*Article* **Integration of Renewable Hydrogen Production in Steelworks Off-Gases for the Synthesis of Methanol and Methane**

**Michael Bampaou 1,2,\*, Kyriakos Panopoulos 2,\*, Panos Seferlis 1,2, Spyridon Voutetakis 1, Ismael Matino 3, Alice Petrucciani 3, Antonella Zaccara 3, Valentina Colla 3, Stefano Dettori 3, Teresa Annunziata Branca <sup>3</sup> and Vincenzo Iannino <sup>3</sup>**


**Abstract:** The steel industry is among the highest carbon-emitting industrial sectors. Since the steel production process is already exhaustively optimized, alternative routes are sought in order to increase carbon efficiency and reduce these emissions. During steel production, three main carboncontaining off-gases are generated: blast furnace gas, coke oven gas and basic oxygen furnace gas. In the present work, the addition of renewable hydrogen by electrolysis to those steelworks off-gases is studied for the production of methane and methanol. Different case scenarios are investigated using AspenPlusTM flowsheet simulations, which differ on the end-product, the feedstock flowrates and on the production of power. Each case study is evaluated in terms of hydrogen and electrolysis requirements, carbon conversion, hydrogen consumption, and product yields. The findings of this study showed that the electrolysis requirements surpass the energy content of the steelwork's feedstock. However, for the methanol synthesis cases, substantial improvements can be achieved if recycling a significant amount of the residual hydrogen.

**Keywords:** blast furnace gas; coke oven gas; basic oxygen furnace gas; methanation; methanol synthesis; aspen plus; gas cleaning; hydrogen; steelworks sustainability

## **1. Introduction**

The iron and steel industry is among the industrial sectors with the highest production volumes, having indispensable end-products for modern society [1]. The European steel industry, in particular, is a world leader in steel production accounting for approximately 16% of the world production (8.5% belongs to the European Union countries), coming second only to China. In market and economic terms, in 2019 it generated 140 bn € of gross added value and employed around 2.67 million people [2]. Steelworks, however, are one the most energy- and carbon-intensive industries in the world, accounting for 27% of the total industrial CO2 emissions and 4–5% of the total anthropogenic CO2 emissions [3]. Since world steel production is expected to rise in the following years, CO2 and carbon emissions will increase accordingly, if no proper countermeasures are adopted [1].

During the primary steel production route, carbonaceous off-gases are generated during the main production steps of: (1) conversion of coal to coke in the coke oven, (2) pig iron production in the blast furnace and (3) processing of pig iron to steel in the basic oxygen furnace [1]. Since the usage of fossil fuels (usually coal and natural gas) as reducing agents in the blast furnace is intensely optimized [4], alternative ways are investigated for the reduction in those emissions. Generally, a common way to avoid the flaring of

**Citation:** Bampaou, M.; Panopoulos, K.; Seferlis, P.; Voutetakis, S.; Matino, I.; Petrucciani, A.; Zaccara, A.; Colla, V.; Dettori, S.; Annunziata Branca, T.; et al. Integration of Renewable Hydrogen Production in Steelworks Off-Gases for the Synthesis of Methanol and Methane. *Energies* **2021**, *14*, 2904. https://doi.org/10.3390/ en14102904

Academic Editors: Markus Lehner and Dmitri A. Bulushev

Received: 5 March 2021 Accepted: 14 May 2021 Published: 18 May 2021

**Publisher's Note:** MDPI stays neutral with regard to jurisdictional claims in published maps and institutional affiliations.

**Copyright:** © 2021 by the authors. Licensee MDPI, Basel, Switzerland. This article is an open access article distributed under the terms and conditions of the Creative Commons Attribution (CC BY) license (https:// creativecommons.org/licenses/by/ 4.0/).

steelmaking off-gases is their use as internal energy sources both for heating and power production. As a consequence, a reduction in natural gas and external produced electricity use is obtained, resulting in a decrease in emissions, primary resources consumption, and operating costs. Recent works focused on the optimization of the management of steelworks off-gases networks using a decision support system [5] including machine learning-based forecasting models [6–8] and advanced optimization strategies [9,10].

An alternative/complementary way of utilizing those gases without deviating from the already established steel production route is their conversion to added-value chemicals. The proposed utilization strategy involves the use of the carbonaceous feedstocks for the production of methane and methanol (MeOH) through the addition of renewable hydrogen by proton exchange membrane (PEM) electrolysis. Apart from the environmental perspective, target is to partially replace the fossil fuel demands of the steel plant and/or to generate revenue by utilizing a by-product stream. Methanol has already broad commercial uses, as chemical intermediate and fuel [11], whereas methane apart from its commercial value, can be used within the steel plant for power production and/or reused as reducing agent in the BF process [12]. The proposed strategy, however, has to surpass or match the benefits obtained through the conventional off-gases exploitation strategy (i.e., heating and power production) from an energetic, economic and environmental perspective.

The three mentioned steelworks off-gases (Blast Furnace Gas: BFG, Coke Oven Gas: COG, Basic Oxygen Furnace Gas: BOFG) are commonly stored in dedicated gasholders, that act as buffers. These off-gases contain more or less the same compounds but at different proportions: the most common are CO2, CO, H2, and N2. Small amounts of impurities are also contained; however, they do not pose environmental threats when combusted in their traditional use within the plant. However, when advanced catalytic processes are pursued using these gases as feedstock, then, further gas cleaning steps are required to avoid catalyst poisoning. The present work considers an already existing gas cleaning setup prior to the gas holder short storage as the starting point for the process formulation. In addition, further gas cleaning steps are proposed upstream the catalytic processes, considering a possible presence of residual impurities in the off-gases, before entering the catalytic syntheses units.

The scope of this work is to study the integration of renewable hydrogen into steelworks off-gases for the efficient production of methane and methanol and to exploit the largest amounts of steelworks off-gases as carbon sources. This is a novelty in respect to past works [13,14] that exploit only limited amounts of these off-gases and focus mainly on the exploitation of the COG as feedstock for the synthesis reactors (as it is or mixed with other off-gases, due to its high hydrogen content). This study has been conducted using flowsheet simulations in AspenPlusTM. The key points of this work can be summarized as follows:


This work is organized as follows: Section 2 describes the main features and characteristics of the considered off-gases; Section 3 illustrates the investigated process and the

comprising sub-systems; and Section 4 presents and discusses the obtained results. Finally, Section 5 provides the conclusions of this work and hints for future work.

## **2. Steelworks Off-Gases**

Table 1 depicts the total volumetric amounts and the mean composition of the steelworks off-gases, for a steel plant producing 6 MT steel per year [15].


**Table 1.** Mean composition of steelworks off-gases.

As shown in Table 1 from an overall perspective, the component that prevails through the three gases is nitrogen. The inert nature of nitrogen lowers the partial pressures of the reactants and raises the volume of the feed gases. Thus, it increases the capital expenses and the costs associated to compression and could lead to accumulation within a recycling loop. The contained CO and CO2 can be used as feedstock for the production of chemicals (e.g., CH4 and/or CH3OH). The reactivity of CO is always higher than the one of CO2 for a considered chemical, resulting, thus, in higher activation energies for the CO2 conversion [15]. The insufficient amount of H2 contained in the off-gases, dictates the addition of additional hydrogen for the synthesis. In order to increase the carbon efficiency of the process, attention should be paid on the addition of renewable hydrogen instead of fossil based. In addition, it is assumed that after cleaning, the off-gases are water saturated. This water content should be removed prior to compression, in order to: (i) avoid condensation that could damage the compressors [15], and (ii) avoid the promotion of the Water Gas Shift (WGS) reaction that could lead to the consumption of CO for the formation of additional CO2 [17]. Finally, the off-gases also contain small amounts of oxygen, which need to be removed for safety reasons prior to the conduction of adsorption processes, such as pressure swing adsorption [18].

From a particular point of view, the BFG contains large quantities of nitrogen due to the use of hot air as oxidant within the furnace and low amount of H2 [6]; enrichment is required for its use as feedstock for methane and methanol production. The coke oven gas is generated in the coking plant during the heating of coal to produce coke. In contrast to the BFG, it contains large amounts of hydrogen and can be mixed with the other gases to reduce the required amounts of additional hydrogen by electrolysis. In addition, it can be easily valorized within the plant as fuel or feedstock for producing chemicals, due to the contained hydrogen and methane [19]. The BOFG is generated in the basic oxygen furnace, where oxygen is injected to oxidize part of the carbon in the pig iron produced from the blast furnace; it contains predominantly CO [15].

The three steelworks off-gases are generally used internally for heating and electricity production purposes. For instance, the COG is used for firing coke ovens, as heat input for rolling mills and to produce energy at the power plant [20]. The blast furnace gas serves also as a fuel for firing the coke ovens, the hot blast stoves heating the wind to be injected into the blast furnace and the power plant [21], whereas the basic oxygen furnace gas, apart from power applications, can also be used for upgrading the heating value of the BFG in a gas mixing station [16]. Gasholders are used for storing the surplus of

those gases. However, in some cases the gasholders capacities are not sufficient to contain the generated quantities and as a consequence, the excess off-gases are flared. In other cases, the gases do not satisfy the internal requirements and natural gas is purchased. An optimized off-gas distribution management can improve the efficiency [5], as well as the consideration of an alternative use, such as for methane and methanol production. Regarding the available amounts for CH4 and MeOH production, it is assumed that 50% of the total generated amount is available, after the rest being utilized in internal applications within the plant [16,20].

## **3. Process Description**

In this section, the outline of the five case studies is described. These case studies differ on the quantities of the utilized gases for the syntheses and on the produced chemical (methane/methanol). The case studies are:


The selected scenarios want to cover the short-, medium-, and long-term technology deployment horizon and to provide useful information in order to reduce the relevant costs when moving towards the large deployment of the proposed technological option. In particular, the Cases 1, 2, and 4 are adopted because they actually present the medium- to long-term capacities required to deploy the proposed technological option. On the other hand, Cases 3 and 5 represent a shorter-term demonstration of that technological option to move towards decarbonization of steelmaking.

After the description of the case studies in Section 3.1, the overall process scheme is presented in Section 3.2 that includes the aforementioned systems (gas cleaning, hydrogen production, methane, and methanol synthesis) as well as the power plant. A gas cleaning strategy is proposed in Section 3.3 based on the possible contained impurities, whereas Sections 3.4 and 3.5 involve the description of the methanol and methane synthesis processes and the followed AspenPlusTM modelling methodology. Sensitivity analyses on crucial modelling approaches and operating parameters are conducted on both processes. Finally, Section 3.6 includes the description of PEM electrolysis for the production of renewable hydrogen.

