Due to the very high fossil energy demand, the glass industry is looking for innovative approaches for the reduction of CO 2emissions and the integration of renewable energy sources. In this paper, we present a novel power-to-gas concept, which has no impact on established melting processes and discuss it for this purpose. A special focus is set on the required CO 2capture from typical flue gases in the glass industry, as this process has not been investigated in detail yet. We used a process simulation approach to investigate post-combustion CO 2capture by absorption processes, followed by a techno-economic evaluation. Our inves- tigations found the designed CO 2capture plant to be approx. 400 times smaller than absorption based CO 2 separation processes for conventional power plants. Due to the many options for waste heat utilization, the waste heat required for CO 2desorption can be generated in a particularly efficient and cost-effective way. The resulting CO 2avoidance costs range between 41-42 C/t CO 2, depending on waste heat utilization for desorp- tion, and thus offer a cost effective way of CO 2removal from glass industry melting processes. These costs are well below the values of 50-65 C/t CO 2described so far for comparable industrial applications. In addition, we describe optimization options, like solvent and process improvements, to enable further cost reductions. These results motivate further research and development on the overall process presented in this work. Keywords: Power-to-Gas, methanation, oxyfuel, glass industry, CO 2-separation, economic evaluation 1 Introduction Commercial glass production is a very energy-intensive industrial process. Converting the raw materials, such as silica sand, sodium carbonate, lime, dolomite and others into molten glass requires high process temperatures of up to 1600 °C. This melting process is the central phase of glass production and accounts for 50-80 % of total energy consumption in overall glass production. The dominating source for achieving the required process temperatures has been the combustion of fossil fuels such as natural gas (NG) and crude oil for a long time. [1] Due to this dominance of fossil fuels, glass production is currently still associated with high carbon dioxide (CO 2) emissions. In 2007, about 27.5 million tonnes CO 2-equivalent (eq.) were emitted within the European Union, the world’s largest glass producer (approx. 30 % of world output). Approx. 17 mio. t CO 2-eq. of these emissions, i.e. 81 %, are exclusively due to emissions from combustion for the melting process [2]. Although there has been a slight reduction in CO 2emissions since then, it fails to meet climate protection targets. The causes of this decline also include economic fluctuations and consumer demand. For example, the level of these emissions in Germany has stagnated at a constant level of about 4 mio. t CO 2-eq. per year since 2007 [3]. 1Corresponding author: Sebastian G ¨artner; E-mail: s.gaertner@tum.de; Postal address: Seybothstraße 2, D-93053 Regensburg, Germany 1 Of course, there are efforts towards increased sustainability and climate protection in the glass industry. Accordingly, there are many research approaches and developments for decarbonization and the integration of renewable energies into the glass melting process. The most important concepts are all-electric melting and the switch to hydrogen (H 2) as combustion fuel instead of NG [4]. However, all of these concepts, despite their promising CO 2emissions reduction potential, have certain restrictions: All-electric melting is well established at small scale glass melting systems, but large scale applications are still controversial [5]. Recent all-electric melting projects have improved and could reach melting capacities of up to 250 t/d. However, complex melting tank design and extensive heat control strategies of such large scale melting tanks are challenging and result in high investment and operating costs [6]. In addition, low glass production rate (so called pull rate) flexibility, short melting tank lifetime, high electricity costs and low operating experience are disadvantages. Besides that, not all glass types are feasible for all electric melting, for example non-ionic glasses [4]. Recent projects have investigated hydrogen as a fuel substitution option, like HyGlass [7], HyNet [8] and Koper-nikus P2X [9]. Since the combustion of H 2does not cause any CO 2emissions, this would significantly reduce the CO 2footprint of glass melting. However, this reduction highly depends on the way of hydrogen production (green, blue, yellow,... hydrogen). So far, only green H 2can guarantee a high reduction of CO 2 emissions if electricity comes exclusively from renewable energies, as for example blue H 2also comes with 9-12 % less CO 2emissions than conventional (= grey) hydrogen [10]. However, green H 2has been cost- intensive and therefore not considered to be economically competitive with established fossil fuels. By contrast, recent price dynamics on the large-scale markets may indicate a turnaround in terms of cost competitiveness of green H 2. Nevertheless, H 2shows much different combustion behavior, like higher flame temperatures, different flame velcoity, faster ignition behavior and changed heat radiation properties. Some of these effects are compensated by oxyfuel combustion, for example changes in adiabatic flame temperature, as shown in [6]. Nevertheless, there are still some questions unanswered, such as effects of higher water content in the flue gases in the firing chamber on glass quality, and melting tank life time. Due to these limitations of currently discussed decarbonisation options, there is an urgent need for innovative energy concepts which allow the integration of renewable energies, a reduction of CO 2emissions, and which have ideally no impact on established melting processes. This work is the first to demonstrate the integration of a power-to-methane (PtM) system into the glass melting process that meets all of these requirements. Since the PtM concept has been known for a long time as a sector coupling option in energy systems [11, 12], there are a lot of established technical options for most of the process steps. The novelty of these established options is briefly described, before the remaining open process of CO 2capture from exhaust gases is discussed in more detail. In addition, a techno-economic analysis is provided, highlighting the economic aspects of CO 2 separation for this specific application. 