Simulation of Conventional WWTPs Acting as Mediators in H₂/CO₂ Conversion into Methane

Abstract

CO₂-biomethanation was studied in the present manuscript by considering the direct injection of hydrogen into a conventional anaerobic digester treating sewage sludge within a simulated wastewater treatment plant (WWTP). The plant was simulated using the Python 3.12.4 software, and a Monte Carlo simulation was conducted to account for the high variability in the organic content of the wastewater and the methane potential of the sludge. Two modes of operation were studied. The first mode involves the use of an anaerobic digester to upgrade biogas, and the second mode considers using the digester as a CO₂ utilization unit, transforming captured CO₂. Upgrading biogas and utilizing the extra methane to generate electricity within the same plant leads to a negative economic balance (first scenario). A hydrogen injection of 1 L of H₂/L_r d (volumetric H₂ injection per liter of reactor per day) was required to transform the CO₂ present in the biogas into methane. The benefits associated with this approach resulted in lower savings regarding heat recovery from the electrolyzer, increased electricity production, and an additional oxygen supply for the waste-activated sludge treatment system. Increasing the injection rate to values of 5 and 30 L of H₂/L_r d was also studied by considering the operation of the digester under thermophilic conditions. The latter assumptions benefited from the better economy of scale associated with larger installations. They allowed for enough savings to be obtained in terms of the fuel demand for sludge drying, in addition to the previous categories analyzed in the biogas upgrading case. However, the current electricity price makes the proposal unfeasible unless a lower price is set for hydrogen generation. A standard electricity price of 7.6 c€/kWh was assumed for the analysis, but the specific operation of producing hydrogen required a price below 3.0 c€/kWh to achieve profitability.

1. Introduction

The biological treatment of municipal wastewater typically involves the conventional activated sludge process, where aerobic microorganisms break down organic material, producing a microbial biomass, which is subsequently removed through physical separation. The larger the scale of the treatment plant, the greater the amount of biological material that requires suitable disposal. Many large-scale plants commonly treat primary and waste-activated sludge through anaerobic digestion, producing biogas as an energy product and a slurry referred to as biosolids, which often serves as an organic amendment when the material complies with the country’s regulations. The land application of biosolids provides the benefit of recycling nutrients (nitrogen and phosphorus), improves soil quality, and avoids the depletion of organic carbon from soils, a feature attained at modest expenses [1,2,3,4].

The digestion of sewage sludges may be carried out under mesophilic or thermophilic conditions. In the case of digesters working under mesophilic conditions, the degradation of organics takes place at a slower pace because of the effect of temperature on kinetics. Despite the benefits of a high degradation rate in the thermophilic regime, they are insufficient to tilt the balance in favor of this process due to the lower quality of thermophilic digestate and the poor properties of the rejected supernatant [5]. In addition, no significant differences in biogas production have been reported when operating digesters under mesophilic and thermophilic conditions [6]; therefore, the increase in temperature has not always resulted in higher gas production [7,8]. For these reasons, many wastewater treatment plants (WWTPs) still operate under mesophilic conditions and typically implement other options for increasing productivity, such as co-digestion. The addition of organic waste increases the loading and aids in achieving a better balance of nutrients, which may enhance biogas formation by 13–176% [9,10,11].

Given the high energy demands of the waste-activated process, treating sewage sludge by anaerobic digestion provides the dual benefit of reducing its volatile content and generating some of the plant’s energy needs. Therefore, increasing digestion productivity is essential for improving the energy balance. García-Cascallana et al. [12] reported that if a sufficient digestion capacity was available, the biogas generated from co-digestion could even cover the full energy demand of the plant. Finding a suitable co-substrate all year round, without dealing with odor or the discomfort from waste handling operations, is often challenging. In addition, the increase in organic loading also causes an unavoidable increase in digested sludge, along with other unexpected outcomes such as solid accumulation inside the reactor, a nitrogen backload, and lower dewaterability [13].