## *3.1. Case Studies Description*

The integration of methane and methanol synthesis is evaluated for the previously described scenarios considering a steelmaking plant of medium size with an annual steel production of about 6 MT. The different case scenarios are evaluated in terms of carbon conversion, product yields, hydrogen requirements and consumption, electrolysis demands, as well as overall efficiency of the process. For the cases where power is produced, it is assumed that a gas-fired boiler is in operation within the plant [16]. Figures 1–5 depict the different flowrates and utilization factors of the steelworks off-gases.

**Figure 1.** Case 1—100% of off-gases as input for methanation.

**Figure 2.** Case 2—methanation of 80% of by-product gases.

**Figure 3.** Case 3—replacement of natural gas demands by methanation.

**Figure 4.** Case 4—methanol synthesis of 80% of by-product gases.

**Figure 5.** Case 5—100% replacement of natural gas demands and methanol synthesis.

1. 100% utilization of the produced by-product gases for the production of methane

The first case represents the utilization of the entire available amount of the three off-gases for the production of methane. Renewable hydrogen is added in basic stoichiometric ratio to produce methane. This represents a boundary scenario for the utilization of the steelwork gases, which is restrictive in terms of hydrogen flows and electrolysis power requirements.

2. Methanation of 80% of the available by-product gases and the remaining fraction used in the power plant

The steelworks off-gases have generally various uses within the steel plant and the main one is for the production of electrical power. However, in this scenario, 80% of the total amounts of the gases are used for the production of methane with the addition of renewable hydrogen and the other fraction is sent to the power plant. Before the methanation process, the enrichment step serves as mixing/upgrading process before entering the power plant. A part of the COG is dispatched directly to methane synthesis, due to its higher hydrogen content compared to the other gases. Although, the CH4 content of COG could have a negative impact in methanation activity, the amount of available COG is relatively small compared to the total utilized gases, reducing significantly the methane quantities (i.e., <1%) in the reactor inlet of the methanation cases (i.e., after the H2 addition).

3. Methanation of specific amounts of the by-product gases in order to replace the natural gas demands of the plant

This case investigates the possibility of valorizing the steelworks off-gases for the replacement of the internal steelworks needs of natural gas—assuming an overestimated case of approximately 50,000 Nm3/h internal natural gas demands for a 6 MT/year steel plant [12]. The remaining portion of the gases is combusted in the power plant.

4. Methanol synthesis of 80% of the by-product gases and the rest goes to the power plant

Similar with Case 2, 80% of the amounts of the gases are used for methanol production with the addition of renewable hydrogen, whereas the remaining portion is used in the power plant.

5. Methanation of specific amounts of the by-product gases in order to replace the natural gas demands and the production of significant quantities of MeOH

Case 5 represents the most integrated valorization scheme for steel gases for the simultaneous production of methane and methanol. After the enrichment step, half of the amount employed in Case 4 is used for methanol production and another part is used for the replacement of the industry's natural gas demands. Finally, a part of the gases is sent to the power plant.

Figure 6 shows a comparison of the energy content of the mixed feedstock of Case 1 with respect to the energy content of the other cases, by highlighting the different off-gases contribution. In the first scenario, BFG comprises 73% of the total energy content of the feed stream to the syntheses processes, due to the larger used flowrate, whereas COG, although in lower quantity (20,000 Nm3/h compared to 365,000 Nm3/h of the BFG; 5% of the total amount) contains a significant portion of the overall energy content (19%). This indicates a higher energy content per m3, due to the contained CH4 and H2 and the lower CO2 and N2 contents in the COG feedstock. Regarding the energy contents of the feedstock used in the different scenarios, Cases 2 and 4 use the same feed quantities for the production of chemicals (81% of the energy compared to Case 1). Case 3 represents the feedstock energy content for the replacement of the natural gas demand of the plant, while Case 5 is a combined case for the replacement of natural gas as well as for methanol synthesis (33% and 70%, respectively).

**Figure 6.** Feedstock energy content for the different case scenarios.

## *3.2. Integration Options of CH4/MeOH Syntheses Concepts into Steelworks*

Figure 7 shows the overall process flowsheet that includes the major sections of the proposed concept: gas conditioning, methane production, methanol synthesis, hydrogen production, and the power plant (for the cases where power is produced).

**Figure 7.** Integration of synthesis units into steelworks—superstructure flowsheet.

The mixture of the steelworks off-gases, after an ad hoc conditioning for removal of unwanted impurities, is fed either to the methanol synthesis or to the methanation section. For the methanol synthesis, the feed gas undergoes compression in three stages in order to reduce the associated compression ratio costs; intermediate cooling between the stages is provided. Afterwards, hydrogen is added to reach the required stoichiometric number

and the inlet mixture is preheated before inserted to the synthesis reactor. The produced mixture is separated using a flash separator into the liquid (mainly methanol and water) and gaseous products that consist of the unreacted hydrogen and the rest of the initial feedstock. The last step is the purification of the methanol product in a distillation column, which removes the contained product water.

The first step for methanation requires compression only in one step, since methanation takes place at low pressures (<10 bar). Renewable hydrogen is added to achieve the required stoichiometric ratio and the inlet feed is preheated and directed to the reactor. A flash separator is also used to separate the gaseous products from the produced liquid water. Table 2 depicts the assumptions that refer to the overall process flowsheet simulations. The off-gases composition reported in Table 1 is taken as the starting point of the subsequent flowsheet simulations.


**Table 2.** General assumptions and specifications of the overall process.

## *3.3. Impurities and Gas Conditioning*

The three steelwork off-gases undergo different cleaning steps in order to remove the contained, unwanted components before being stored in gas holders. Typical gas cleaning steps involve dust removal, cooling, scrubbing (for ammonia and BTX removal), and demistering [19,25]. After the initial steps, additional gas cleaning is required to protect the methane [26] and methanol [27] syntheses catalysts.

As shown in Table 1, several sulfur-containing compounds can be found in the steelworks off-gases, which cause corrosion and poisoning of Cu-based catalysts. Other common impurities include nitrogen-containing species such as ammonia or hydrogen cyanide. At the high temperatures of the steel production processes, nitrogen oxides NOx can be formed, which have to be removed from the exhaust gases, whereas at lower temperatures, NH3 can be adsorbed at catalyst sites, reducing the catalyst activity [28]. Halogens (HCl, HF, and HBr) are also contained in the off-gases and are known to cause corrosion and poison catalysts. In particular, experimental works have shown that HCl poisoning could cause loss of the active surface area of the catalyst and promote sintering of the copper crystallites [27]. Furthermore, additional reactions could occur between HCl and other contaminant-forming species such as NH4Cl and NaCl, which when condensed, could cause fouling and create deposits in cooler downstream pipes and equipment [29,30]. Finally, trace elements and heavy metals are also contained in the off-gases due to the diverse nature of the feedstocks. Besides corrosion problems, other trace elements pose a threat to human health and the environment. The distribution and partitioning of these contaminants play an important role on the undertaken cleaning strategy. For example, particle filters could be used for solid particles, but if those compounds appear in the

gaseous phase, more advanced cleaning efforts should be employed, such as solid sorption. Whether a trace element appears in the gas or particulate phase and in which form, depends on following factors [31]:


Based on the contained impurities, Figure 8 depicts the proposed off-gases cleaning strategy. Each of the cleaning steps targets aims at a specific impurity group. However, possible interactions between an impurity and a precedent/succeeding step cannot be ruled out.

**Figure 8.** Proposed gas cleaning scheme.

A first step is devoted to the removal of any contained solid particles through fine filters. Afterwards, the contained halogens (HCl, HF, etc.) are removed using inexpensive sorption materials such as NaHCO3 (Nahcolite) or Trona (Na2CO3-NaHCO3-2H2O) [28]. For instance, in the case of nahcolite, HCl is removed in the form of NaCl, whereas H2O and CO2 are also formed, according to the following reaction:

$$\text{NaHCO}\_3 + \text{HCl} \rightarrow \text{NaCl} + \text{H}\_2\text{O} + \text{CO}\_2\text{O}$$

Regarding the sulfur-containing compounds, H2S is more easily removed at ppb levels with respect to other sulfur species. A common strategy consists in converting organic sulfur compounds to H2S and then employing adsorption technologies for the deep removal of H2S [32]. The avoidance of acid gas removal process, such as SelexolTM or RectisolTM, despite their efficiency in reducing H2S to ppm levels, lies within their affinity to physically absorb CO2, which should otherwise be used as feedstock for the production of methanol/methane [26].

At the hydrodesulfurization (HDS) reactor, organic sulfur compounds and COS are converted to H2S through the addition of hydrogen. The usual employed catalysts are based on cobalt and nickel. A possible reaction network for the conversion to hydrogen sulfide is the following [33]:

$$\text{COS} + \text{H}\_2 \rightarrow \text{H}\_2\text{S} + \text{CO}$$

$$\text{CS}\_2 + 4\text{H}\_2 \rightarrow 2\text{H}\_2\text{S} + \text{CH}\_4.$$

Afterwards, a sorption bed containing metal oxides, such as CaO and ZnO, can be used for the removal of H2S. For the case of ZnO, H2S is removed in the form of ZnS [34]:

$$\text{H}\_2\text{S} + \text{ZnO} \rightarrow \text{H}\_2\text{O} + \text{ZnS}.$$

It is an exothermic process, conducted at T < 250 ◦C and as shown in the reaction stoichiometry, the reaction equilibrium is not affected by pressure, whereas the inlet content of water could affect the H2S removal efficiency. Studies have shown that H2S can be effectively removed at ppb levels employing the ZnO strategy [34,35]. However, due to the contained CO and CO2, additional reactions could occur, with a consequent deterioration of the H2S removal efficiency [34,36].

Finally, a guard bed is placed, containing nickel or other inexpensive material to protect the subsequent synthesis units. It restricts impurities that could have escaped from the former gas cleaning steps and acts as a final protection before the production of chemicals. In addition, the gases are dried to remove the contained water to avoid condensation during compression and/or the promotion of unwanted side-reactions.

## *3.4. Methanol Synthesis*

3.4.1. Process Description and Modelling Approach

Methanol synthesis is based on the following three reactions:


The first two hydrogenation reactions can be combined to form the reverse water gas shift (RWGS) reaction, indicating thus a dependency in-between the reaction system [37]. The catalytic methanol synthesis is exothermic and thermodynamically favored by lower temperatures and higher pressures. Today most of the world methanol production is covered with natural gas derived synthesis gas that after H2/CO ratio adjustment is catalytically processed at 50–100 bar and temperatures between 200–300 ◦C (temperatures required for the activation of the employed catalyst) [38]. An alternative consideration could be a process occurring at much higher pressures (above 100 bar), which would result into an increase in the CO and CO2 conversion rates and thus lowering the needs for carbon recycling [39,40]. This would, however, result in increasing compression costs and power demands and therefore, it was not adopted in this study.