2 Integration of Power-to-Methane into Oxyfuel Glass Melting Figure 1 shows a simplified process flow sheet for the integration of PtM into oxyfuel glass melting processes. In this section, each process unit in the flow sheet is described, along with established technical options for its technical implementation. We paricularly focus on CO 2separation from the flue gases of oxyfuel glass melting tanks, since this process has not been investigated yet. 2.1 Electrolysis In the first process unit, water (H 2O) is separated into (H 2) and oxygen (O 2) by electrolysis. Both H 2and O 2 are subsequently stored to overcome the intermittent supply of renewable electricity from wind or solar power (sections 2.2). Besides H 2, also O 2is an important product for the downstream PtM processes, as it can be used in oxyfuel combustion (section 2.4). Alkaline (AEL), solid oxide (SOEC) and polymer electrolyte membrane electrolysis (PEMEL) are currently considered as the technically most established processes for this purpose [11, 13]. AEL has the highest 2 Storage system 1 Hydrogen (H2)Storage system 2 Oxygen (O2) H2H2O2 O2 SNGRaw flue gasCleaned Filter residues H2O Electricity Process unit Storage system Energy stream Material streamCO2 T = 200 - 600 °C p = 20 - 80 bar Waste heat of methanationWaste heat of melting tankWaste heat of flue gas coolingDesorption heat demandFlue gas treatment H2O Electrolysis 4 H2O → 4 H2 + 2 O2Methanation 4 H2 + CO2 → CH4 + 2 H2OCH4 + 2 O2 → CO2 + 2 H2OFlue gases CO2 Separationlean on CO2 CO2 from batchOxyfuel melting tank flue gas rich on CO2Figure 1: Simplified flow sheet of the integration of Power-to-Methane (PtM) into oxyfuel glass melting pro- cesses. technology readiness level (TRL) 9, but offers only limited part-load capability (30-100 %) due to thermal restrictions [13]. SOEC would offer the highest efficiency of all options (up to 95 % [14]), but requires a operating temperature level of 600-1000 °C [13]. These temperatures can be provided in an energy-efficient way by integrating waste heat from catalytic methanation. However, heating and cooling creates a thermal inertia of SOEC, which impedes rapid adaptation to fluctuations of renewable energy sources. In addition, the TRL of 5-6 is the lowest of all technical options described here. The most promising technology for the application proposed in this work is PEMEL, as it has the widest part-load capability (0-100 %) and can thus adapt to the fluctuations of renewable energy sources. With a slightly lower TRL of 7, PEM is behind AEL in terms of development and availability [13]. Nevertheless, PEM electrolysers are already available in power categories of multiple MW (e.g. Siemens Silyzer 300 [15]). Development potential is primarily focused on the industrial series production of this electrolysis technology, which is expected to result in a significant reduction of costs [11]. 2.2 Gas Storage The intermediate storage of H 2(storage system 1 in Figure 1) is necessary because there are currently technological restrictions on the flexibility of the down-stream catalytic methanation process (see section 2.3). In general, there are various options for H 2storage: • Physical storage technologies like compressed gas, liquid, or cryo-compressed H 2storage [16]; • adsorption technologies, relying on carbon based materials like multi-walled carbon nanotubes (MWC- NTs), metal organic frameworks (MOFs) or zeolites [16]; • absorption technologies, based on metal hydrates, like iron-oxide pellets [17]. • chemical H 2storage like liquid organic hydrogen carriers (LOHCs) [18], For the process described in this work, H 2storage in pressurized gas tanks is currently the most suitable option. It is described as technologically mature and therefore a comparatively cost-effective option, especially 3 in combination with battery storage systems [16]. Nevertheless, the application of innovative H 2storage technologies within the described PtM process, should be focused in further research. The high waste heat potential in glass melting processes offers an attractive option for an energy-efficient use of LOHC or iron-oxide based H 2storage technologies. Storage system 2 in Figure 1 is necessary to ensure a constant flow of O 2for a stable combustion process in the melting tank. Various options for O 2storage exist at industrial scale since a long time and can therefore be considered robust [19]. So far, the O 2required for oxyfuel combustion in the glass industry has been produced primarily by air separation using vacuum pressure swing adsorption (VPSA) processes [19]. This is done either directly at the industrial site, or by delivery and storage in compressed gas tanks (liquid O 2). In both cases, oxygen storage facilities are usually available on site. Methanation, melting furnace, flue gas conditioning and CO 2capture can be considered as stationary pro- cesses with only minor transient behavior. However, a constant educt gas flow is required for a stable thermo- cathalytic methanation process. Nevertheless, a constant educt gas flow is required for stable methanation. In order to compensate for the slight fluctuations in the product stream of the CO 2capture plant, a CO 2storage system may also be required. The storage of CO 2is a long-proven and available technology and can therefore be regarded as mature [20]. Due to the comparatively low storage capacity required, no further details of this storage system are discussed in this work. 2.3 Methanation Thermocatalytic and biological methanation are the main tech currently used in power to gas systems [13, 21]. Due to the higher TRL and technical performance, catalytic methanation is mainly used in large-scale PtM plants [11]. However, this methanation process requires constant operating conditions to achieve a high CH 4content in the synthetic natural gas (SNG) [13, 21]. For this purpose, the educt gases CO 2and H 2must be supplied in the ideal stoichiometric ratio of 1:4. In addition, pressure and temperature conditions must be constant to avoid damages on the catalysts through hot-spot formation. From a thermodynamic point of view, low temperatures and high pressure would be ideal for a high conversion rate. However, the exothermic nature of the dominant methanation reactions challenges the operation at thermodynamic optimum. For many commercially available fixed-bed reactors, operating temperatures of 300 °C and pressures of 20 bar have proven to be suitable operating conditions. The operating temperatures of established reactor concepts vary between 200-600 °C and a pressures range of 20-80 bar [13, 21]. In any case effective cooling of the reactor is required. In the PtM system described above, this waste heat could be used for SOEC or CO 2-desorption (see section 2.7). In addition to the technical properties mentioned above, the purity of CO 2is crucial for a stable methanation process. In order to investigate the effects of CO 2impurities on methanation, experiments investigating direct methanation of flue gases for different types of power plants have been performed. The flue gases investigated were obtained both from lignite-fired power plants with conventional combustion and from pilot plants with oxyfuel combustion. In both cases, direct methanation of the flue gases led to heavy damage on the catalyst, so that the experiments had to be stopped. However, oxyfuel flue gases could ensure stable methanation after purification and treatment. Within these investigations, mainly SO 2and halogen compounds (F, Cl, Br, At, TS) have been described as strong acting catalyst toxicants [22, 23]. 