A completely different strategy for increasing methane production may involve using hydrogen as a co-substrate. CO₂ is transformed by hydrogenotrophic methanogens into methane, requiring four moles of hydrogen (H₂). CO₂-biomethanation has garnered the attention of the scientific community due to the ease of adaptation of anaerobic microflora and the broad application of digestion technology operating at an industrial scale. Several researchers have reported on the experimental performance and technical feasibility of the approach [14,15,16]. The strategy of increasing methane productivity by recirculating biogas has been reported by Poggio et al. [17], indicating that H₂ gas transfer limitations were reduced by attaining higher circulating rates and increasing gas residence time. In situ biogas upgrading can significantly reduce energy consumption by utilizing endogenous CO₂, resulting in a methane content that is compatible with the natural gas grid [18]. Martínez et al. [19] tested the conversion of H₂ in anaerobic reactors treating sewage sludge at injection rates of 0.5–2.0 L of H₂/L_r d. These authors reported an increase in biogas production but not in composition. After analyzing the microbial population, the reactor performance was explained by the conversion of CO₂ into acetate, which was subsequently converted into methane. Nguyen et al. [20] tested hydrogen injection rates of 4.39 L of H₂/d (1.0 L of H₂/L_r d), achieving a H₂ utilization efficiency of 92–99% with a methane composition up to 92%.

Large-scale anaerobic digesters may play a role in the hydrogen economy by utilizing the captured CO₂ to produce methane, thereby serving as units for energy storage [21]. Hydrogen derived from water electrolyzers can be integrated into existing large-scale digesters without negatively impacting performance. This approach combines two benefits. The first is upgrading biogas, and the second is using excess renewable energy, which becomes available when solar and wind power account for a high share of the energy mix. Methane can be used as fuel in different sectors, such as industry and transportation. Thus, wastewater treatment plants (WWTPs) can significantly contribute to the shift towards a decarbonized economy by supplying renewable fuel that helps reduce the dependency on fossil fuels [22]. The decision to phase out coal and nuclear generation and increase the share of renewables has created a phenomenon where negative prices are more frequent in the energy spot market due to excess energy available when demand is low [23]. In addition, the Spanish shutdown in April 2025 underscored the importance of maintaining a reliable grid system with a sufficient energy storage capacity to ensure the continuity of essential services.

One way of storing this extra energy is by transforming electricity into fuel such as methane. However, alternatives like storing energy in batteries as well as potential, kinetic, or thermal energy are also possible. The conversion into chemicals seems to be the best option for long-term storage [24]. This goal can also be achieved by using catalysts in a process known as the Sabatier reaction, where noble metals (Ru, Rh, Pd, and Pt) are required to catalyze the conversion of CO₂ at high pressures and temperatures between 300 and 400 °C [25,26]. The feasibility of this approach depends on the price of catalysts, the cost of producing hydrogen, and the capture of CO₂ [27]. Recent work has focused on the development of new nickel-based catalysts and bimetallic catalysts [28,29] as well as the simultaneous capture and transformation of CO₂ into methane by combining adsorbents and metallic catalysts. This allows for the simplification of the capture and transformation stages by having these take place in a single reactor using dual-functional materials [30,31,32]. Another interesting route for transforming CO2 is electrochemical reduction. However, the high stability of the CO₂ molecule (with the C=O bond having a dissociation energy of 750 kJ/mol) translates into the need to provide a large overpotential to activate the reaction, leading to low energy efficiencies [33,34].

On the contrary, biological systems may also attain the transformation of CO₂ into methane with the added advantage of milder reaction conditions. This approach offers several benefits, as previously mentioned, in addition to the advantage of utilizing existing large-scale anaerobic digesters. It leverages the experience gained from operating standard digestion units and transforms these systems into a process capable of storing energy as methane. The application of milder conditions and the use of mixed microflora make biological methanation a promising technology due to its ease of implementation. The use of existing reactors reduces the financial burden associated with installation and operating costs, as nutrient supply and specific inoculation are no longer necessary. This is one of the arguments put forward by some authors for relegating biological methanation to small-scale applications [35,36,37].