The most common MeOH catalyst employed in industrial scale is based on CuO/ZnO/Al2O3, which is also considered in this study. At higher synthesis temperatures, sintering could take place resulting in higher deactivation rate of the catalyst [41]. The produced water, mainly by CO2 hydrogenation, apart from affecting the equilibrium, could also adsorb on the catalyst sites and promote catalyst sintering [42]. In past works, in-situ water removal was proposed for the enhancement of the thermodynamic equilibrium concentration [43]. The methanol synthesis reaction is characterized by the stoichiometric number (S.N.) where [H2], [CO], and [CO2] refer to the molar flows of the feed components: S.N. = [H2]−[CO2] [CO]+[CO2] . A value of S.N.=2refers to a stoichiometric correlation between the components, whereas the optimum case is slightly above the stoichiometric number [41].

Methanol synthesis applications result in conversion close to what the thermodynamic equilibrium dictates. Any additional hydrogen is not consumed throughout the process and remains unexploited [38]. Therefore, the process economics could benefit from separating the residual H2 and reuse it in the synthesis reactor.

In this work, MeOH synthesis reactor is simulated using two different approaches: a thermodynamic and a kinetic approach. The thermodynamic approach is represented by an AspenPlusTM RGibbs reactor model (based on Gibbs free energy minimization), which for a given pressure and temperature, calculates the equilibrium concentration of selected components. The kinetic approach utilizes the kinetic model developed by Vanden Bussche (with WHSV = 2 kgfeed kgcat−<sup>1</sup> h−<sup>1</sup> and Bed Voidage: 0.33) [44]. Simulation results have shown that for the studied conditions, the deviation of the two modelling approaches is within an acceptable range (<5%) and therefore, the thermodynamic approach is being employed in the investigations.

## 3.4.2. Modelling Assumptions

The modelling of methanol synthesis is based on chemical equilibrium by means of minimization of the Gibbs free energy. Certain components included in the feed mixture are assumed as inert components, having thus no influence in the reaction. Apart from nitrogen, ethane, and methane are also treated as inert gases that do not affect reaction equilibrium. The property method that is used in the flowsheet simulations is Soave– Redlich–Kwong equation of state [45], as past works have proven that it is suitable for methanol synthesis applications [46,47]. Table 3 shows the assumptions and specifications for the thermodynamic MeOH synthesis model.


**Table 3.** General assumptions and specifications of the MeOH AspenPlusTM model.

## 3.4.3. Sensitivity Analysis

Since the methanol synthesis reaction is exothermic, it is thermodynamically favored by lower temperatures. However, the temperature range of the catalyst's activation should also be taken into consideration in order to find the optimum operating conditions. Higher pressures are thermodynamically preferred for methanol production (Figure 9a–c), but the higher compression costs should also be taken into account. Figure 9c depicts the lower conversion rate of CO2 compared to CO, whereas at higher temperature and pressure values, CO conversion tends to decrease and CO2 to increase. This fact can be attributed to the WGS reaction, which is an endothermic reaction and is thermodynamically favored by higher temperatures.

Figure 9d–f shows the influence of increasing stoichiometric number, e.g., increasing input hydrogen. Higher stoichiometric numbers result in higher methanol yields for a given operating temperature (Figure 9d). However, relatively to the input, less hydrogen is consumed in the reactor (Figure 9e) and more remains unexploited in the outlet gaseous fraction (Figure 9f), which refers to the gaseous stream after the separation of methanol and water. Figure 9g illustrates the need for drying of the feed mixture before entering the synthesis reactor. It can be seen that an increase in the water content of the inlet feed leads to a strong decrease in the maximum attained methanol yield. This behavior can be attributed to the promotion of the WGS reaction and consequent CO conversion to additional CO2, at the expense of the methanol synthesis reactions.

The higher the input hydrogen flowrates, the higher the quantity that remains unexploited during the process. Even in sub-stoichiometric ratios, the remaining hydrogen is in considerable portions, which illustrates the need for efficient hydrogen management throughout the process. This could be achieved either through operating in substoichiometric numbers or employing hydrogen recirculation technologies to lower the needs for additional hydrogen and increase the overall efficiency of the system.

**Figure 9.** Methanol synthesis at different operating conditions: (**a**) methanol yield at constant stoichiometric number, (**b**) hydrogen consumption at constant stoichiometric number, (**c**) carbon conversion at constant stoichiometric number, (**d**) methanol yield at constant pressure, (**e**) hydrogen consumption at constant pressure, (**f**) residual hydrogen at constant pressure, and (**g**) methanol yield—inlet water content at constant stoichiometric number and temperature.

## *3.5. Methane Production*

3.5.1. Process Description and Modelling Approach

Syngas methanation is a highly exothermic process aiming at the production of Substitute Natural Gas (SNG) from CO and CO2 with the addition of H2 at the required stoichiometries. The simplicity and high efficiency of the process have been crucial parameters for the establishment of this technology for the production of methane from

waste feedstocks such as biomass [48,49] or steelwork off-gases [19]. For the production of methane from syngas, the main occurring reactions are:


However, based on experimental results, it is assumed to consist of a more complex reaction network, taking provision also for the formation of solid carbon throughout the process [22].

A variety of catalysts are employed for the catalytic methanation process. In this work, a nickel-based catalyst is considered, since it is mostly employed in commercial applications due to the high activity and low associated costs [50]. Similar to methanol synthesis, methane production is also favored by low temperatures and higher pressures because it results in reduction in the total volume. There is currently a variety of established methanation concepts operating at different conditions and reactor configurations [50]. In this work, methanation is conducted in low temperatures 200–300 ◦C and pressures < 10 bar.

Again, two approaches based on kinetics and thermodynamics are compared using the reactor inlet composition of Case 1, as a comparison basis. The first approach is based on the Langmuir-Hinshelwood (LHHW) kinetics derived from Kopyscinski et al. (for WHSV = 2 × <sup>10</sup>−<sup>5</sup> kgfeed kgcat−<sup>1</sup> <sup>h</sup>−<sup>1</sup> and Bed Voidage = 0.33) [51]. Because the LHHW kinetics cannot be implemented directly in AspenPlusTM, a revised form was adopted [52]. The second approach is also based on the thermodynamic model using the Gibbs free energy minimization method, as already explained for the methanol synthesis system. Simulation results have shown that the two modelling approaches agree well for the studied conditions (deviation < 5%) and therefore, for the subsequent sensitivity analyses and case studies evaluation, the thermodynamic model was used.

## 3.5.2. Modelling Assumptions

Modelling of the methanation section is also based on the minimization of the Gibbs free energy. In this case, nitrogen and ethane are treated as inert gases and the used property method in the flowsheet simulations is Soave–Redlich–Kwong equation of state [22,23]. Table 4 illustrates the assumptions for the AspenPlusTM model of methane synthesis.


**Table 4.** General assumptions and specifications of the AspenPlusTM methanation model.

## 3.5.3. Sensitivity Analysis

Figure 10 shows some crucial characteristics of the process at different operating parameters of methane production. As depicted in Figure 10, CO conversion is almost complete, irrespectively of the operating pressure and temperature range (T = 200–300 ◦C, P = 1–10 bar). On the other hand, CO2 conversion, strongly depends on the operating parameters; lower temperature and higher pressure favor the CO2 conversion thermodynamically. Higher conversion rates mean higher hydrogen consumption (see Figure 10b), resulting in >95% consumption in any pressure and temperature range. This results in low

portion of the hydrogen remaining unexploited in the off-gases of the methanation process. In addition, Figure 10c shows that methane production is favored in any of these operating conditions obtaining a methane yield greater than 95%.

**Figure 10.** Methane production at different operating temperature and pressure: (**a**) CO2 conversion, (**b**) hydrogen consumption, and (**c**) methane yield.

## *3.6. Hydrogen Production*

Due to the composition of the steelworks off-gases, hydrogen addition is required in order to reach certain stoichiometric ratios and improve the efficiency of methane and methanol syntheses. In order to obtain both economic and environmental advantages, hydrogen needs to be produced in an environmentally friendly way, i.e., by exploiting renewable sources. In this work, the adopted process is water electrolysis fed by renewable energy.

During water electrolysis, the water molecules are split into hydrogen and oxygen by means of electricity. There are three main electrolysis processes, each one differing on the operating principles and conditions: alkaline exchange membrane (AEM), proton exchange membrane (PEM) and solid oxide electrolysis (SOE) [53]. In the present work, PEM electrolysis is considered as the option for renewable hydrogen production, since it is an already established technology, it is used in large-scale industrial applications and it is not sensitive to the fluctuations in power supply, such as in the case of renewable energy sources [24].

For the calculation of the power requirements, an AspenPlusTM PEM electrolysis model has been developed. The model incorporates the following main reactions occurring in the two PEM sections:

$$2\mathrm{H}\_{2}\mathrm{O} \rightarrow \mathrm{O}\_{2} + 4\mathrm{H}^{+} + 2\mathrm{e}^{-} \qquad\qquad\text{anode section}$$

$$2\mathrm{H}^{+} + 2\mathrm{e}^{-} \rightarrow \mathrm{H}\_{2} \qquad\qquad\text{cathode section.}$$

In addition, further phenomena taking place inside a PEM electrolysis module are also considered such as hydrogen and oxygen permeations [54,55] and water diffusion [55,56], which are estimated in ad hoc configured calculator blocks. Highly pure hydrogen is

assumed to be produced (<99.9 vol.%), and the overall electrical energy consumption for the stack is 54.8 kWh/kg H2. The produced hydrogen is stored into pressurized vessels and compressed to achieve the conditions required for the syntheses units.

A detailed description of the PEM electrolysis model, as well as of other renewable hydrogen production technologies that are considered possible solutions for the enrichment of steelworks off-gases, is included in another publication by the authors [57].

## **4. Results**

In this section, the proposed case studies are evaluated using the aforementioned AspenPlusTM models. The results presented in Table 5 focus on the hydrogen requirements and consumption, electrolysis demands, product yields, and carbon conversion. Figure 11 shows these key results. Further specific indicative stream results are available in the Appendix A.

For the three first cases, which refer only to methanation, carbon is almost completely converted, compared to Case 4, which includes only methanol synthesis, and Case 5, which is a combination between methane and methanol syntheses. In Case 4, in particular, the low CO2 utilization rates indicate that CO (see Figure 11a) and not CO2 (see Figure 11b) is consumed for methanol synthesis. In addition, for the different case studies, the higher the carbon conversion rate, the higher the hydrogen consumption throughout the process (see Figure 11c). Figure 11d shows the produced electrical power of the different cases compared to the base-case, that refers to the traditional, full-scale utilization of the steelworks offgases for power production. Case 1 is not included in the comparison, as the whole amount of off-gases is used for methane production. For Cases 2 and 4, the same power is produced since the same feedstock amount is used in the syntheses units (19% of the total power). Case 3 produces 60% of the power produced in the base-case and Case 5, which is the most integrated scheme, involves the production of 30% of the total power.