2.4 Oxyfuel Glass Melting Tank In oxyfuel combustion processes, a fuel like natural gas is burned in a pure oxygen atmosphere instead of ambient air. Oxyfuel combustion is more energy-efficient because nitrogen content in the furnace atmosphere is significantly reduced, thus lowering the thermal capacity of the exhaust gases. This also enables improved mass transport in the combustion chamber. In addition, a higher adiabatic flame temperature is achieved [24]. The higher overall efficiency of this combustion technology also enables a reduction in CO 2emissions compared to the regenerative processes that are primarily used at present. A major limitation of this technology has been the energy-intensive production of oxygen and the associated higher investment costs. Oxyfuel 4 Table 1: Critical parameters of the melting system investigated in [25]. Parameter Value Melting technology oxyfuel Operating temperature 1600 °C Thermal power 6.08 MW Fuel Natural Gas H Glass type borosilicate Nominal pull rate 40 t/d Operating hours 8760 h/year furnaces are established state-of-the-art despite these limitations, especially in the special glass production sector. Due to the much more energy-intensive glasses that are molten for this purpose, the higher efficiency of oxyfuel melting offers an economic advantage [6]. 2.5 Flue Gas Purification Systems Currently, mainly fabric filters and electrostatic precipitators are used for flue gas purification in the glass industry. These cleaning systems are primarily aimed for compliance with air pollutant limits set by country- specific regulations. The air pollutants relevant for the glass industry are mostly dust resulting from turbulence in the furnace, carbon monoxide (CO), sulfur and nitrogen oxides (SO x, NO x), as well as hydrochloric and fluoric acid (HCl, HF). While NO xand CO emissions are mainly caused by combustion, SO x-, HCl- and HF-emissions occur due to glass batch impurities. To remove these air pollutants from the flue gas mixture, sorbents such as lime (calcium hydroxide, Ca(OH) 2) are used, which then deposit as so-called filter dust. In order to prevent the filter systems from being damaged by excessively high gas temperatures, a heat exchanger is usually installed upstream to cool the uncleaned flue gases. Especially fabric filters are much more sensitive to high gas temperatures than electrostatic filters. In addition, cooling by ambient air leaks in the flue gas system is common. 2.6 Flue Gas Properties and Composition Detailed data on the flue gas composition and properties must be available to enable an evaluation of the CO 2 separation process. However, the composition of flue gases in the glass industry is highly variable. The main influencing factors are: batch composition, cullet fraction, glass type, melting technology and fuel, as well as flue gas treatment and cooling. Therefore, we conducted a literature screening for data from flue gas analyses of melting systems, suitable for the application of the PtM process presented in this paper. We identified the exhaust gas study by Roger et al. [25] was identified as the most promising source. In this study, flue gas analyses were performed on several special glass melting tanks. The aim was to investigate the potential for removal of boron compounds (namely boric acid (H 3BO3) and meta boric acid (HBO 2)) from the flue gases generated during the melting of the typical borosilicate glasses. Although these investigations do not fully meet the scope of this work, several oxyfuel melting systems were investigated that would be well suited for the use of a PtM system. The most promising system is the one with the data given in Table 1. A conventional soda-lime glass was assumed to be molten in this furnace in these studies, rather than borosilicate glass. As a result, boron compounds will be absent in the flue gas. The effects of H 3BO3- and HBO 2-contaminated flue gases on currently established absorbents such as MEA have not been investigated so far and are therefore unpredictable. Based on these assumptions, the flue gas properties and composition shown in Table 2 are used for the examination of the CO 2absorption plant in this work. Additional thermodynamic properties for the flue gases were calculated using the open source software Cantera with GriMech 3.0 reaction mechanism [26, 27]. The 5 properties of the clean gas lean on CO 2(3⃝) are resulting from the ASPEN Plus model, described in section 3.1. Trace substances were neglected in this modeling approach. Table 2: Flue gas composition and properties for examination of the CO 2absorption plant. Based on experi- mental results of [25]. Additional properties are calculated using [26, 27], and the developed ASPEN Plus model. STP = standard temperature and pressure conditions (0 °C; 1.0135 bar). Numbers of gas compositions refer to Figure 2. 1⃝ 2⃝ 3⃝ Properties Unit Raw gas Clean gas Clean gas rich on CO 2lean on CO 2 Temperature °C 470 226 26 Volume flow rate - dry m³/h 4780 6910 4290 - wet m³/h 5910 8020 4437 Density at STP kg/m³ 1.266 1.268 1.161 Spec. heat capacity ( cp) J/(kg ·K) 1219.2 1138.4 1022.5 Composition H2O Vol% 19.1 13.8 3.3 O2 Vol% 19.8 20.5 24.8 CO2 Vol% 11.3 7.0 0.8 N2 Vol% 49.8 58.7 71.1 Trace Substances Dust mg/m³ 1482 0.3 - HCl mg/m³ 2.6 0.2 - HF mg/m³ 72.7 0.5 - SO2 mg/m³ 19.0 8.1 - The volume fractions of CO 2in the flue gas are very low, while N 2and O 2are very high for an oxyfuel combustion system. As a result in this particular case, high amounts of ambient air are entering the flue gas system. The rise of O 2fractions from raw gas to clean gas may also be explained due to ambient air leaks in the system. 2.7 CO 2Separation Technologies for CO 2capture from combustion processes can be categorized as pre-combustion, post-com- bustion processes and processes that are directly connected to combustion (in this work further referred to as ”in-combustion” processes). Pre-combustion processes primarily involve the separation of CO 2from fuels such as lignite, hard coal or natural gas, in order to prevent CO 2emissions prior to combustion. The coal is first converted into a CO 2+ H2gas mixture. In the case of natural gas, the main component methane is converted to CO 2+ H 2by steam reforming. The CO 2is subsequently separated by absorption processes such as rectisol, selexol or purisol. Such processes are widely established in the petrochemical industry, for example, to avoid acidification of natural gas and to be able to offer a high-purity product [28]. The remaining gas mixture, consisting mainly of H2, can then be further used as fuel. In-combustion processes such as chemical looping or so-called oxyfuel processes are directly involved in the combustion process. In the chemical looping process, metal oxides are added to the combustion chamber in fluidized bed reactors in order to effect a redox reaction, binding the CO 2[29]. Oxyfuel processes are closely 6 related