The idea of transforming C1 gases dates back to the 1990s. It has been previously studied not only as a method for upgrading biogas but also for transforming the H₂/CO/CO₂ components of syngas to produce methane [38,39,40]. The interest in this technology has experienced a resurgence due to concerns about greenhouse gas (GHG) emissions and the goals associated with achieving climate neutrality by 2050, which are at the heart of the European Green Deal [41]. Efforts to increase fermentation efficiency are associated with reducing mass transfer limitations by favoring the gas–liquid interphase, operating at lower temperatures, increasing pressure, and increasing the biomass concentration in the reactor via immobilization [16,42,43]. Other approaches include the use of biocathodes, the introduction of electrodes (electro-fermentation), and operating under thermophilic conditions using mixed cultures [44,45,46]. Nevertheless, the low development of these recent proposals and the lack of enough experience at a pilot scale make the direct injection of a gas phase into an anaerobic operating digester the most feasible option in the short term.

The present manuscript studies the feasibility of using anaerobic digesters working in WWTPs as biological units for producing extra methane thanks to CO₂-biomethanation. The novelty of the present contribution is assessing the feasibility of using anaerobic digesters in WWTPs as reactor units for either upgrading biogas or as CO₂-conversion systems. Different hydrogen injection rates were established, assuming the use of water electrolyzers, with the digestion unit’s main aim being the conversion of H₂/CO₂ mixtures into methane. The main parameters to attain profitability were assessed and the electricity price required to attain profitability was estimated.

2. Materials and Methods

The WWTP model was based on Ellacuriaga et al. [47] as well as González et al. [48] and used the assumptions of Martínez et al. [19].

Table 1. Main model parameters used in the WWTP flow calculations.

Parameter	Value	References
Number of equivalent inhabitants	150,000	[19]
Specific wastewater production (L/inhab. d)	330	[49]
Percentage of water removed with particle separation unit at WWTP inlet	2%	[47]
Biomass yield (WAS 1 process)	0.6	[47]
Volumetric air supplied to WAS (m3 air/m3reactor min)	0.025	Based on SuperPro designer model assumption
Power WAS process (kW/m3reactor)	0.3	Based on SuperPro designer model assumption
WAS stream recirculation	35%	[47]
Primary sludge total solid (TS) content (g/L)	60 ± 12
Percentage of volatile solids (%VS) primary sludge	75 ± 15
Secondary sludge TS content (g/L)	45 ± 9
(%VS) secondary sludge	65 ± 13
Organic matter (COD mg/L) 2	760 ± 152	[50]
Organic matter (BOD mg/L) (50% of COD value)	380 ± 77	[51]
Biochemical methane potential (BMP) (Average value from references)	300 ± 73	[52,53,54]
Digester maximum volume (m3)	4000	[47]
Digester diameter–height ratio	1:2
Digester free head space (%)	25
Biogas methane content (%)	60
Methane LHV (MJ/m3)	35.8

¹ WAS: Waste-activated sludge process. ² Average value reported for Spain.

shows the list of main model assumptions, considering a 20% variation in sludge solid content and volatile solid composition. The WWTP process assumptions were based on the SuperPro designer model used in Ellacuriaga et al. [47].

The digester’s thermal demand was calculated using the equations described by González et al. [55]. The thermal demand considered the heat required to increase the sludge temperature from the inlet stream (15 °C in summer conditions and 5 °C in winter conditions) to the fermentation temperature (37 °C under mesophilic and 55 °C under thermophilic conditions), assuming a 95% heat transfer efficiency, 5% of heat losses in the summer, and 10% of losses in the winter period. A hydraulic retention time of 21 days was used for dimensioning the digester volume.

The digestate was subjected to dehydration using horizontal decanter centrifuges. The solid content of dehydrated sludge was 22.0 ± 4.4%. The land application of the dehydrated digestate for agronomic purposes was assumed. Sludge was transported to a nearby site located 30 km away (tortuosity factor of 1.4). A truck with a load capacity of 40 m³ and a fuel (diesel) consumption of 35 L/100 km was used to estimate the transport energy demand [56]. The lower heating value (LHV) and density of diesel were 44.8 MJ/kg and 0.84 kg/L, respectively [57,58]. The price of truck renting was 1.25 €/km with a diesel fuel price of 1.6 €/L. A combined heat and power (CHP) engine was considered with a maximum electrical efficiency of 39.7% and a maximum thermal efficiency of 52% for an electrical output of a 249–330 kW range [59]. The thermal exhaust gas temperature was assumed to be 640 °C, which allowed for the recovery of 50% of the CHP exhaust gases as heat [56,60]. The natural gas price was 45 €/MWh for the period from 1 November 2024 to 30 April 2025 [61].