If different stoichiometric ratios are chosen (other than the stoichiometric for methane and 1.7 for methanol synthesis) the required hydrogen feed inputs are greatly affected, as illustrated in Figure 12. Case 1, which refers to the full-scale utilization of the steelworks byproduct gases, requires more hydrogen compared to the other cases at any stoichiometric ratio. Although Cases 2 and 4 refer to utilization of the same feedstock flowrates, methane synthesis (Case 2) requires more hydrogen compared to methanol synthesis (Case 5), due to the higher carbon conversion. As a consequence of the lower carbon conversion, the rest of the hydrogen remains unexploited in the off-gases of the methanol synthesis process.


**Table 5.** Case studies key results (methanation: T = 250 ◦C, P = 5 bar and stoichiometric H2, MeOH synthesis: T = 250 ◦C, P = 70 bar, and S.N. = 1.7).

**Figure 11.** Key Results of the case studies (methanation: T = 250 ◦C, P = 5 bar and stoichiometric H2, MeOH synthesis: T = 250 ◦C, P = 70 bar and S.N. = 1.7): (**a**) carbon conversion, (**b**) CO2 utilization, (**c**) hydrogen consumption, and (**d**) power production.

**Figure 12.** Hydrogen requirements for the different stoichiometries per case.

This is also verified in Figure 13, which depicts the energy content of the residual off-gases after each synthesis processes and after the separation of the total amounts of produced methane and methanol, in each respective case. Especially for Cases 4 and 5, the remaining off-gases have a significant energetic value due to the large fraction of unreacted hydrogen. These residual off-gases could either be used for combustion to support heat-intensive processes, or hydrogen recycling should be included to avoid producing additional hydrogen by electrolysis. In conventional methanol synthesis loops, a flash drum separates the methanol and water products from the unreacted gaseous components, which are recycled back to the synthesis reactor [41]. Alternative processes to recover only the residual hydrogen include technologies such as pressure swing adsorption [58], membranes [59], and/or electrochemical hydrogen compression [60].

**Figure 13.** Residual off-gases energy content for the different stoichiometries per case.

The PEM electrolysis requirements, as illustrated in Figure 14, are in the range of GWs, which are restrictive for employment in full-scale, by considering the capacities of currently available commercial electrolyzers.

**Figure 14.** PEM electrolysis requirements for the different cases.

In terms of comparison between the energy content of the total feedstock (used in each case) and the electrolysis requirements, Figure 15 shows the major energy streams of the Cases 1, 4, and 5. In Case 1, 383% of the energy content is contained in the methane product, 89% is contained in the methanol of Case 4, whereas 105% and 46% are contained in the methanol and methanol products of Case 5. The electrolysis requirements of each case are noticeable. In Case 1, 631% of the energy of the feedstock is required for electrolysis, in comparison to 304% and 322% in Cases 4 and 5. Regarding the MeOH synthesis cases, however, these figures could be further reduced, if certain amounts of the residual hydrogen are recycled.

Figure 16 compares the PEM electrolysis energy requirements of the base case without H2 recycling, to recycling 25%, 50%, and 75% of the residual hydrogen. If recycling 75% of the residual hydrogen is pursued, almost 50% less power is required for electrolysis, indicating, thus opportunity for further optimization of the process.

**Figure 15.** Sankey diagrams—energy analysis for three key scenarios.

**Figure 16.** Influence of hydrogen recycling in PEM electrolysis requirements of Case 4.

Regarding Case 5, the benefits of recycling 75% could result in almost 25% lower electrolysis requirements (see Figure 17). The lower savings percentage compared to Case 4 is due to the methanation section of Case 5, which consumes a major fraction of the input hydrogen, resulting in less available hydrogen for recycling, which is also verified in Figure 11.

The benefits of recycling can also be seen in Figure 18, which refer to Cases 4 and 5 and the required electrolysis power to the energy content of the steelworks feedstock (as shown in Figure 15). The 304% of Case 4 could be reduced to 158% and from 322% to 244% for Case 5 if recycling 75% of the hydrogen is pursued.

**Figure 17.** Influence of hydrogen recycling in PEM electrolysis requirements of Case 5.

**Figure 18.** Electrolysis requirements compared to feedstock energy content for the different recycling cases for Case 4 and Case 5.

## **5. Discussion**

In this work, the integration of hydrogen-intensified methane and methanol synthesis is investigated for three available steelworks off-gases by means of AspenPlusTM flowsheet simulations. The composition of the off-gases is analyzed and a generic gas cleaning scheme is proposed for the removal of the contained impurities that can affect the catalyst operation, which is used in the CH4 and CH3OH syntheses. Thermodynamic and kinetic AspenPlusTM models are compared for the investigation of methanol and methane synthesis processes. The studied conditions for methane synthesis include 200–300 ◦C, stoichiometric number 1–1.1 and pressure 1–10 bar and for methanol synthesis: 200–300 ◦C, stoichiometric number 1.7–2.1, and pressure 50–100 bar. Furthermore, case studies corresponding to different usage of the steelworks off-gases for chemicals production are investigated in terms of hydrogen requirements and consumption, carbon conversion, product yields, and PEM electrolysis requirements. The cases of methane synthesis depict high CO and CO2 conversion rates that almost eliminate the CO2 emissions of the steel plant. In case of increasing carbon credits, this would represent a significant financial benefit and therefore, a carbon credit avoidance could be of high importance. On the other hand, methanol synthesis produces a product with higher market value, but only converts approximately 40% of the carbon emissions into a renewable fuel/chemical. The choice between methane and methanol production or a combination of the two, will be a result of an upcoming cost estimation study but also on the relevant prices of the products, as well as, on the carbon credits. The energy content of the hydrogen employed for this transformation far overcomes the energy off-set of the steel plants. However, recycling of the residual hydrogen in the methanolinvolved cases could lead to substantial benefits in terms of electrolysis requirements. Calculations show that reductions in the range of 50% for Case 4 and 25% for Case 5 of

electrolysis requirements could be achieved, when recycling 75% of the residual hydrogen. Future studies will involve the capital and operating cost estimation analysis of the case studies as well as the application of state-of-the art and alternative hydrogen recirculation technologies for recycling the residual hydrogen from the methanol synthesis cases. In addition, the application of an advanced dispatch controller will be investigated in order to optimize the management of steelworks off-gases among internal users, power plant and syntheses processes, and the related hydrogen requirements.

**Author Contributions:** Conceptualization, M.B., K.P., I.M. and V.C.; methodology, M.B., K.P., I.M., V.I. and V.C.; software, M.B., K.P. and S.V.; validation, M.B., K.P., I.M. and S.D.; formal analysis, K.P., T.A.B. and V.C.; investigation, M.B. and K.P.; resources, K.P., S.V., P.S. and V.C.; data curation, M.B., K.P., A.P. and A.Z.; writing—original draft preparation, M.B. and K.P.; writing—review and editing, M.B., K.P., I.M., V.C., S.D., V.I., A.P., A.Z., T.A.B., P.S. and S.V.; visualization, M.B. and K.P.; supervision, K.P., P.S. and S.V.; project administration, K.P. and V.C.; funding acquisition, K.P. and V.C. All authors have read and agreed to the published version of the manuscript.

**Funding:** This research was funded by the European Union through the Research Fund for Coal and Steel (RFCS), Grant Agreement No. 800659.

**Institutional Review Board Statement:** Not applicable.

**Informed Consent Statement:** Not applicable.

**Data Availability Statement:** Not applicable.

**Acknowledgments:** The work described in this paper was developed within the project entitled "Integrated and intelligent upgrade of carbon sources through hydrogen addition for the steel industry" (i3upgrade, GA No. 800659), which has received funding from the Research Fund for Coal and Steel of the European Union. The sole responsibility of the issues treated in this paper lies with the authors; the Commission is not responsible for any use that may be made of the information contained therein.

**Conflicts of Interest:** The authors declare no conflict of interest.

## **Appendix A**

Indicative process stream results are presented in the following tables, namely, the steelworks off-gases feedstock, the added hydrogen, the reactor inlet, and outlet of the methanol and methane sections (see Figure 7). Note that the depicted cases refer to methane synthesis conducted at T = 250 ◦C, P = 5 bar, and stoichiometric H2, whereas for methanol synthesis, T = 250 ◦C, P = 70 bar, and S.N. = 1.7.


**Table A1.** Stream results of Case 1.


**Table A2.** Stream results of Case 2.

**Table A3.** Stream results of Case 3.





**Table A5.** Stream results of Case 5—MeOH synthesis.

**Table A6.** Stream results of Case 5—methane production.


## **References**


## *Article* **In Situ Catalytic Methanation of Real Steelworks Gases**

**Philipp Wolf-Zoellner 1, Ana Roza Medved 1, Markus Lehner 1,\*, Nina Kieberger <sup>2</sup> and Katharina Rechberger <sup>3</sup>**


**Abstract:** The by-product gases from the blast furnace and converter of an integrated steelworks highly contribute to today's global CO2 emissions. Therefore, the steel industry is working on solutions to utilise these gases as a carbon source for product synthesis in order to reduce the amount of CO2 that is released into the environment. One possibility is the conversion of CO2 and CO to synthetic natural gas through methanation. This process is currently extensively researched, as the synthetic natural gas can be directly utilised in the integrated steelworks again, substituting for natural gas. This work addresses the in situ methanation of real steelworks gases in a lab-scaled, three-stage reactor setup, whereby the by-product gases are directly bottled at an integrated steel plant during normal operation, and are not further treated, i.e., by a CO2 separation step. Therefore, high shares of nitrogen are present in the feed gas for the methanation. Furthermore, due to the catalyst poisons present in the only pre-cleaned steelworks gases, an additional gas-cleaning step based on CuO-coated activated carbon is implemented to prevent an instant catalyst deactivation. Results show that, with the filter included, the steady state methanation of real blast furnace and converter gases can be performed without any noticeable deactivation in the catalyst performance.

