generally con-vert 94–98% of the sulphur compounds in the feed gas into elemental sulphur. To increase the sulphur recovery rate in order to reduce SO2 emissions, the Sulfreen processes have been devel-oped by Lurgi, Germany; and Elf Aquitaine, France. The first Sulfreen plant was started in 1970 in Lacq, France; by 2005, more than 50 had been built. A wide degree of flexibility is possible with the Sulfreen processes. Three versions are commercially available: 1) The Sulfreen process, based on an extension of the Claus process, where residual H2S and

SO2 in the Claus tail gas reacts on a fixed catalytic bed at a temperature below the sulphur dewpoint. At such a low temperature (125–150°C) equilibrium conversion of H2S and SO2 is nearly complete and sulphur recovery of 99.5 % or more can be achieved [7] [8].

2) The Hydrosulfreen process, which overcomes the limitation of the Sulfreen process due to

COS and CS2 present in the Claus tail gas. It is implemented by converting these com-pounds to H2S in a catalytic pretreatment step upstream of the Sulfreen reactor. This process can achieve an overall sulphur recovery of 99.5–99.7 % [7] [8].

3) The Doxosulfreen process catalytically converts residual H2S and SO2 from the upstream

Sulfreen step with oxygen into elemental sulphur. A sulphur recovery rate of 99.9 % can be obtained [7] [8].

5.4.1.1 Sulfreen The Sulfreen process is based on the well-known Claus reaction in which H2S and SO2 in the process gas are catalytically converted to elemental sulphur, see Figure 5.10. The process oper-ates in the gas phase, the operating conditions being those at which the tail gas leaves the up-stream Claus plant. The catalyst, arranged in fixed beds, is activated alumina (ultramacropo-rosity, between 0.1 and 1 μm, reduces diffusional constraints and a low sodium (Na2O) content ( < 2500 ppm) reduces the sulphation, thus slowing down the deactivation of the catalyst), with properties similar to those of Claus catalysts. The tail gas from the Claus plant is passed at a temperature of 120–140°C through one of the two reactors where, by means of the catalyst, most of the H2S and SO2 are converted to elemental sulphur and adsorbed on the catalyst. The sulphur-loaded catalyst can be regenerated by using part of the Claus tail gas heated at 315°C. The regeneration gas is heated in an indirectly fired heater. Hot flue gas from thermal incineration or directly fired heaters can also be used. The sulphur in the hot regeneration gas is recovered in the sulphur condenser. After subsequent cooling of the catalyst bed with purified tail gas, the reactor is again ready to be switched to adsorption.

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 Figure 5.10 Sulfreen process schema [7] [8]

 

5.4.1.2 Hydrosulfreen The Hydrosulfreen process is based on the pretreatment of the Claus tail gas by hydrolysis of COS and CS2 to H2S, and direct oxidation of H2S to sulphur, see Figure 5.11:

COS + H2O → CO2 + H2S CS2 + 2 H2O → CO2 + 2 H2S 2H2S + O2 → 2/n Sn + 2 H2O

These reactions are performed by bringing the Claus tail gas directly into contact with an acti-vated titanium oxide catalyst. In a second process step, the remaining sulphur compounds in the process gas are removed as in the Sulfreen process. The tail gas from the upstream Claus unit (temperature ca. 130°C) is heated in tail gas preheaters to ca. 200–250°C, and passed to the hydrolysis/oxidation reactor where the temperature rises to ca. 250–300°C, owing to the exo-thermicity of the reactions. For gas heating, a gas–gas heat exchanger in the waste gas down-stream of the final incinerator, or directly fired heaters can be used.  

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 Figure 5.11 Hydrosulfreen process schema [7] [8]

 

 Figure 5.12 Hydrosulfreen process schema balance

 

5.4.1.3 Doxosulfreen The Sulfreen process is based on the Claus equilibrium, which at 130–140°C lies almost com-pletely on the product side. At the outlet of the Sulfreen reactor the concentrations of H2S and SO2 are already greatly lowered, and the rather high concentration of water vapours favours the back reaction to a certain extent. Hence, the potential of the Claus reaction is almost exhausted. Further desulphurization in the Doxosulfreen process is therefore based on direct oxidation of H2S with oxygen at temperatures similar to those of the Sulfreen process. The upstream plant is operated with a small excess of H2S so that the tail gas of the Sulfreen step contains ca. 2500 ppm H2S and 250 ppm SO2. A controlled amount of air is mixed with the gas (O2/H2S ratio ap-proximately twice the stoichiometric ratio), and the gas is passed through the catalytic reac-tor. On the catalyst surface, with the excess of H2S, the Claus reaction is further completed ac-cording to:

10 H2S + SO2 → 3/n Sn + 2H2O + 8 H2S

The remaining H2S is oxidized to elemental sulphur:

2H2S + O2 → 2/n Sn + 2 H2O  

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Depending on the type of catalyst, a small amount of sulphuric acid is formed by oxidation of SO2. The products are adsorbed in the pore system of the catalyst such as impregnated modi-fied alumina with Cu. The plant essentially consists of two reactors for pretreatment of the Claus tail gas (Sulfreen stage) and two reactors filled with direct oxidation catalyst, see Figure 5.13. One of these reactors operates in the adsorption mode and the other in the regenera-tion mode. The regeneration loop comprises the sulphur condensers as well as the blower and the heater for the regeneration gas. Both stages are regenerated simultaneously. The tail gas from the Claus plant is passed through the reactor containing the Sulfreen catalyst at 130–140°C. The gas leaving the Sulfreen unit is cooled to ca. 125°C before adding the air for oxida-tion. The tail gas from the Doxoreactor is passed to the incinerator, where the very small amounts of sulphur components are converted to SO2. Regeneration of the catalyst is carried out with tail gas from the Claus plant, which is heated to 300–330°C by a special gas heater or by heat exchange with hot flue gas from the incinerator. This gas contains 1–2% H2S, so any sulphuric acid formed during adsorption is reduced to elemental sulphur. The gas is cooled to ca. 125°C in a sulphur condenser, and the elemental sulphur is drawn off in liquid form. The regeneration gas is recycled. Following de-sorption, the reactors are cooled to operating temperature again.  

 Figure 5.13 Doxosulfreen process schema [7] [8]

 

 Figure 5.14 Doxosulfreen process schema balance

 

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5.4.2 The CBA/MCRC Process

5.4.2.1 The CBA Process The Cold Bed Adsorption (CBA) process is a dry-bed, sub-dewpoint, adsorption process based on the extension of the modified Claus process to lower temperatures (below the sulphur dew-point). The sub-dewpoint reactors use activated alumina catalyst operating below the sulphur dewpoint (120-150°C) to form elemental sulphur. The Claus conversion is increased because of the greater equilibrium conversion at sub-dewpoint temperatures, and because the sulphur is ad-sorbed on the catalyst which further shifts the reaction to the right.

2 H2S + SO2 ↔ 3/n Sn + 2 H2O However, when the sulphur load on the catalyst exceeds a certain quantity, the catalyst activity decreases. To restore catalyst activity to an acceptable level, the catalyst must be periodically thermally regenerated and the sulphur recovered. This regeneration step is accomplished by flowing hot process gas, usually the first reactor effluent, through the deactivated bed. After sulphur desorption, the catalyst is cooled to the operating temperature. High levels of COS and CS2 in the tail gas reduce the overall sulphur recovery as they are not recovered in the CBA process. The CBA process is offered in a number of configurations from two to five reactors. Most ear-lier installations consisted of four reactors (two high temperature Claus reactors and two sub-dewpoint CBA reactors as shown in Figure 5.15), while most recent CBA installations have a total of three (a single high-temperature reactor and two CBA reactors as shown in Figure 5.16). As of January 2000, there were a total of 28 CBA units in design or in commercial operation. For sulphur recovery applications with overall recovery requirements in the 99% range, CBA is an option to be considered.

Figure 5.15 Four-Reactor CBA process schema [7] [8]

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Figure 5.16 Three-Reactor CBA process schema [7] [8]

 

5.4.2.2 The MCRC Process The MCRC™ process is also a dry-bed sub-dewpoint adsorption process based on the extension of the modified Claus process. The process uses a proprietary high macro porosity alumina cata-lyst in a reactor operating below the sulphur dewpoint to form elemental sulphur. The MCRC™ process can use either a three- or four-converter configuration. The three bed process is a three-stage Claus unit with the last two stages converted to sub-dewpoint reactors [11]. The Canadian licensor, Delta Hudson Ltd., states that typical overall sulphur recoveries for a sulphur recovery unit with a MCRC™ unit are 98.5 to 99.0% for three converters and 99.0 to 99.5% for four converters. There are several operational and configurational differences that distinguish MCRC process from CBA one. With MCRC, the gases leaving the first converter (Claus converter) are cooled to partially condense the sulphur formed in the reaction, then reheated before entering in the re-generating converter. The MCRC process does not include a cooling period In the cycle, but switches the converter directly from regeneration to adsorption. With a three converter unit this result in a slight dip in sulphur recovery for a period of time after bed switching while cool down occurs, although the overall recovery remains above 99 %. There is especially no dip in recovery with the four converter unit as the two sub-dewpoint beds operating in series dampen any impact caused by bed switching. A distinctive feature of the MCRC process is that each sub-dewpoint converter has its own designated condenser. The converter and condenser operate together as a unit during each step of the cycle. This arrangement is claimed to require fewer switching valves and less plot space, which reduces piping costs. Also, sulphur emission to the atmosphere are minimized when the catalyst beds are purged to shutdown.

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 Figure 5.17 MCRC process schema [11]

 

5.5 SCOT / Super-SCOT processes SCOT Process (Shell Claus Off-gas Treating Process) was developed by Shell, and introduced in the early seventies as an attractive process for improving the efficiency of a Sulphur Recov-ery Unit (SRU; Claus Unit). Since the first unit was started-up in 1973, more than 120 units have been built with a wide range of capacities. By using the SCOT process, the overall sulphur recovery of SRU can reach about 95% to 99.9%. while the Super-SCOT process can give 99.95% recovery, which makes it capable of coping with the most stringent legislation. The Claus tail gas, after being reduced in the reactor, SO2, COS, CS2, and elemental sulphur, are