to the combustion process described above. When natural gas is burned in pure oxygen, the exhaust gas theoretically consists only of CO 2and water vapor. The H 2O can then be condensed by cooling and a highly pure CO 2stream would remain [28]. Both processes are unsuitable for the glass industry. Metal oxides in the combustion chamber would react with the glass batch as well, thus negatively affecting the glass quality. The separation of CO 2by cooling is unsuitable, since the exhaust gases in the glass industry also contain evaporation from the molten glass (Section 2.6). Accordingly, the CO 2product stream would be contaminated after the condensation of H 2O. Post-combustion processes concern the capture of CO 2from combustion flue gases. Membrane processes as well as adsorption and absorption processes are discussed for this purpose. Membrane processes provide for the separation of CO 2from the flue gas stream by means of molecular lattices. However, these mem- branes are currently not yet sufficiently temperature-stable and scalable for rapid implementation. Adsorption processes are currently mainly used for CO 2capture from ambient air, known as direct air capture [30]. How- ever, de- and adsorption cycles are still too slow for application in exhaust gas streams and are not available in sufficient capacity [28]. For CO 2capture from exhaust gases, absorption processes are currently being discussed as the main option [28]. In particular, the use of monoethanolamine (MEA) as an absorbent is widely investigated and can be considered a standard. MEA suffers from degradation, caused by carbamate polymerization, as well as oxidative and thermal degradation. Therefore, a switch to more stable and efficient solvents is widely discussed [31, 32]. Nevertheless, CO 2absorption processes can be considered the most promising technology for the glass industry, since they have been investigated for a long time and have already been tested in several worldwide projects for CO 2separation from flue gases. They were also proposed for applications in other industry sectors, such as the cement industry [33]. A integration of such a post-combustion CO 2absorption plant is preferably realized after the already existing flue gas purification systems. At this point, the exhaust gases are already cleaned of harmful residues such as dust and a large proportion of the SO 2emissions. Most of the remaining SO 2is removed in the flue gas cooler in order to prevent excessive solvent degeneration. A flow sheet for this integration is shown in Figure 2.a). 2.7.1 Process Description of Amine based CO 2Absorption CO2separation from flue gases using absorption process can be divided into two main process steps. First, CO2reacts with the lean aqueous amine solution (30 wt.-% MEA in this case), in the absorber column. The rich solvent, containing chemically bounded CO 2enters the stripper column by passing a lean/rich heat exchanger. This preheats the rich solution close to the stripper operating temperature of about 120 °C and consequently cools the lean solution. In the stripper column, the MEA solution is regenerated by an endothermic process. Low-pressure steam is supplied to the reboiler to maintain stable regeneration conditions as a result. To ensure an efficient desorption process, saturated steam at a temperature level of at least 140 °C, and a pressure of 3.6 bar is required. This thermal energy demand for desorption of CO 2is considered to be the main cost factor during long-term operation of amine based CO 2separation plants. In power plant technology, usually steam from intermediate stages of turbines is used for this purpose. This reduces electrical power output of the turbines and thus has a significant impact on increased electricity production cost of carbon capture power plants. However, the PtM concept introduced in this work, offers several options for waste heat utilization to meet the desorption heat demand in an energy efficient way. 2.7.2 Heat Supply for CO 2Desorption The ability to provide the necessary heat for desorption process is crucial for the economic viability of a post- combustion CO 2absorption plant. The use of waste heat from these processes for steam generation is evident, especially for heat-intensive processes at common temperature levels of the glass industry. Figure 2. a) shows the most promising waste heat sources of the PtM process for the glass industry, namely i) methanation, ii) melting process and iii) flue gas filter unit. These options which will be described in more detail below. 7 flue gasesWaste heat of flue gas coolingWaste heat of melting tank CO2 Separation rich on CO2Oxyfuel melting tankFilter Unit Dust CO SO2 NOx HCl HFCa(OH)2 CO2flue gasesCleaned Raw lean on CO2flue gasesCleanedheat demand Desorption (MEA)CO2 Absorbent SNG Glass batch O2 1 2 3 Process unit Energy flow Material flow Gas composition 13-Sorbent Waste heat of methanation MethanationCO2 H2(a) Flue gas coolerFlue gasBlower ReboilerHeating steam COL 1COL 4 Water coolerHEX 1PUM 1 AbsorberCOL 2Washing section lean amine solutionCleaned gas Rich amine pumpPUM 2Rich/lean exchangerHEX 3 rich amine solutionlean amine solutionlean amine coolerHEX 2 StripperCOL 5 Condenser Vapour LiquidCO2 product HEX 4 (b) Figure 2: a) Principal flow sheet of the integration of a CO 2absorption plant in the glass industry. b) Basic flow sheet of a CO 2absorption plant with an amine-based solvent. It should be mentioned that waste heat is also generated during PEM electrolysis. Depending on the per- formance and design of PEM electrolysis systems, either water or air cooling is used. The temperature level of this waste heat is at approx. 80 °C [14]. The utilization of this waste heat potential for steam generation is limited as a result. Further processing, for example by means of high-temperature heat pumps would be required to generate steam on the required temperature and pressure levels of CO 2desorption. Due to these limitations, the waste heat option of PEM electrolysis is not considered in detail in this work. Waste Heat of Thermochemical Methanation For thermochemical methanation processes, continuous cooling of reactors is required to prevent degeneration of established catalyst materials. The exothermic natue of the dominant reaction processes would otherwise lead to catalyst particle sintering. Commercially established tubular bulb reactors for methanation processes are usually operated at a temperature level of 300 °C and a pressure of 20 bar. Usually water cooling is a common way to maintain constant operating conditions in these reactors, generating steam at a temperature of approx. 