Supplementary Material S1 shows a full description of the WWTP model used. A Monte Carlo simulation was performed to take into account the variability of the sludge and wastewater composition. A normal distribution was assumed for the values reported in Table 1. The Python 3.12.4 software with the ‘rng. normal’ command was used for modeling the process, running 10,000 simulations.

2.1. CO₂-Biomethanation

The biological methanation of CO₂ involves the following reactions, as proposed by Schwede et al. [62] and Rafrafi et al. [35]. The conversion of CO₂ into acetate was demonstrated to be dependent on the dissolved CO₂ levels in the reactor, with values above 2.0 ± 0.2 mmol/L favoring direct methane conversion from CO₂ by keeping H₂/CO₂ ratios above 4.0 units [63]; therefore, in the present document, only reaction (1) was considered to estimate the methane productivity of the reactor:

CO₂ + 4 H₂ → CH₄ + 2 H₂O

(1)

$$2 {\mathrm{CO}}_2 + 4 {\mathrm{H}}_2 \leftrightarrow {\mathrm{CH}}_3\mathrm{COOH} + 2 {\mathrm{H}}_2\mathrm{O}$$

(2)

CH₃COOH → CH₄ + CO₂

(3)

Continuous stirred tank reactors (CSTR) attain better specific surface areas than other types of reactors if the gas bubbles are small and well dispersed, thus allowing better transfer efficiencies (higher K_La values) [64]. However, other reactor configurations, such as airlift reactors, perform better in terms of their mass transfer efficiency [65].

Table 2. Main operating parameters and performance of anaerobic reactors working as CO₂-methanation units.

Experiment Characteristics	Performance	Reference
CSTR 1 with 5 L of working volume. Mesophilic regime (37 °C). Injection flow: 1.5–7.2 mL/min (0.43–2.07 L of H2/Lr d). Evaluation period of 151 d	28 days were enough for attaining microflora enrichment. H2 conversion rates were above 80%. Long-term and stable CH4 production were attained without any extra addition of nutrients during the 5-month experimental period.	[66]
CSTR with 3.5 L of working volume. Mesophilic regime (37 °C). Cattle manure used as substrate (HRT 2: 35 d, OLR 3: 0.5 g of VS/Lr d). The reactor was operated to transform CO2 contained in the biogas and subsequently transform extra CO2 injected into the system.	The injection rate increased from 0.03 to 0.37 L of H2/Lr d when operating under a biogas upgrading configuration. The volumetric CH4 production rate increased from 117.4 ± 8.6 mL/Lr d to 985 ± mL/Lr d when changing operation from the basic manure digestion condition to external CO2 addition.	[67]
Anaerobic filter (130 L of reactor volume; the fixed bed had a volume of 58 L). The reactor used a biomass hydrolysate with an OLR of 3.0 g of COD/Lr d. Mesophilic regime (37 °C). Injection flow: 2.97 and 5.98 L of H2/h (0.54 and 1.1 L of H2/Lr d). Evaluation period of 131 d.	H2 conversion was below 80%. The gas residence time in the reactor was between 18 and 50 h. Stable hydrogen conversion during the whole evaluation period.	[68]
Bubble column reactor with 20 L of working volume. Thermophilic regime (55 °C). Mesophilic inoculum was used for starting up the reactor. Injection flow: 0.67–9.4 L of H2/Lr d. Evaluation period of 405 d.	H2 and CO2 conversion exceeded 90%, with great resilience to starvation periods and temperature drops (25 °C under N2 atmosphere).	[69]
Trickle bed reactor (73 L of reactor volume) with 58 L of packed bed. Thermophilic regime (55 °C). Injection flow: 1.7–62.1 L of H2/Lbed d). Evaluation period of 313 d.	The maximum CH4 production rate was 15.4 L/Lbed d at a 98% CO2 conversion.	[70]
Trickle bed reactor (73 L of reactor volume) with 58 L of packed bed. Thermophilic regime (55 °C). Stepwise increase in H2 injection from 4.5 to 52.5 L of H2/Lbed d for each restart. Evaluation period of 240 d.	Focusing on repetitive standby (25 °C, no H2 feed)/restart (55 °C, H2 feed) operation with the aim of using the reactor for energy storage from intermittent operating renewable systems. Standby periods adversely affect subsequent performance.	[71]
Inverse membrane reactor. Two membrane modules connected in series (0.25 m2 of membrane area, 0.6 L of retentate and 0.31 m2 of membrane area, 0.65 L of retentate). Mesophilic regime (35 °C). Injection flow: 9.6 L H2/LLiquid d. Evaluation period of 412 d.	Maximum methane production was 83 L CH4/m2 membrane area (2.28 L of CH4/LLiquid d). Methane concentration was greater than 90%.	[72]
CSTR with 2 m3 of working volume and 3.4 m3 of headspace volume. Thermophilic regime (55 °C). Sewage sludge digestion (HRT: 20 d, OLR: 1.0–1.1 g of VS/Lr d) Injection flow: 0.12–0.67 L of H2/Lr d. Evaluation period of 140 d.	The pilot experiment with a working volume of 2 L showed higher H2 conversion capacity, assimilating a H2 flow of 0.67 L H2/Lr d, whereas the pilot-scale system tolerated a flow of 0.37 L H2/Lr d. The pilot experiment presented a lower abundance of acetoclastic methanogens.	[73]
CSTR with 5 L and 4.5 L for the thermophilic regime and 4 L and 3.6 L for the mesophilic regime. Evaluation period of 400 d.	Thermophilic reactors showed better utilizations of H2 gas, although the adaptation period was longer. Maximum H2 assimilation was 2 L of H2/Lr d for a biomass concentration of 4 g/L.	[74]