**Keywords:** power-to-gas; catalytic methanation; steelworks; real gases; activated carbon; catalyst poison and degradation

## **1. Introduction**

Integrated steelworks are major contributors to today's global CO2-emissions. Review publications screening the steelmaking process around the globe revealed that approximately 27 to 30% of any industrial CO2 emissions are directly linked to this sector [1,2]. With a world-wide crude steel production of 1869 million tonnes in 2019 and a 3.6% per annum average growth rate, the steel demand of our society is increasing strongly. These large amounts of steel are mainly required for building and infrastructure (~52%), mechanical equipment (~16%) and the automotive sector (~12%) [3]. Figure 1 shows the most common route of steelmaking globally, which includes a blast furnace for the reduction of iron ore to hot metal and a converter or basic oxygen furnace for the batch-wise production of molten steel. The accumulating by-product gases, such as the blast furnace gas (BFG), basic oxygen furnace gas (BOFG) and coke oven gas (COG), have a very rich content of CO2 and carbon monoxide (CO), among other gases (Table 1). At the current stage, these by-product gases are buffered within the steelwork and utilised as an energy carrier internally. Nevertheless, additional fossil energy sources, such as natural gas, are needed to cover the whole energy demand for the power plant and auxiliary energy conversion. Prior to any further use, the product gases are cleaned in a two-stage process, including, for example, a dust collector for the separation of coarse dust, and a venturi scrubber for fine dust and water-soluble components.

**Citation:** Wolf-Zoellner, P.; Medved, A.R.; Lehner, M.; Kieberger, N.; Rechberger, K. In Situ Catalytic Methanation of Real Steelworks Gases. *Energies* **2021**, *14*, 8131. https://doi.org/10.3390/ en14238131

Academic Editor: Dino Musmarra

Received: 9 November 2021 Accepted: 29 November 2021 Published: 3 December 2021

**Publisher's Note:** MDPI stays neutral with regard to jurisdictional claims in published maps and institutional affiliations.

**Copyright:** © 2021 by the authors. Licensee MDPI, Basel, Switzerland. This article is an open access article distributed under the terms and conditions of the Creative Commons Attribution (CC BY) license (https:// creativecommons.org/licenses/by/ 4.0/).

**Figure 1.** Schematic of steelmaking process—blast furnace/basic oxygen furnace route.


**Table 1.** Gas composition of by-product gases from a typical steelworks plant [4].

With the challenging targets of the climate agreements being set, the steel industry sector logically seeks for possibilities to reduce their greenhouse gas (GHG) emissions, as well as to incorporate green energy sources in the steelmaking process itself. One way of reducing the GHG emissions, and simultaneously substituting the need for fossil fuels, is the implementation of synthesis processes, such as methanation. In this process, CO2 and CO react with hydrogen (H2), gained from green energy sources—for instance, renewable power driving water electrolysis—to create methane (CH4) and steam (Equations (1) and (2)) [5].

$$\text{CO} + 3\,\text{H}\_2 = \text{CH}\_4 + \text{H}\_2\text{O(g)}\qquad\qquad\Delta\text{H}\_r^0 = -206\,\text{kJ/mol}\tag{1}$$

$$\text{CH}\_2 + 4\,\text{H}\_2 = \text{CH}\_4 + 2\,\text{H}\_2\text{O}(\text{g}) \qquad\qquad\qquad\Delta\text{H}\_r^0 = -165\,\text{kJ/mol}\tag{2}$$

These two reactions are highly exothermic and are linked via the reverse water–gas shift reaction (Equation (3)).

$$\text{CO}\_2 + \text{H}\_2 = \text{CO} + \text{H}\_2\text{O(g)}\qquad\qquad\Delta\text{H}\_r^0 = \text{41 kJ/mol}\tag{3}$$

Although these reactions are well-known, the behaviour with real, untreated steelworks gases, as the COx source, are yet to be investigated. The detailed fundamentals behind the methanation concept, possible reactor designs and available catalysts are documented by Rönsch et al., combining them with an up-to date overview on methanation projects and the state-of-the-art in the research [6].

Current work on the methanation of steelworks gases primarily focuses on the usage of COG due to its favourable composition. The high amount of hydrogen (up to ~60%) makes it an attractive feedstock to produce synthetic natural gas (SNG), as it also works as an alternative hydrogen source compared to a solution involving an electrolysing unit (e.g., PEM). Müller et al. [7] investigated the direct conversion of CO and CO2 from synthetic COG into methane using nickel (Ni)-based catalysts in a fixed-bed reactor. It was shown that additional CO2 from other sources (e.g., air, flue gas) is required to compensate for the high surplus of hydrogen in the COG to achieve a desirable methane yield. Razzaq et al. [8] tested various Ni-based catalyst support materials (SiO2, Al2O3, ZrO2 and CeO2) for the methanation of synthetic COG. The COx conversion rates and CH4 selectivity in a fixed-bed reactor were evaluated. Results showed that ZrO2-CeO2-coated catalysts have the highest activity and selectivity for CO and CO2 for synthetic gases with COG composition.

Medved et al. [9] showed in their work that, although the gas from the coke oven seems to be the favourable by-product gas for methanation, it is already fully energetically integrated into the process chain of integrated steelworks. Due to its high calorific value of up to 19.000 kJ/Nm3, it is used plant-internally at the blast furnace for firing processes, as well as for the power plant. Consequently, it is not readily available as an input for a methanation unit without significantly affecting the steelmaking process and disturbing the energy balance of the plant. The utilization of COG necessitates its substitution by external energy sources, such as electric power and natural gas, respectively. Therefore, the other two by-product gases, BFG and BOFG, with their very high amount of CO2, CO and N2, have been evaluated for their applicability as a feed gas for a methanation plant. Furthermore, the authors concluded that the enrichment of BFG and BOFG via methanation, without the necessity of nitrogen removal, as a lean product gas shows a utilisation potential in the integrated steel plant as a substitute for natural gas. Schöß et al. [10] concluded that, although both gases (BFG and BOFG) are suitable carbon sources for the SNG process with the addition of hydrogen for reaction stoichiometry, the necessary specifications of the natural gas grid cannot be met, due to the high content of nitrogen and the resulting low calorific value. In addition, the significant amount of catalyst poisons present in the already cleaned by-product gases needs to be addressed. Lehner et al. [11] added that, for converting steelworks gases to methane, a load flexible reactor setup is favourable to meet the fluctuations in the process gas and hydrogen availability.

Studies on methanation are mainly based on synthetic gas mixtures simulating real gas compositions. Nevertheless, Müller et al. [12], for example, evaluated the direct CO2 methanation of flue gases at a lignite power plant, showing that the same commercial Ni-based catalyst used in this work did not degrade during the time frame of the experiments. The real gases included the following catalyst poisons: 63 ppm SO2, 36 ppm NO2. The same authors further investigated the CO2 methanation of flue gases emitted by conventional power plants [13]. They used synthetic gases simulating the real gas composition, including contaminations of up to 100 ppm NO2 and 80 ppm SO2. The experiments showed a decrease in the conversion, yield and selectivity by 17% in 12.5 h, or 1.36% per hour. The authors also concluded that the deactivation caused by SO2 is low in relation to a possible degradation caused by traces of H2S. Rachow [14] studied the influence of catalyst deactivation by sulphur compounds and NOx components. The author confirmed through experiments with bottled flue gases from coal-fired power plants, as well as with real gases from the cement industry, that Ni-based catalysts strongly degrade in the range of hours when exposed to SO2 and sulphur compounds in general. The degree of deactivation depends on the catalyst, its active surface area, the SO2 concentration and the total volume flow rate. No deactivation was observed during experiments with NOx contaminations. Méndez-Mateos et al. [15] studied the CO2 methanation over modified Ni catalysts, integrating promoters (transition metals, such as Mo, Fe, Co or Cr), which were added to the catalyst formulation in different portions. The target was to improve the catalyst's resistance over sulphur, and H2S in particular. The authors showed that the catalyst activity between 573 and 773 K at 10 bar increased when transition metals were added, except for Mo. In addition, it was possible to regenerate the Co-modified catalyst with oxygen, recovering to a 13% methane yield compared to the fresh catalyst.

Calbry-Muzyka et al. [16] reported on the technical challenges and recent progress made when using biogas as an input for direct methanation. Due to the varying composition of the biogas feedstock, no standard gas cleaning solution has been developed so far. Nevertheless, thorough H2S and non-H2S sulphur removal to the sub-ppm level is necessary in order to prevent catalyst deactivation. Witte et al. [17] demonstrated the stable operation of a catalytic direct methanation with biogas in a fluidised bed reactor for over 1100 h. Only a slow deactivation by organic sulphur compounds was identified, which broke through the gas cleaning unit with concentrations between 0.5 and 3 ppmv. The installed unit contained a two-stage adsorption-based biogas cleaning system for deep desulphurization [18]. Fitzharris et al. [19] concluded in their work that Ni-based catalysts are highly sensitive to poisoning by sulphur due to geometric, and not electronic, effects. Concentrations of H2S as low as 13 ppb in H2 reduced the steady state methanation activity by more than two orders of magnitude. Bartholomew et al. [20] showed that, under typical low-pressure reaction conditions for methanation (525 K, 1 atm) in a reaction mixture containing 10 ppm H2S, Ni-based catalysts lost most of their activity within a period of 2 to 3 days. The rate of deactivation due to H2S poisoning increased with an increasing H2/CO ratio, as well as with an increasing reaction temperature.

Although the applicability of steelworks gases as feed for the synthesis of methane was addressed and confirmed by multiple authors [7–11], experiments with real gases, including contaminations poisonous to catalysts, have not yet been performed. Consequently, the main aim of this work is to show the degree of degradation when using real BFG and BOFG (including nitrogen) as an input for a methanation unit in a composition, as given in Table 1, as well as to present a working solution to overcome the problem of catalyst deactivation.

## **2. Methods and Methodology**

## *2.1. Experimental Setup*

The experiments presented throughout this paper were carried out with the lab-scaled methanation test rig shown in Figure 2 on the left [21]. The technology was validated in relevant environments, consequently representing a Technology Readiness Level of 5 (TRL 5). The unit consists of three cylindric reactors, each made from austenitic stainless steel (304H chromium-nickel (1.4948)), with a height of 300 mm and an inner diameter of 80 mm. The reactors are connected in series but can also be operated individually (Figure 2, right). At the bottom, each reactor is filled with 3/8" stoneware balls over a height of 100 mm in order to ensure an evenly distributed gas stream through the catalyst section located on top of the inert material. The catalysts used are either a commercial Ni-based bulk catalyst (Meth 134®, 3–6 mm spheres with a Ni-loading of 20 wt.-%), or Ni/boehmite wash-coated ceramic honeycombs [22]. For the experimental results shown throughout this work, only the bulk catalyst was used. The remaining volume towards the top of the reactors is again filled with the same inert balls.