260 °C and a pressure of 45 bar [34]. A Sankey diagram of tubular bulb reactors for methanation and their cooling process is shown in Figure 3. Waste Heat of Melting Tank Due to the melting temperatures of up to 1600 °C, and an efficiency of about 42 % of oxyfuel furnaces [6], a high waste heat potential occurs in the direct furnace environment. However, the use of this heat reservoir is limited by a number of restrictions. For example, the installation of heat 8 Figure 3: Sankey diagram of a tubular bundle methanation reactor [34]. exchangers close to the furnace wall is restricted, since permanent access for industrial maintenance and servicing is necessary. The remaining option is to use the heated ambient air close to the tank environment. However, the low temperature level of the ambient air (below 100 °C) is an obstacle for the generation of process steam at the temperature and pressure level required for CO 2desorption [35]. Waste Heat of Flue Gas Treatment In many cases, the waste gases from the furnaces are still at too high a temperature level to be directly passed on to a filter system. In particular, the textiles in cloth filter systems are sensitive to excessively high temperatures. Heat exchangers are therefore installed in many flue gas cleaning systems to cool flue gases to a suitable level for filter systems. The heat recovered in this way is used, for example, to supply heat to office and administration buildings at glass industry sites [36]. In some cases reheating of the waste gases is also necessary to ensure sufficient flow conditions for the removal of the waste gases through pipes and stacks [35, 37]. 3 Materials and Methods 3.1 Simulation Approach The technical and economic analysis of CO 2capture from flue gases of the glass industry requires design and sizing of the plant components. For this work, the process simulation tool Aspen Plus was used for component design. Numerous other studies found this approach to be useful and validated [38, 39]. The most important equipment of the CO 2capture plant, namely the absorber and stripper column (see Figure 2) were designed in detail, using the simulation results of the Aspen Plus model (v12). The flowsheet of the Aspen Plus model is depicted in Figure 4. The columns are modeled in detail using a rate-based approach. The other components as the compressor, the pump or the heat exchanger are modeled in a simplified way. The ELECNRTL property method in Aspen is used to calculate media data using the eNRTL model for the liquid phase and the Redlich-Kwong equation of state for the vapour phase. The dissociation reactions are considered by defining equilibria reactions and whereas the kinetics are considered for all reactions including CO 2as a reactant using built-in power laws in Aspen Plus. Because of the fast reaction of the CO 2, the liquid film is discretized in five parts. As structured packings, the Melapak 250Y from Sulzer are used in all columns. The mass transfer coefficients, interfacial areas and column hold- up are calculated using the correlation of Bravo et al. [40, 41]. The heat transfer coefficients are determined according to the theory of Chilton and Coburn [42]. In order to calculate the column diameters the fractional approach to maximum capacity is used in Aspen Plus [43]. A factor of 0.65 is chosen. A constant capture rate of 90 % is set using the lean solvent mass flow. The heat flow into the reboiler in the stripper column is varied in such a way that the CO 2mass balance is met. The lean solvent loading is varied to ensure an optimal operation point of the plant and minimum heat demand in the reboiler. The packed column height was set according to [38], and the relevant parameters of the plant are listed in Table 3. In order to calculate the equipment costs, a comparison with an identical system is required (see section 3.2). A simulation model was developed for the CO 2separation plant with the exhaust conditions 2⃝shown in 9 ABSORBER STRIPPER HEATER PUMP COOLER WASHER CMP LEANIN FLUEGAS GASOUT RICHOUT RICHIN CO2OUT LEANOUT TOHEATER FGCMPIN LIQOUT LIQIN WOUT LIQWASH LIQWOUT FGABSINFigure 4: Main flow sheet overview of Aspen Plus Model used for the simulation Table 3: Critical input design parameters of amine scrubbing plant parameter value Solvent temperature at absorber inlet 35 °C MEA-concentration 30 w.-% Desorber pressure 2 bar Absorber flue gas outlet pressure 1 bar CO2product temperature 25 °C Washing fluid temperature 25 °C Desorber solvent inlet temperature 105 °C carbon capture rate (set point) 90 % Table 2, as well as a model for the reference CO 2absorption plant described in [39, 44]. 3.2 Economic Analysis The costs of the plant components were calculated using the “estimating equipment cost by scaling” method, described in [45]. For this purpose, the costs of closely related reference plant equipment kris compared to equipment cost of the proposed plant keby an exponential factor m: ke=krG Grm (1) The basis of characteristic component parameters of the reference equipment Grsuch are for example column volume, pump mass flow rate or heat transfer area. These parameters are in relation to the one of the proposed plant G. Various degression exponents m < 1for the respective components allow the cost comparison for each equipment type. Other components of the investment costs, such as installation, instrumentation and control, piping or elec- tricity supply, can subsequently be calculated as a proportion of total equipment costs, as proposed by [45] and already used for similar approaches for cost calculations of CO 2absorption plants [44]. The well established equivalent annual cost (EAC) method was used to calculate costs of owning and operating assets over their entire life time. 10 The “ProcessNet - Chemical Plant Price Index for Germany” was used to take inflation-related price devel- opment since 2007 as well as increased material, distribution and logistics costs into account. According to this index, prices have risen by 21.5 % since 2008. Due to the impact of the global COVID-19 pandemic, the effects of production cutbacks and limited logistics capacities are particularly evident from 2020 onward. By 2019, the increase would have been only 16 % (see Figure 5) [46]. 4 Results and Discussion 4.1 Simulation Results For CO 2absorption plants, the absorber and stripper column, as well as the gas blower and the cooling water pump for the pre-scrubber are the main cost factors with a share of more than 80 % [44]. The relevant comparative parameters for equation 1 were considered during simulations according to Table 6. Columns 4 and 6 in Table 6 show the simulation results for the CO 2absorption plant considered in this work, as well as the re-evaluation of the reference plant, considered in [39, 44]. The most important results are shown in Table 4: Table 4: Major simulation results for the design of the amine scrubbing plant parameter value Absorber diameter 0.95 m Desorber diameter 0.5 m Scrubber diameter 0.95 m Solvent flow rate 2.28 kg/s Heat demand in stripper 556.5 kW Specific heat demand in stripper 3.354 MJ/kg CO 2 CO2-purity 99.3 % CO2product mass flow 0.167 kg/s Washing fluid temperature 25 °C Desorber solvent inlet temperature 105 °C carbon capture rate (Sollwert) 90 % 4.2 Cost Calculation Results Cost calculations for the CO 2absorption plant were determined considering the following steps: i) calculation of equipment cost (section 4.2.1), ii) calculation of capital expenditures (CAPEX, section 4.2.2)), iii) calculation of operational expenditures (OPEX, section 4.2.2) and iv) the calculation of resulting CO 2avoidance cost (section 4.4). Basic assumptions for operating parameters of the CO 2absorption plants are shown in Table 5. 