¹ CSTR: Continuously stirred tank reactor. ² HRT: Hydraulic retention time. ³ OLR: Organic loading rate.

shows a list of the different experimental work found in the literature reporting on the conversion of H₂/CO₂ to produce methane.

The hydrogen injection values reported in the scientific literature range from 0.02 up to 0.56 L of H₂/L_r d [66,67] when studying direct injection into the anaerobic reactor and higher values when operating with other reactor configurations, thanks to the feasibility of applying higher gas recirculation rates and thus, improving mass transfer. Laguillaumie et al. [69] applied injection rates between 0.7 and 9.4 L of H₂/L_r d using a bubble column when working under thermophilic conditions and attained a gas conversion close to 100%. Illi et al. [68] reported values between 0.54 and 1.1 L of H₂/L_r d using an anaerobic filter under mesophilic conditions, and Strübing et al. [71] reported an injection rate of 52.5 m³ of H₂/m³ trickle bed/d operating in this case under thermophilic conditions. Haitz et al. [72] reported injection values of 4–6 L of H₂/L_r d when testing a hollow fiber system. Given the complexity of operating other reactor configurations, the present work assumed a direct hydrogen injection into the digester, thereby taking advantage of the nutrients already present in the sludge, operating the reactor as a typical digestion unit, and serving H₂ as a co-substrate.

Two operational modes were studied (see Figure 1), one involving mesophilic conditions for the anaerobic digester and thermophilic conditions for the other. The injection rate in the first case was assumed to be 0.2 L of H₂/L_r d for the low injection case and a value of 1 L of H₂/L_r d for the high injection case, based on the assumption that the maximum theoretical value was 2.27 L of H₂/L_r d [75]. The second operational mode evaluated higher injection rates by establishing thermophilic conditions. In this latter mode, two injection rates were also tested: 5 L of H₂/L_r d for the low injection case and 30 L of H₂/L_r d for the high injection case. No increase in sludge-specific methane production (SMP) was considered for the thermophilic conditions based on reports of Gavala et al. [6] and Chen et al. [8]. The H₂ injection assumptions yielded k_La values in the range of 0.23 and 0.32 h⁻¹ when working under the mesophilic regime and for the thermophilic case the range for k_La was between 9.0 and 42.0 h⁻¹. These values are in accordance with those reported by Liu et al. [76] when evaluating the k_La values for syngas fermentation in CSTRs.