Figure 3 shows a basic flow chart of the described reactor setup. A maximum flow rate of 3 m3/h (STP, ~50 NL/min) is possible for the input gas stream. Operating pressures of up to 20 bar(abs) can be maintained and the maximum temperature for the reactors is limited to 700 ◦C. The methanation plant is fed with H2 (hydrogen 5.0 purity), N2 (nitrogen 5.0 purity), CO (carbon monoxide 2.0 purity) and CO2 (BIOGON® C, E290, 99.7% purity) from gas bottles, allowing for the preparation of synthetic gas mixtures to meet the specifications of the by-product gases of interest. In addition, bottled real gases from an integrated steel plant in Austria can be connected to the input stream (additional information provided in Section 2.2). Through thermal mass flow controllers (Bronkhorst), the individual gases enter the gas-mixing station. Before entering the first reactor (R1), the gas stream is preheated in a heat exchanger (W1) to temperatures above 200 ◦C. Additionally, to reach and keep the required temperature of the catalyst at 260 ◦C prior to the methanation synthesis, the reactors are equipped with infrared panels (RS Pro, 4 panels per reactor, 500 Watt each) on the outside. Between the reactors, two further heat exchangers (W2, W3) are installed. The final product gas stream is cooled and guided through a condensate trap

to extract the H2O formed during the synthesis. The product gases are combusted in a flare, which is connected to the aspiration system. Four gas sampling stations (at the input, as well as downstream of each reactor) allow for analysis of gas composition with the use of an infrared photometer URAS 26 for CO, CO2 and CH4, as well as a thermal conductivity analyser CALDOS 27 for H2 (both from ABB GmbH) with a deviation of ≤1% per component. The gas analysers are calibrated once every week.

**Figure 2.** Lab-scale methanation test rig (**left**); reactors and heat exchangers (**right**).

**Figure 3.** Basic flow sheet of lab-scale methanation plant at Montanuniversität Leoben.

The methanation test rig is equipped with a series of type K thermocouples, as well as pressure and flow rate measurement devices. In addition, a multi-thermoelement is added to each reactor measuring the axial temperature profile 22 mm eccentric from the reactor middle axis. In total, five temperatures are measured inside the catalyst section, as well as two further ones directly below and above the catalyst bed (Figure 4 left) [23]. When using a reactor setup with structured honeycomb catalysts, the locations of the temperature readings are modified to measure the actual temperature inside the channels at the bottom, middle and top of the honeycomb in the centre and at 50% of the reactor's radius (Figure 4 right).

**Figure 4.** Schematic of one reactor with multi-thermoelement—bulk (**left**), honeycomb (**right**).

The conversion rate of CO and CO2 as COx is calculated based on the feed and product gas composition. The input gas volume flow and the input gas concentrations are known from the mass flow controller setpoints of each species. Whereas CO, CO2, CH4 and H2 can be measured by the gas analysis system at each reactor outlet, the missing species H2O, as well as the outgoing total gas volume flow, is determined by component and atom balances. The COx conversion is then calculated from the ingoing and outcoming component mole flows of CO, CO2 and CH4, as shown in Equation (4).

$$\text{CO}\_{\text{X}}\text{ conversion}\left[\%\right] = \frac{\text{n}\_{\text{in}} \times \left(\text{x}\_{\text{CO2in}} + \text{x}\_{\text{COin}}\right) - \text{n}\_{\text{out}} \times \left(\text{x}\_{\text{CO2out}} + \text{x}\_{\text{COout}}\right)}{\text{n}\_{\text{in}} \times \left(\text{x}\_{\text{CO2in}} + \text{x}\_{\text{COin}}\right)}\tag{4}$$

## *2.2. Analysis of Bottled Real Gases*

Prior to any methanation experiments with real steelworks gases, the content of the provided gas bottles is analysed for gas composition and any catalyst poisons potentially present. These gas bottles are filled within a mobile gas-filling station located in the steel works plant. Compressed BFG or BOFG can be filled into such gas bottles, with a volume of 20 L up to pressures of ~150 bar. For the analysis, a ThermoFisher Trace GC-ultra equipped with three gas channels is used. Hydrocarbons are resolved on a 30 m Rtx-alumina capillary column (ID 0.53 mm; filling Na2SO4, 10 μm film thickness) and detected by FID. Permanent gases are resolved on two packed columns, HayeSep Q (2 m × 1/8" OD) and MolSieve 5A (2 m × 1/8" OD) and detected by TCD. Sulphur and phosphorus compounds are determined using an Rtx-Sulphur packed column (2 m × 1/8" OD) and an FPD detector. Helium is used as carrier gas for all three channels.

The analysis of the bottled real gases showed that the samples taken after the gas cleaning station and the gasometers are composed like typical average values in the steel industry [4] (Table 2).


**Table 2.** Gas composition of bottled real gases from steelworks per gas type.

n.d.—not detectable.

In addition, the following catalyst poisons were detected in the blast furnace gas [24,25]: carbon disulphide (CS2) with a very small amount of 0.26 mg/Nm3. COS stabilised at 110 mg/Nm3, and the SO2 content was evaluated with 2.2 mg/Nm3. HCl was below the detectable value of the used equipment (<1.0 ppm), but hydrogen sulphide (H2S) was detected with values around 28 mg/Nm3, ammonia (NH3) with 0.15 and HCN with 0.12 mg/Nm3. The values for antimony (Sb), mercury (Hg) and other heavy metals that are poisonous to Ni-based catalysts could not be analysed with the selected method. For the converter gas, only small amounts of COS, H2S and SO2 were detected; the other catalyst poisons were all below the detectable value of the analysis method. Table 3 summarises the catalyst poisons present in the bottled real gases.


**Table 3.** Catalyst poisons in bottled real gases for BFG and BOFG [24,25].

n.d.—not detectable.

## **3. Experiments and Results**

*3.1. Initial Experiments with Bottled Real BFG*

For the experiments performed in this work, a reference base case was defined, which serves as a performance comparison between the methanation of synthetic and real gases under steady state and dynamic operating conditions. The experiments have been carried out first with synthetic mixed gas from gas bottles, and additionally with unconditioned and with pre-cleaned real gases from the steel industry.

The parameters of this reference case are:


These parameters are based on Medved et al. [9], who analysed the influence of nitrogen on the methanation of synthetic steelworks gases. The authors concluded that a 4–5% H2 surplus is required within the tested GHSV to achieve a full methane yield for a three-stage methanation setup outlined above. Hauser et al. [26] reported the same value for a heat pipe cooled structured fixed-bed reactor. For the expression of the reaction stoichiometry, the parameter σH2 is introduced, which describes the ratio of the molar hydrogen flow to the molar flows of CO and CO2 present in the feed gas (Equation (5)). σH2 is equal to 1.0 for stoichiometric mixtures, and is <1.0 for under- and >1.0 for overstoichiometric mixtures, respectively.

$$
\sigma\_{\text{H}\_2}[-] = \frac{\mathbf{n}\_{\text{H}\_2}}{4 \times \mathbf{n}\_{\text{CO}\_2} + 3 \times \mathbf{n}\_{\text{CO}}} \tag{5}
$$

The definition of the gas hourly space velocity (GHSV), which is the ratio of the total feed gas volume flow (Qgas) and catalyst volume Vcatalyst, is given in Equation (6).

$$\text{GHSV}\left[\text{h}^{-1}\right] = \frac{\text{Q}\_{\text{gas}}}{\text{V}\_{\text{catalyst}}}\tag{6}$$

Figure 5 shows the time-based measurement data for the first experiment with real gases and base case parameters. No gas cleaning system was installed yet. At the starting point, one reactor (R2) was used, and steady-state conditions with a synthetic BFG gas mixture, simulating the real gas composition according to Table 2, were established over a period of two hours prior to the experiment start. The feed gas contained CO, CO2 and the inert gas N2, and hydrogen was added to the input stream to reach a surplus of 5% to the reaction stoichiometry (σH2 = 1.05). At time "0", the bottled real BFG substituted for the synthetic gas mixture, whereby the required hydrogen was kept at σH2 = 1.05. During the following eight minutes, an immediate and consistent drop in synthesis performance was observed, as the CH4 content in the product gas stream dropped significantly from ~31 to 24 vol.-%, and, consequently, the CO, CO2 and H2 content increased. After a duration of approximately 16 min, the curves started to flatten, but the performance drop remained at a lower rate. This shows that the conversion for the real BFG is less compared to the one with the synthetic gas mixture at the same H2 surplus, indicating an instant catalyst degradation. As no parameter was varied other than the catalyst poisons, the performance drop can be linked to their presence. This is supported by experimental campaigns performed with synthetic steelworks gases, each with times on stream of 80 h and more (max. 158 h) [9,23,27]. In no case was a catalyst degradation that quick or in this magnitude observed. Although GHSV, σH2 and the operating pressure were varied for synthetic gas mixtures, the overall performance stayed nearly constant during the whole experiment duration.

**Figure 5.** First methanation experiment with real BFG (BFGreal, 4000 h<sup>−</sup>1, 4 bar, 5% H2 excess).

## *3.2. Variation of Hydrogen Excess Rate at Constant GHSV*

With the base case parameters established, experiments with a dynamic H2 excess rate compared to reaction stoichiometry were performed in order to evaluate the performance of the catalyst and its possible degradation over time. The GHSV, as well as the pressure, were kept constant during these experiments (4000 h−1, 4 bar), and all three available reactors were used. Figure 6 shows the COx conversion rates (Equation (4)), as well as the average reactor temperature, for each H2 excess rate per gas type (real vs. synthetic BFG) and reactor (R1 to R3). The measurements were taken in intervals of 20 min, after which, a steady state of the system, as well as a stable gas sampling, was achieved. Furthermore, the experiment start time for the real gases is plotted at the top. Starting with a σH2 of 1.05 at 0 min, a slightly lower COx conversion rate was measured for R1 after 20 min and for R2 after 40 min, compared to the measurements taken with synthetic BFG (BFGsyn). Downstream of R3, still full COx conversion was achieved with the measurement after 60 min. Afterwards, the H2 excess rate was increased to 9% (σH2 = 1.09), which should result in an improved methane yield according to the literature. However, the performance dropped below the one of the 1.05 experiment that started at 0 min. Compared to synthetic gas, no full COx conversion could be achieved with all three reactors. The additional experiments with values for σH2 of 1.02 and 1.0 starting after 120 and 180 min continued the trend towards a significant performance decrease when comparing the real gases with the synthetic gas mixture. This is especially noticeable for the first reactor R1, with a difference of 7.6%-points in COx conversion for the σH2 = 1.0 experiment after approximately three hours. The decrease in catalyst activity is clearly attributable to the catalyst poisons in the methanation feed gases.

**Figure 6.** Comparison of experiments with real and synthetic BFG with varying H2 excess rate (σH2 = 1.05, 1.09, 1.02, 1.00) at 4000 h−<sup>1</sup> and 4 bar, COx conversion in %, reactor temperature in ◦C.