4.2.1 Equipment Cost Table 6 gives an overview of the equipment cost for both of the modeled CO 2absorption plants. Details concerning the reference plant are given in [44]. The plant for this work was designed for the same CO 2 capture rate of 90 % and stripper operating conditions of 2 bar and 130 °C. The most expensive equipment is the absorber column, which accounts for approx. 65 % of total equipment costs. The second most important contribution to equipment costs is related to the stripper column, responsible for another 20 %. As a result, 85 % of the total costs can be attributed to the most important columns in the process. The costs for cold water pump, blower and DC water cooler were also calculated in detail. These equipment parts contribute another 10 % of total costs. The remaining 10 % can be attributed to the lean/rich 11 Table 5: Operation parameter and economic assumptions for cost calculations of CO 2absorption plant. All parameters are based on [44], except others are named.*Results of own simulations. Parameter Value Specific desorption heat demand 928 kWh/t CO 2 Reboiler heat power 915 kW Life time 20 years Interest rate 5.0 % p.a. Operating hours 8688 h/year Maintenance cost 4.0 % of FCI MEA price [44] 1000 C/t MEA degradation rate 1.5 kg/t CO 2 Cooling water make up 1.0 m³/ GJ thermal Cooling water cost 0.20 C/m³ Water cooler duty 640 kW CO2product temperature 25 °C CO2product pressure 20 bar Electricity demand for CO 2compression [47] 72 kWh/t CO 2 CO2emissions before*9601 t CO 2/year CO2separation rate*90 % CO2avoided*8641 t CO 2/year CO2emissions after*960 t CO 2/year heat exchanger, amine pumps, reboiler, and other minor important equipment parts. The distribution of these expenses is in line with equipment cost calculations of [44]. The costs for CO 2compression after separation is calculated following the approach given by [47] for CO 2 compression for carbon capture and storage (CCS) processes. Lower pressure levels are required for metha- nation than for underground storage of CO 2. [47] suggest a four stage compression process for a CO 2product pressure of 120 bar. We assumed a one stage compression process to 20 bar is assumed. This pressure level is sufficient for established large scale methanation reactors like [49]. Consequently, equipment costs for CO2compression in the PtM concept are much lower than for CCS processes [44] 4.2.2 Capital Expenditures Further direct and indirect costs for calculation total CAPEX of the CO 2separation plant were estimated as a factor of the overall equipment cost, as suggested in [45]. Table 7 shows the parameters taken into account and their influence on the total CAPEX. Due to the comparably small size of the CO 2capture plant, the lower limits of the ranges proposed in [45] were selected. The OSBL costs for buildings, yard improvements, and service facilities are assumed to be lower than recommended by [45] due to low column volumes and compactness of the plant. As a result, a container solution for plant housing is a viable alternative. Considering this option, OSBL costs of over 100,000 C building infrastructure is justified. The price index correction, according to Process Net [46], since 2007 has a significant impact of 21 % on total CAPEX. The extent to which prices develop according to current influences such as the COVID-19 pandemic and recent conflicts should be further examined. These effects have not yet been taken into account in the current values of [46]. CAPEX of plants for CO 2capture from flue gases of power plants are far above the values calculated in this study. [44] and [50] propose C 147 million and C 179 million, respectively. By considering column 12 Table 6: Equipment cost calculation. Indicatiors (Ind.) refer to Figure 2.b for easier identification of equipment. *Remaining heat exchanger and pumps with less significant influence on total component costs. Ind. Equipment Type Exponent Reference plant Own plant m G r krin MC G k einC COL 2 Absorber Column 0.60 [48] 12560 m³ 10.94 39.25 m³ 340,000 COL 4 Stripper Column 0.60 [48] 3391 m³ 3.43 8.75 m³ 110,000 – Cold water pump Pump, reciprocating 0.34 [45] 1.21 m³/s 2.04 0.0015 m³/s 21,000 PUM 1 Blower Blower, centrifugal 0.59 [45] 485.44 m³/s 3.10 2.236 m³/s 13,000 COL 1 Flue gas cooler Column 0.60 [45] 1570 m³ 0.54 4.75 m³ 17,000 – Others* mixed - - 3.89 - 20,000 Total equipment cost: 23,94 521,000 CO 2compressors Compressor, rotary, 0.79 [45] 31.49 m³/s 31.73 0.07761 m³/s 69,000 single stage geometry and main equipment dimensions, the plant presented here is about 400 times smaller. This factor is also outside the limit proposed by [45] for the applicability of the cost comparison method. The determined cost frameworks can be considered valid, by comparing specific CAPEX to other literature sources. [33] investigated CO 2capture for carbon-intensive industrial processes and reported specific CAPEX of 160 C/t CO2and year for amine based post-combustion CO 2capture in cement industry processes. The resulting specific CAPEX of this work are 188 C/t CO 2and year. Considering the price increases since 2012 (approx. 16 % according to [46], see also Figure 5), the deviation of 15 %, calculated in this work seems justified. Different boundary conditions and assumptions of this work can influence the costs of the implementation of a CO2capture plant. 4.2.3 Operational Expenditures The total operational expenditures (OPEX) include fixed charges for insurance and taxes, direct production costs, general expenses and plant overhead costs. The total annual OPEX of the novel CO 2separation plant are shown in Table 8. The yearly OPEX were calculated to be approx. 146,500 C. The main OPEX factors are operating labor costs, contributing about 27 % to total OPEX. We assumed a 0.1 job per shift for operation of the CO 2sepa- ration plant. In the glass industry, a three-shift operation of 8 hours each for 365 days per year is common to maintain labor supply. The assumption that 10 % of the working time of a maintenance employee per shift is spent on maintenance and servicing of the CO 2separation plant thus corresponds to a total time expenditure of approx. 2.5 h/day or 912.5 h/year. Others studies such as [44] assumed a two job per shift staff expen- diture for plant operation. However, as the separation plant discussed in this work is much smaller, such a reduction in operating labor expenses seems justified. In addition, operating labor costs influence supervision and support labor costs, laboratory charges and plant overhead costs, which result in operating labor costs contributing to approximately 60 % of total OPEX. MEA makeup and maintenance costs contribute about 18 % of total OPEX. Other works report this factors to have an influence of aprox. 