2.2. Hydrogen Production from Water Electrolyzers

Alkaline electrolyzers are the most widespread technology due to their high level of maturity [77,78]. However, proton exchange membrane (PEM) electrolyzers allow for higher current densities, with the prices for these units expected to decrease significantly by 2030 [36,79]. The size of the electrolyzer is based on the hydrogen injection rates considered in the previous scenarios. The oxygen produced by the electrolyzer was assumed to be added to the air supply system of the conventional WAS unit, thus reducing the volume of fluid handled by the air compressor. The energy demand of the air compressor was 0.2 kWh/m³ of air. The electrolyzer specific energy demand was 4.3 kWh/m³ of H₂ with a heat production equivalent to 20% of the power and 80% of the heat recovery capacity [80,81]. The heat recovered from the electrolyzer is used to cover the digester’s thermal demand either under mesophilic or thermophilic conditions, depending on the working mode evaluated. Under the mesophilic regime, the biological reactor works as a biogas upgrading unit. Conversely, the thermophilic regime permits higher injection rates and utilizes the biological reactor as a treatment unit capable of transforming captured CO₂ (intensive CO₂ conversion system). The electrolyzer water demand was estimated by assuming a conversion factor of 9 L of ultrapure water being required for producing 1 kg of H₂ and 4.5 L of tap water being necessary for producing 1 L of ultrapure water.

The PEM electrolyzer and auxiliary equipment costs were 1337 (2020) $/kW using the 2020 CEPCI index of 596.2, and 2023 CEPCI index of 800.8 [82,83]. The US dollar to euro conversion was 1.14$, equivalent to 1 €. The operating and maintenance costs were 5% of the initial investment. The profitability of the approach was based on the savings attained through the extra methane available and the reduction in electrical demand resulting from the use of pure oxygen in the WAS unit. The methane derived from the digester was used to produce electricity in the CHP engine. The heat recovered from the CHP engine covered the thermal demand of the digestion system and, whenever possible, the energy demand for sludge drying. The time horizon of the economic assessment was 25 years, using linear depreciation with a 10% salvage value. The depreciation period was 15 years. The net present value (NPV) and payback period were used to estimate profitability, assuming a discount rate (r) of 3.5%.

$$NPV=-CI+{\sum }_{t=1}^n{\displaystyle \frac{CF}{{\left(1+r\right)}^t}}$$

(4)

where CI stands for capital investment; CF represents cash flow, which, in the present document, is derived from the savings and revenues obtained when the reactor operates as a capture-CO₂ utilization unit. In this latter case, the profitability was assessed by assuming revenues equivalent to 50 €/t CO₂ and 100 €/t CO₂. The present study focuses on assessing the feasibility of using existing anaerobic digesters as CO₂-conversion units. For this reason, the analysis does not consider the costs associated with carbon capture or the investment costs associated with reactor installation.

3. Results and Discussion

Figure 2 shows a scheme of the WWTP where sludge digestion is also represented. The primary and secondary sludge are treated in the anaerobic digester. The amount of biogas produced was 3914 ± 1347 m³ biogas/d based on a 60% methane concentration in the biogas. Two digesters with a mean size of 3130 m³ are needed for treating the sludge flow. The model equations considered winter and summer conditions for estimating the digester’s heat demand.

The main performance parameters of the WWTP are listed in

Table 3. Results from the WWTP simulation derived from model equations and a Monte Carlo simulation.

Parameter	Value
Inlet wastewater flow (m3/d)	49,500
Primary sludge flow (m3/d)	104.5 ± 33.0
Secondary sludge flow (m3/d)	120.2 ± 23.4
Methane production (m3/d)	2354 ± 798
Energy in biogas (MJ/d)	84,268 ± 28,590
Electricity production (kW)	370.6 ± 125.7
Heat production(kW)	507.2 ± 172.1
Digester’s thermal demand (kW) under summer conditions	263 ± 49
Digester’s thermal demand (kW) under winter conditions	402 ± 75
Dewatered digestate flow (m3/d)	33.1 ± 9.7
Energy needs for dewatered digestate transport (MJ/year)	333,246 ± 101,290
Transport costs for dewatered sludge (€/year)	45,795 ± 13,919
Thermal demand for sludge drying (kW)	925 ± 320
Specific sludge drying demand (GJ/t water evaporated)	3.1 ± 1.5
Auxiliary fuel required during winter conditions (kW)	1118 ± 348
Annual auxiliary fuel demand without considering sludge drying (GJ)	2437 ± 1537
Annual costs for auxiliary fuel demand without considering sludge drying (€)	30,472 ± 19,213
Annual auxiliary fuel demand with sludge drying (GJ)	32,823 ± 10,910
Annual costs for auxiliary fuel demand with sludge drying (€)	410,296 ± 136,377