Consequently, additional experiments with a 5% H2 excess rate were performed after this alteration of hydrogen addition to determine the degree of catalyst deactivation. Therefore, the COx conversion rates prior to and after 4 h of methanation with real gases (0–240 min) were compared against each other. Figures 7 and 8 show the COx conversion and product gas composition per reactor, respectively, for the experiment at 0 min (E-1) and another measurement taken under the same operating conditions after 240 min (E-2). Within a period of 4 h, the performance of the first reactor (R1) dropped by 5.3%-points in COx conversion. The second reactor (R2) took over the load as the first reactor's performance dropped, keeping the overall performance of the first two reactors stable. This is confirmed by the temperatures measured inside the reactors, as they decreased for

R1 (median of −71.3 ◦C) and increased for R2 (median of +35.7 ◦C). The small drop in COx conversion for reactor three (R3) is a result of a too-low average reactor temperature adjusted during experiment E-2.

**Figure 7.** Comparison of COx conversion rates in % of two experiments with real BFG, first one (E-1) taken at 0 min, and second one (E-2) taken after 4 h of real gas experiments with varying H2 excess rate.

**Figure 8.** Comparison of real gas experiments E-1 and E-2 with 5% H2 excess rate (4000 h<sup>−</sup>1, 4 bar), product gas composition in vol.-% dry.

## *3.3. Extended Experiment Duration with Bottled Real BFG*

After the initial tests, a first long-term experiment was performed with bottled real BFG and a hydrogen excess of σH2 = 1.05 to further assess the degree of catalyst deactivation. Therefore, base case conditions and parameters were used. It needs to be mentioned that, due to the catalyst deactivation discovered in the previous experiments, only one reactor (R1) was used this time, in order to spare the catalyst in the remaining two reactors (R2 and R3).

Figure 9 shows the results of this extended experiment duration. The product gas composition for the four gases H2, CH4, CO2 and CO is shown in vol.-% dry on the *y*-axis, and the *x*-axis shows the duration in hours. Again, a decrease in the overall performance

of the catalyst is observed, starting at a higher rate at the beginning that stabilises after 16 min. Over 16.5 h, the CH4 concentration in the product gas dropped from a starting value of ~29 to ~16 vol.%, indicating the already-mentioned deactivation of the catalyst over time. This is equivalent to a drop in COx conversion of ~30% (1.8% per hour).

**Figure 9.** Time-based data for extended experiment duration with real BFG (4000 h−1, 4 bar, 5% H2-excess) over 16.8 h.

Figure 10 shows the temperature measurements taken inside the catalyst bed of the first reactor (R1) at the beginning of the experiments under synthetic gas conditions (reference base case), and at the end of the last experiment with real gases. The operating conditions are the same (4000 h−1, 4 bar, 5% H2-excess). A clear shift of the typical bellshaped temperature curve towards the top of the catalyst bed can be seen. This confirms that the catalyst at the bottom of the first reactor was deactivated by the poisons present in the bottled real gases. This behaviour was not seen for any experiments with synthetic gas mixtures under steady state or dynamic conditions [9,27].

**Figure 10.** Comparison of reactor R1 temperature profiles at start (red) and end (blue) of long-term experiments with real BFG.

## *3.4. Analysis of Gas Condensate and Catalyst*

Table 4 compares the analysis results of the gas condensate taken after the experiments with real gases with the ones for the synthetic gas mixture. The parameters have been measured according to the norms ISO 10304-1, 17294-2 and 10523.


**Table 4.** Gas condensate analysis for BFG (real gases vs. synthetic gas mixture).

The Ni content in the real gases' condensate is more than three times higher compared to the one for the synthetic gas mixture from the reference base case. This confirms the theory that more Ni atoms are taken off the catalyst with real gases and blown out of the reactor setup with the product gas, resp., leaving it through the condensate, indicating a mechanical deactivation through attrition [28]. The decrease in available Ni atoms on the surface of the catalyst can be another reason for the performance drop observed, resulting in a further deactivation of the catalyst. The other parameters relevant for catalyst deactivation (Cl, SO4, S) were all below the determination limit of the selected analysis methods.

Once the experiments were performed, the possibility to reactivate the catalyst inside the first reactor was evaluated. Therefore, the reactor was purged with pure hydrogen for a duration of 4 h with a GHSV of 2000 h<sup>−</sup>1, keeping a constant reactor pressure of 4 bar and temperatures above 260 ◦C. Confirmation experiments afterwards showed no significant improvement of the catalyst's performance for the methanation synthesis. Consequently, the whole methanation test rig was flooded with N2 to clean the piping and reactors. Afterwards, the catalyst was deactivated with compressed air and withdrawn from the reactors. A sample of the catalyst spheres inside the first reactor was analysed in a scanning electron microscope (SEM) and compared with a new catalyst sample (Figure 11). Just by comparing the two samples visually, a clear change in the surface structure and morphology can be noted. The surface of the used catalyst (right) does not show the crystalline structure of the fresh catalyst anymore (left). SEM reference analyses of used catalyst spheres after 96 h of methanation with synthetic steelworks gases revealed that the surface structure matches the one of the fresh catalyst rather than the one of the real gas experiments, which goes hand in hand with the observation that there was no significant degradation in the overall performance detected during these experiments [9,27]. Due to the high methanation temperatures (up to 600 ◦C) in the first reactor, thermal sintering is certainly a method of catalyst deactivation that needs to be adressed. Again, the experiments with synthetic steelworks gases and without poisons did not show any signs of a decreased catalyst performance during the experimental campaigns, although maximum reactor temperatures at a similar level were measured. Compared to the real gas experiments, no other parameter than the presence of catalyst poisons was changed. Consequently, this degradation can be clearly attributed to their presence in the real gases.

**Figure 11.** SEM pictures of new (**left**) and used (**right**) bulk catalyst sphere.

## *3.5. Implementation of Activated Carbon Filter*

Based on the results achieved with the direct use of the bottled real gases, including the listed catalyst poisons, which resulted in a quick catalyst degradation, an activated carbon filter was implemented upstream of the first reactor (R1). For this purpose, metal oxide impregnated activated carbon pellets were added to the gas-mixing station. These pellets are specifically developed to remove hydrogen sulphide, organic mercaptans, sulphur dioxide, carbonyl sulphide and nitrogen oxides from oxygen-deficient gas streams, such as CO2, N2, CO and H2. They have a copper content of 7% (as CuO), a diameter of 3 mm, a length of 7 mm, a specific surface area of 936 m2/g, a pore volume of 0.53 cm3/g and a bed density of 0.48 g/cm<sup>3</sup> (Figure 12). A total of ~300 mL (187 g) of these pellets was added to the gas mixing station (ID 36 mm, height 300 mm).

**Figure 12.** Activated carbon pellets with copper oxide coating.

In addition to the implemented filter based on activated carbon, a fresh Ni-based catalyst was added to the reactors. Furthermore, the piping and fittings were exchanged to assure no catalyst poisons remained inside the plant. To test the functionality of the implemented gas cleaning stage with activated carbon, only one reactor was used in the beginning under base case operating parameters. This decision was made to limit the exposure of the plant's components to the real BFG as much as possible, in case of failure of the activated carbon solution.

Figure 13 shows the time-based measurement data for the real gas experiments with one reactor and base case operating conditions. The data include four days of methanation, as well as their individual start-up and shut down phases. Methanation during the first two days (~6 h each) showed very constant measurement values for all gas components on the product side. The actual values vary in a range of ±2%-points around their averages (Table 5) and are within a 1.5%-point range per gas component compared to reference experiments with synthetic BFG, with the same composition and operating conditions. The small fluctuations and peaks, respectively, for H2 and CO2, are only process-related due to the measurement technique of the infrared gas analytic station. In addition, the temperature profile is within a ±1 ◦C range for all temperatures during day one and in a ±2 ◦C range for all temperatures except TI2 during day two. The temperature at the bottom of the catalyst bulk (TI2) started to decrease after approximately two hours in day two and dropped by 11.1 ◦C during the remaining four hours of the experiment time. This drop in temperature did not have any influence on the overall catalyst performance, as the COx conversion and product gas composition remained stable.

**Figure 13.** Time-based data for extended experiment duration with real BFG and activated carbon filter (4000 h<sup>−</sup>1, 4 bar, 5% H2-excess) over 4 days (25.2 h net).

**Table 5.** Comparison of average product gas composition of methanation with real BFG on day one and synthetic BFG (BFGsyn); values in vol.-%, activated carbon filter used for real gases.


On day three, the temperature at measurement point TI2 continued its downwards trend with a rate of 38 ◦C/h. Furthermore, TI3 and TI4 started to decrease by 21.6 ◦C/h and 8.8 ◦C/h. respectively, whereas TI5 and TI6 remained stable. Even though the temperatures at the bottom and middle section of the catalyst bulk decreased by 95 and 22 ◦C in total, the product gas composition remained almost constant. On day four, all temperatures decreased further, except the one for T6, which started to increase as the upper part of the catalyst took over the synthesis load from the catalyst in the lower section. After a

net duration of 25.2 h of methanation with real gases, the experiment was stopped, as the temperature at TI2 came close to the lower boundary of 200 ◦C, below which, poisonous nickel carbonyl is formed [21].

An analysis of the temperature profile along the reactor proves that there is no drop in catalyst performance noticeable on day one, as the red and grey lines in Figure 14 are almost identical. Day two shows the reported minor shift at the bottom of the reactor towards cooler temperatures (yellow line). With day three and four (green and blue lines), the shift towards the left for the lower half of the catalyst bulk is clearly noticeable. At TI5, the temperature remained almost constant during all four days. At the top, the reported temperature increase during the days three and four can be seen, as well as a shift in the hotspot temperature from TI3 to TI5/TI6 over time, clearly indicating the loading of the installed adsorbent bed, as well as a breakthrough of catalyst poisons downstream towards reactor R1.

**Figure 14.** Comparison of reactor R1 temperature profiles at the start of day one, and at the end of each experiment day; activated carbon filter used for real BFG.

Further to the comparison of the temperature profile, reference measurements were also performed prior to and after the real gas experiments on day one and day two. For these, a synthetic gas mixture consisting of hydrogen and CO2 (H2:CO2 = 8:1.5) was used to evaluate if any drop in performance can be recognised. Table 6 lists the product gas composition for these reference measurements. Ref. #1 was taken prior to the real gas experiments, Ref. #2 after day one and Ref. #3 at the end of day two. When exclusively addressing the product gas composition, there is no sign for catalyst deactivation observed until the end of day two.

**Table 6.** Product gas composition of reference measurements taken with synthetic H2/CO2 gas mixture (values in vol.-% dry).


Nevertheless, as the temperature profile in Figures 13 and 14 indicates, the first sign of a breakthrough of catalyst poisons occurred on day two after approximately two hours of experiments. This results in a service life of 7.2 h for the implemented adsorbent and is exclusively based on analysing the temperature profile measured inside the reactors. During this time, a total of 7.3 m3 of real gases flowed through the adsorbent bed, with a volume of 300 mL. Considering the product gas composition, the first catalyst deactivation can be observed after 12.4 h (~12.4 m<sup>3</sup> of real gases). This results in a required amount of

absorbent material of 15.1 g per hour of operation with base case parameters, or 15.0 g/m<sup>3</sup> of feed gas, respectively.