32 % [44, 50], but consider a smaller plant size. Due to the comparatively small amounts of CO 2involved, correspondingly less MEA makeup and maintenance is required. We found OPEX to be 9 % of total CAPEX. In literature, a OPEX range of 7-12 % of total CAPEX for a post combustion MEA separation plant in the cement industry was considered [33]. The calculated OPEX is consistent with other literature sources, indicating potential for further optimization. 13 Table 7: Total capital investment (CAPEX) for MEA CO 2separation plant. Range [45] Used Cost % % C Direct cost Inside battery limit (ISBL) cost Total equipment cost (EC) 100 521,000 Installation 25-55 25 130,250 Instrumentation and control 8-50 8 41,680 Pipes 20-80 20 104,200 Electrical equipment 15-30 15 78,150 Total ISBL: 875,250 Outside battery limit (OSBL) cost Building and building services 10-80 5 26,050 Y ard improvements 10-20 5 26,050 Service facilities 30-80 10 52,100 Total OSBL: 104,200 Total direct cost: 875,250 Indirect cost Engineering 10 10 52,100 Construction expenses 10 10 52,100 Contractor’s fee 0,5 0,5 2,605 Contingency 17 17 88,570 Total indirect cost: 106,805 CO2compressor equipment 69,000 Fixed capital investment (FCI): 1,155,285 Working investment 12-28 12 136,634 Start-up cost and MEA cost 8-10 8 92,423 Price index correction (2007-2022); % of FCI [46] 21,5 248.386 Total capital investment (CAPEX): 1,634,278 4.3 Cost for Desorption Heat Generation The supply of the necessary heat for desorption is crucial for the economic viability of a post-combustion CO 2 absorption plant (section 2.7.2). As shown in Table 5, the heat demand for the designed CO 2separation process is 928 kW, at a steam temperature and pressure of 150°C and 2 bar. In particular, the methanation process as well as the flue gas filtration systems are the most promising options regarding the use of waste heat for steam generation. 4.3.1 Methanation Waste Heat Utilization Considering the thermal power demand of the melting tank described in Table 1, a 7.60 MW methanation reactor is required to ensure a constant heat supply. This reactor can provide high pressure and temperature steam (HP-steam) and low pressure and temperature steam (LP-steam, see Figure 3). Table 9 shows the power distribution of a tubular bulb reactor, based on vendor information [34]. It is evident that the HP steam of the methanation reactor can already cover approx. 120 % of the heat demand for CO 2desorption. According 14 Table 8: Total operating expenditures (OPEX) for MEA CO 2separation plant. Range [45] Used Cost % % C/year Fixed Charge Local taxes 1.0-4.0 % of FCI 1.0 % 11,533 Insurance 0,5-1,0 % of FCI 0.5 % 5,776 17,329 Direct production cost Cooling water 4,003 MEA Makeup 1.5 kg/t CO 2 14,401 Maintenance (M) 1.0-10 % of FCI 1.0 % 11,553 Operating labor (OL) 0.1 job/shift 45 C/h 39,096 Supervision and support labor (S) 30 % of total labor cost 30 % 11,729 Operating supplies 15 % of maintenance 15 % 1,733 Laboratory charges 10-20 % of operating labor 10 % 3,910 86,425 Plant overhead cost 50-70 % of M + OL + S 50 % 31,189 General expenses Administrative cost 15-20 % of OL 15% 5,864 Distribution and marketing 2-20 % of OPEX 2 % 2,816 Research and Development cost 2-20 % of OPEX 2 % 2,816 Total operating expenditures (OPEX): 146,440 15 to vendor information, a specific CAPEX of 40 C/kW thermal can be expected for the cooling equipment of the methanation reactor. Based on the total steam power of 1.22 MW, a total CAPEX of approx. 49.000 C can be considered for cooling equipment of the methanation reactor in the considered PtM system of this work. Thus, assuming a lifetime of 20 years and an interest rate of 5 %, specific costs of 0.0004 C/kWh thermal are achieved. Table 9: Waste heat potentials of methanation reactor equipment. HP = high pressure and temperature, LP = low pressure and temperature. Reference point Power MW Thermal power heating value of SNG 6.08 HP-steam 45 bar, 260 °C 1.09 LP-steam 30 bar, 240 °C 0.12 Power loss heat losses 0.30 Methanation reactor combined 7.60 4.3.2 Flue Gas Treatment Waste Heat Utilization In addition to methanation, waste heat recovery from flue gas cleaning for desorption heat demand is a promis- ing option. However, considering the properties shown in Table 2, the enthalpy content in the investigated flue gas is about 850 kW for a ∆Tof 247 °C. This is not sufficient to provide the required thermal power of the reboiler of 915 kW. The flue gas needs to be cooled to approx. 130 °C to ensure sufficient heat output for adequate steam generation. At this ∆Tof 340 °C an enthalpy of 1171 kW is contained in the flue gases. This heat content is sufficient to cover the reboiler’s heat demand as well as to compensate for heat losses through a flue gas heat exchanger. Based on vendor information and calculations, a 5 bundle plain tube heat exchanger is sufficient to generate steam at the required desorption conditions. Depending on the exhaust gas properties, conventional steel or, in the case of high exhaust gas corrosivity, stainless steel would be suitable as a construction material for this heat exchanger. A cost estimate for these material options is around 40,000 C for conventional steel, and 90,000 C for stainless steel respectively [36]. Assuming a lifetime of 20 years and an interest rate of 5 %, specific costs of 0.0007 C/kWh thermal for common steel and 0,001 C/kWh thermal for stainless steel are achieved. The lower outlet temperature of the flue gases from the filter system would also affect the CO 2separation system. As a result, the capacity of the flue gas cooler could be reduced, which in turn would result in lower investment costs. This would also influence flow characteristics in the flue gas channel. Consequently, existing flue gas systems may no longer be able to achieve the required flow capacities and may require re-planning and design. 