. A significant amount of digested sludge is obtained (219 ± 44 m³/d). The sludge stream is subsequently dewatered, to reduce the amount of sludge requiring final disposal, thereby impacting the efficiency of sludge handling and transport costs. The energy associated with sludge drying is excessive. Transporting the dewatered sludge requires an annual energy demand that appears to be extremely high compared to the energy needed when dealing with the dried material (78,055 ± 13,971 MJ/year, equivalent to a transport cost of 10,726 ± 1919 €/year). However, the decision to dry sludge before transport involves a significantly higher amount of energy due to the excessive energy demanded for water evaporation. Thus, the transport of dewatered sludge translates into a mean annual expense of approximately 46,000 €. In contrast, the cost of drying the material reaches a cost that is almost ten times greater. Table 3 also indicates that during winter, auxiliary fuel is required if 50% of the heat generated by the CHP unit is recovered with the engine exhaust gases [56].

Biosolid land application is an environmentally friendly option to valorize digested material, allowing the recycling of nutrients (nitrogen and phosphorus) and the retention of carbon in soils. Biosolids are rich in phosphorus content, particularly when the plant counts with an enhanced system for phosphorus removal, either a chemical or a biological one [84]. Additionally, substituting mineral phosphates for this organic amendment helps mitigate the risks associated with the presence of Cd in low-quality rock phosphate [85,86]. Nevertheless, thermal valorization emerges as a viable alternative when sludge valorization is unfeasible due to location-specific restrictions at WWTPs regarding the presence of metals or micropollutants. The energy contained in sludge can be estimated from its higher heating value (HHV), with values ranging from 12 to 14 MJ/kg [87,88,89]. In the present case, the power derived from the biosolids hardly covers the drying needs, with a value of 1047 ± 187 kW if a mean HHV of 13 MJ/kg is assumed. This simplistic estimation shows the intrinsic difficulties found when attempting sludge thermal valorization. Thermal processes such as gasification and pyrolysis may seem to be feasible technologies for obtaining valuable fuels from sludge. However, these processes may only partly compensate for the high energy demand required for drying when integrating digestion and thermal valorization, with the solid content of the feed in the integrated system having a significant impact on the global energy balance [48,90].

3.1. Addition of H₂ Gas as a Co-Substrate in Anaerobic Digestion

The use of hydrogen as a co-substrate enhances the digester’s productivity while avoiding the inconvenience associated with handling external waste. However, the cost of the electrolyzer is a significant drawback. CO₂-biomethanation can transform the CO₂ present in biogas into methane by requiring additional hydrogen (See Figure 3). The amount of CO₂ present in the biogas was 64.7 ± 22.2 m³/h; thus, upgrading the biogas into a natural gas surrogate is achieved under the assumption of the high injection case (1.0 L of H₂/L_r d), but the extra amount of methane produced does not equilibrate the economic balance regarding the cost of the electrolyzer.

Table 4. Results from the WWTP model with H₂ injection. H₂ injection flows into the anaerobic reactor are 0.2 and 1.0 L of H₂/L_r d. Anaerobic digester works under mesophilic conditions.