All of the information obtained from the experiments in lab-scale can be used to estimate and design a pilot plant for the methanation of real gases, including an additional gas cleaning step based on activated carbon. Medved et al. [9] described several application scenarios for the direct methanation of steel gases. One scenario is the substitution of the plant internal natural gas demand with SNG through the methanation of blast furnace gas (BFG). Such a scenario includes approximately 57,000 Nm3/h of BFG and 100,000 Nm3/h of hydrogen. With the consumption numbers for activated carbon obtained in this work, this would result in the need of ~2.4 t of adsorbent per operating hour. Furthermore, Calbry-Muzyka et al. [18] estimated the capacity of the required adsorbent by calculating the integrated loading of H2S in the adsorbent bed during operation with biogas. With Equation (7), the loading can also be calculated for other catalyst poisons, such as the ones present in the real blast furnace gas. The following parameters are required as an input:


$$\text{Loading } [\text{wt.} - \%] = \frac{\sum \text{C}\_{\text{cat\\_poison}} \times \text{Q} \times \text{M}\_{\text{cat\\_poison}} \times \text{t}}{\text{m}\_{\text{adsorbert}}} \tag{7}$$

Table 7 shows the results for a total flow rate of 1.01 m3/h (representing base case parameters), 187 g of adsorbent bed and 12.4 operating hours. This experiment duration was selected as it flags the first sign of catalyst deactivation based on monitoring the product gas composition. Consequently, it also shows a breakthrough of catalyst poisons through the adsorbent bed, which is sufficient for an effect on the catalyst's activity. The loading is below 1.0 wt.-% for each type of catalyst poison measured, with the highest values for COS, followed by H2S.

**Table 7.** Loading of adsorbent bed in wt.-% per catalyst poison during methanation with base case parameters over 12.4 h of operation.


Neubert [29] used a different approach for describing the adsorption of catalyst poisons by comparing the integral change in the axial temperature profiles between two experiments with the one obtained with a fresh catalyst. The author used Equation (8) to calculate the relative activity loss (Δactivity in %) per experiment, which can be further related to the runtime of an experiment (Equation (9)). As a parameter for upscaling, the catalyst consumption (Δmcatalyst) due to sulphur-based catalyst poisons (ni) can be calculated. This term is multiplied by the amount of fresh catalyst mass (mcatalyst,0) used in the methanation reactor (Equation (10)).

$$\text{Aactivity} = \frac{\int\_0^{\text{z}} \text{T}\_{\text{m}}(\text{z}) - \text{T}\_{\text{n}}(\text{z}) \text{dz}}{\int\_0^{7.2} (\text{T}\_{\text{l}}(\text{z}) - \text{T}\_{\text{l}}) \text{dz}} \tag{8}$$

$$
\Delta \text{activity/h} = \frac{\Delta \text{activity}}{\Delta \text{t}} \tag{9}
$$

$$
\Delta \mathbf{m}\_{\text{catalyst}} = \frac{\Delta \text{activity}}{\sum \mathbf{n}\_{\text{i}}} \times \mathbf{m}\_{\text{catalyst,0}} \tag{10}
$$

The results for the four days of experiments with real BFG are shown in Figure 15, as well as in Table 8. During the first two days, the catalyst activity decreased only by less than 3.0%, with a very small catalyst consumption due to sulphur-based catalyst poisons. On day three and four, the catalyst consumption increased to 6.8 and 3.8 gcatalyst/mmolsulphur, with a combined activity loss of close to 45%. This again shows that, without gas cleaning, respectively, with a fully loaded adsorbent bed, catalyst degradation takes place at an enormous rate. Furthermore, the higher values for Δactivity and Δmcatalyst on day three compared to day four show that an instant, high drop in performance takes place as soon as poisons break through the adsorbent bed, after which, the decrease stabilises at a lower rate. This behaviour was also observed during the initial tests with real blast furnace gas.

**Figure 15.** Axial shift of temperature profiles for Δactivity calculation per experiment day (**left**); relative activity loss per hour and catalyst consumption per mmol of sulphur (**right**).

**Table 8.** Relative activity loss (Δactivity) in % and in %/h, as well as catalyst consumption (Δmcatalyst) in gcat/mmolsulphur and gcat/m3 for experiments with real BFG (4 days); activated carbon in place.


When upscaling the figures obtained for day one and day two for the real application scenario in the steelworks, this would result in the deactivation of 88.1 kg of catalyst per hour of operation. Considering the catalyst consumption of the evaluated 12.4 operating hours with a rate of 1.75 gcat/m3 and upscaling them for the real case, 275 kgcat would be deactivated per hour.

In addition to the experiments with BFG, the upgraded methanation setup, including the activated carbon filter, was also exposed to real gases from a converter (BOFG), with a composition according to Table 2. Again, the required hydrogen to achieve a σH2 of 1.05 was added through gas bottles. During the experiments with base case parameters and one reactor, no catalyst deactivation was detected over a period of 16 h. As the contaminants of the BOFG are far fewer compared to the ones of BFG (compare Table 3), this is a logical result and shows that the implemented solution works for both gas types. Table 9 compares the product gas composition of methanation with real BOFG and synthetic BOFG. As with blast furnace gas, they are again very similar, varying within 0.5 to 3%-points depending on the gas component.

**Table 9.** Comparison of average product gas composition of methanation with real and synthetic BOFG (values in vol.-%).


The temperature profile shown in Figure 16, as well as the analysis of the gas condensate, proves that there is no catalyst deactivation noticeable. The temperatures measured inside the reactor at the start and end of the real gas experiments are almost identical and follow the same bell-shaped profile, with a hot spot at measurement point TI4 (middle of the catalyst bulk). In addition, the amount of Ni atoms measured in the gas condensate sample is the same for synthetic and real gases (22 μg/L).

**Figure 16.** Comparison of reactor R1 temperature profiles at start (red) and end (blue) of experiments; activated carbon filter used for real BOFG.

## **4. Conclusions and Outlook**

In this work, methanation experiments with real gases from the steelworks industry have been performed. These experiments included real by-product gases from the blast furnace, as well as from the basic oxygen furnace (converter), that were directly bottled at an integrated steel plant during normal operation. No further treatment, such as a CO2 or N2 separation step, was performed prior to the filling procedure. Methanation without additional gas cleaning resulted in an instant, as well as steady, catalyst degradation due to the poisons present in the real gases. Over the evaluated periods, a drop in the COx conversion of ~30%, or 1.8% per hour, was detected for blast furnace gas. The usage of unfiltered real gases resulted in a 3.2 times higher amount of Ni transported out of the reactor setup with the condensate, compared to experiments with synthetic gases meeting the same composition and operating conditions.

As a working solution, an activated carbon filter coated with copper oxide was implemented, showing that a further pre-treatment of the already cleaned steelworks gases is essential prior to feeding them to a catalytic methanation plant. With the activated carbon filter in place, methanation experiments could be performed without any noticeable degradation until the 187 g of activated carbon adsorbent were fully loaded with catalyst poisons. The first signs of deactivation appeared after 7.2 h of operation with real BFG, by means of a drop in the reactor temperature measured at the bottom of the catalyst bulk. Over another period of 5.2 h, the product gas composition and overall conversion rate remained constant, after which, the methane yield started to drop. During this period of 12.4 h, on average, 15.1 g of adsorbent was consumed per hour of operation. The loading of catalyst poisons within the adsorbent bed stayed within a range of 0.01 to 0.73 wt.-% depending on the type of poison. While continuing methanation over another 6.7 h, the catalyst consumption increased from 0.4 to 4.8 gcat/mmolsulphur on average, and the relative activity of the catalyst decreased by ~45% compared to its starting performance. In the case of real BOFG, no signs of catalyst deactivation could be observed during the course of the experiments, which is a result of the far lower catalyst poisons present in this type of gas.

For a real application-based scenario in an integrated steelworks, with the target to substitute the demand of any externally sourced natural gas with a plant-internally produced SNG through methanation, the figures obtained through the experiments at lab-scale would result in the need of ~2.4 t of adsorbent and a deactivation of 88.1 kg of catalyst per hour of operation.

Future work will show the usability of the methanation setup for dynamic experiments as they occur in a steelworks plant, including frequent load changes of up to 25% of gas input power in the range of 5 to 45 min simulating a dynamically operated PEM electrolysing unit. The values obtained through the lab-scaled experiments will also assist in the technical design of a pilot plant for the steelworks industry, including a gas cleaning step based on activated carbon prior to feeding the real gases to the methanation units. Furthermore, the catalyst degradation due to poisons present in the real gases will be investigated in detail through Raman spectroscopy, as well as BET analysis.

**Author Contributions:** Conceptualization, P.W.-Z. and A.R.M.; Formal analysis, P.W.-Z.; Methodology, P.W.-Z. and A.R.M.; Validation, P.W.-Z., A.R.M., M.L. and N.K.; Visualization, P.W.-Z.; Writing–original draft, P.W.-Z.; Writing–review & editing, P.W.-Z., A.R.M., M.L., N.K. and K.R. All authors have read and agreed to the published version of the manuscript.

**Funding:** This research was funded by the Research Fund for Coal and Steel RFCS by the EU Commission, grant number 800659 i3upgrade https://www.i3upgrade.eu/ (accessed on 18 August 2021).

**Data Availability Statement:** Not applicable.

**Acknowledgments:** The experiments of this work were conducted as part of the research project "i3upgrade—intelligent, integrated, industries", funded by the European Commission. Besides Montanuniversität Leoben, the following research institutes were involved: Chair of Energy Process Engineering (EVT) and Institute of Chemical Reaction Engineering (CRT) at Friedrich-Alexander University Erlangen-Nürnberg, Germany; Central Mining Institute (GIG) in Katowice, Poland; Institute of Communication Information and Perception Technologies (TeCIP) of Scuola Superiore Sant'Anna (SSSA) in Pisa, Italy; and the Centre for Research and Technology Hellas (CERTH), Thessaloniki, Greece; with the industrial partners AIR LIQUIDE Forschung und Entwicklung GmbH (ALFE), voestalpine Stahl GmbH (VAS) and K1-MET GmbH. The authors would also like to acknowledge the work of Hanna Weiss during her bachelor's studies.

**Conflicts of Interest:** The authors declare no conflict of interest. This paper reflects only the author's view, and the founding sponsors had no role in the design of the study, in the collection, analyses, or interpretation of data, in the writing of the manuscript and in the decision to publish the results. The European Commission is not responsible for any use that may be made of the information contained therein.

## **Abbreviations**



## **References**