4.4 Cost of CO 2Avoidance With a lifetime of 20 years and an interest rate of 5 %, the annual depreciation or EAC costs of the CO 2 separation plant designed in this work are 131,175 C/year. As shown in Table 8, there are additional OPEX of 146,440 C/year. According to [47], a specific electrical power of 73 kWh/ t CO 2is required to compress the captured CO 2to a pressure of 20 bar. Specific electricity costs are taken from [6] where 0.12 C/kWh were assumed for glass industry companies. The three cost factors EAC of CAPEX, OPEX and electricity demand for CO 2compression add up to an annual cost of 353,348 C for the CO 2separation plant designed in this work. 16 The heat demand for desorption in year-round operation is 8017 MWh/year, or 928 kWh/t CO 2. The de- scribed CO 2capture plant can avoid 8641 t CO 2/year (Table 5). The three options for waste heat utilization, described in section 4.3, were used to investigate their influence on CO 2avoidance costs. The results are shown in Table 10. The CO 2avoidance costs are 41,31 C/t CO 2for waste heat utilization of the methanation process. For the flue gas heat exchanger option, slightly higher costs of 41,61 C/t CO 2, or 41,92 C/t CO 2 are achieved. These heat exchangers generate a lower heat output, and the use of stainless steel results in increased material costs. Despite the marginal increased cost, they illustrate the potential for cost-effective waste heat utilization for absorption based CO 2capture in the glass industry. Table 10: Calculation of CO 2avoidance costs for different waste heat utilization options. EAC = equivalent annual costs of CAPEX. Specifc cost Total cost C/year EAC of CAPEX 189 C/(t CO 2·year) 131,175 OPEX 146,440 Electricity for CO 2compression 0.12 C/kWh 8.60 C/t CO 2 82,521 Fixed costs: 353,348 Heat supply cost options 1. Methanation reactor 0.0004 C/kWh 0.37 C/t CO 2 3,264 Total annual costs: 356,912 CO2avoidance cost: 41,31 C/t CO 2 2. Flue gas heat exchanger 0.0007 C/kWh 0.65 C/t CO 2 6,235 conventional steel Total annual costs: 359,584 CO2avoidance cost: 41,61 C/t CO 2 3. Flue gas heat exchanger 0.0010 C/kWh 0.93 C/t CO 2 8,908 stainless steel Total annual costs: 362,256 CO2avoidance cost: 41,92 C/t CO 2 Other studies reported CO 2separation costs from industrial waste gases to range from 25 to 135 C/t CO 2 [33, 51]. But these reported values depend very much on the used CO 2capture technology. Comparable process designs in the cement industry with post-combustion CO2 capture using MEA assume short to mid- term CO 2avoidance costs to be about 65 C/t CO 2at low cost for external steam import [33]. In studies investigating CO 2capture from industrial point sources for methanation, 50 C/t CO 2were found [51]. 4.5 Optimization Options The CO 2avoidance costs of 41-42 C/t CO 2indicate the promising potential for low-cost and energy-efficient integration options for the post combustion CO 2separation within the introduced PtM process for the glass industry. In addition, there are further options for optimizing the CO 2capture process that have not yet been investigated: In this work, a 30 wt.-% aqueous solution of MEA was assumed. Other studies have shown that increasing the MEA content in the aqueous solution up to 40 wt.-% can further reduce the CO 2avoidance costs. This can be achieved by reducing the energy requirement for desorption and lowering investment cost for the plant equipment due to reduced liquid flow rates [44]. MEA is considered to be the standard solvent for post combustion CO 2absorption. In addition, other 17 absorbents such as secondary or tertiary amines, as well as various mixtures, are described to have improved performance compared to MEA. A mixture of the tertiary amine methyl-diethanolamine (MDEA) and piperazine (PZ) is described to have increased resistance to thermal and oxidative degradation, higher CO 2loading capacity and a reduced heat demand of 2.6 GJ/t CO 2vs. the 3.4 GJ/t CO 2of MEA [52, 53]. In addition to solvent replacement and changing solution properties, an advanced process design could positively influence the costs of CO 2separation. Improved designs for the rich/lean heat exchanger [54], vapour recompression and split-stream processes [55] have been investigated and showed promising potential for enhances CO 2removal capacity or reduced energy demand. Combinations of these improved process designs and careful solvent selection could further optimize costs and performance of absorption based CO 2separation from flue gases of glass melting processes. 5 Conclusion An innovative energy concept towards carbon free glass melting processes by effectively integrating a PtM process into an oxyfuel glas melting plant is introduced and discussed. The energy concept utilizes of the well-established infrastructure and know-how related to natural gas based oxyfuel glass melting plants. The integration of the PtM process seems to be a very promising option for the decarbonization of glass melting processes. The novelty of the main process components is described and recommendations for suitable technology options are provided in these studies. The required capture of CO 2from glass industry flue gases has been identified as the process with the highest research expenditure. Post-combustion CO 2absorption capture processes, which have also been applied in power plant technology, were identified as the most promising option to meet this challenge. A process simulation of this technology allowed a detailed design of the most important plant components and their cost evaluation. The following conclusions can be drawn for CO 2absorption from flue gases of the glass industry: • Amine based CO 2capture processes allow a reduction of 90 % of CO 2emissions of combustion based glass melting processes. The designed CO 2absorption plant for the glass industry is approx. 400 times smaller than comparable concepts for power plants, due to the much lower volumes of flue gas. • This enables cost reductions in the required building infrastructure, as well as in staff costs for mainte- nance and monitoring. At the same time, cost-reducing scaling effects for large plants have a negative impact on the small plant described here. • Various options for waste heat utilization are available within the introduced PtM process for oxyfuel glass melting, which offer cost-effective options for desorption heat supply. Especially the methanation process can be seen as the most suitable option, as it provides steam at sufficient temperature, pressure and volume, with no impact on existing glass industry infrastructure. • The low operating and heat supply costs result in low CO 2avoidance costs of 41 C/t CO 2. These could be further reduced by described optimization approaches, like changes in solvent loading, the use of advanced solvents and/or advanced CO 2absorption processes. These promising results for CO 2capture also motivate further research on the entire PtM process presented. Initially, further simulative and techno-economic analyses should be carried out in order to assess the technical and economic parameters. Subsequently, both lab-scale plants and step-by-step integration into real glass melting processes can be rapidly implemented. Acknowledgement Dr. Christian Schuhbauer of MAN Energy Solutions SE, Germany, should be thanked for his valuable advice on methanation reactor designs, energetic analysis and cost evaluations. The authors would also like to thank 18 Bernhard Lattner and Rudolf Steiner of Kelvion Ltd., Austria, for their contribution on flue gas heat exchanger design and cost evaluation. Furthermore, Prof. Dr. Luigi Vanfretti as well as Meaghan Podalski from the ALSET Lab of the Rensselaer Polytechnic Institute (RPI), Troy, NY , USA should be thanked. They supported the work on this paper during a visiting scholarship, funded by German Academic Exchange Service (DAAD). This research was further funded by a scholarship from the Nagelschneider Foundation, Munich. The authors would like to acknowledge the financial support received from the Nagelschneider Foundation. 19