Parameter	Low H2 Injection Case	High H2 Injection Case
Specific H2 injection rate (L H2/Lr d)	0.2	1.0
Hydrogen flow (m3 STP 1/d)	1176	5880
Electrolyzer size (kW)	225	1100
Electrolyzer price (€)	357,000	1,748,000
Oxygen produced from water electrolysis (m3/h)	24.5	112.5
Methane production from CO2-biomethanation (m3/h)	12.3	61.3
Electrolyzer heat recovery (kW)	33.7	168.6
CHP heat recovery (high temperature gases) (kW)	1170	5585
Digester’s thermal demand, winter period (kW)	401 ± 72
Auxiliary fuel required to cover digester’s thermal demand under winter conditions (kW)	150 ± 90

¹ STP: Standard temperature and pressure conditions.

shows the main results derived from the model when assuming mesophilic conditions for the digester. If the digester’s conversion capacity for transforming CO₂ into methane is considered, a yearly methane production equivalent to 5336 MWh/year would be available, translating into 240,106 € of annual savings. This extra methane is enough to cover any demand for auxiliary fuel if digestate drying is not contemplated into the scenario. However, an average extra cost of 170,200 € is still required when a thermal drying operation is included.

The transformation of CO₂ into methane offers the advantage of utilizing the existing equipment to upgrade the biogas, achieving a quality comparable to that of natural gas. However, in addition to the electrolyzer’s investment and installation costs, hydrogen production involves high electricity consumption. Currently, the lowest electricity price in Europe was reported by Finland, with a value of 0.0767 €/kWh, and the highest was that for Cyprus, with a price of 0.2578 €/kWh for the second half of the year 2024 [91]. Applying these prices to the economic balance translates into an hourly cost of producing methane between 81 and 272 € for obtaining the H₂ flow required to upgrade the biogas. In contrast, the market price of methane produced for the same interval only reaches a value of 27.5 €. This discrepancy evidences the difficulties found by technologies dealing with CO₂ utilization. The price of electricity to equilibrate the balance needs to drop below 0.025 €/kWh if methane market prices are kept constant.

The electrolyzer produces oxygen in addition to hydrogen. In the present case (see Figure 3), the oxygen produced may serve as an extra supply for the air in the waste-activated sludge process. Based on the assumptions about the WAS unit, the airflow estimated was 311 m³ air/min (equivalent to 65.0 m³ O₂/min). However, the oxygen flow from the electrolyzer is only 2.0 m³ O₂/min, which hardly contributes to reducing the air supply by 3%. This slight decrease is also extrapolated into a small decrease in power demand.

The heat recovered from the electrolyzer helps meet the thermal demand of the digester during winter. Consequently, if sludge drying is not included, incorporating the electrolyzer reduces the need for additional fuel, resulting in significant savings. In addition, the extra methane derived from the CO₂-methanation process is now available as fuel in the CHP engine. Thus, the amount of electricity is 231 kW, which translates into annual savings of 154,700 € when considering an electricity price of €0.0763/kWh (the average value reported in Spain during the second half of 2024). The previous assumptions result in a negative economic balance, even if the cost of electricity for producing hydrogen is set to zero.

Figure 4a shows the effect of reducing the installation costs by 10% and 20%, along with the proportional decrease in operating and maintenance activities. Even with this specific reduction in the equipment cost, the electricity demand was still considered zero-priced. Any price assumed for the energy demand when producing hydrogen results in a negative economic balance, given the current price of methane. Achieving a positive result is possible if the price of methane doubles (Figure 4b) or a 10% decrease in electrolyzer investment costs is assumed. However, even in the best-case scenario (a methane price of 90 €/MWh and a 20% reduction in the electrolyzer investment costs), the payback period exceeds 10 years.

The price of the energy required for producing hydrogen may be set to zero if this energy is intended for storage or for avoiding disruptions in the electricity market whenever the production of renewable energies is excessive. Even though several researchers propose the biological reaction of transforming H₂ into CH₄ as a feasible option [92,93,94], the current investment price of electrolyzers and electricity prices make this approach unfeasible. Gantenbein et al. [36] reported that an electricity price of 5 c€/kWh and an electrolyzer investment cost below 1000 €/kW were needed to attain economic feasibility. Current market prices in Europe are above 7 c€/kWh and values close to 0.25 c€/kWh are offered for non-household sectors [91]. While lowering electricity prices can help decarbonize the economy by encouraging electrification in residential and industrial sectors, excessively low prices may deter investors and hinder the adoption of efficiency measures to reduce electricity consumption.

3.2. CO₂-Biomethanation as a Technology for Transforming Captured CO₂

The previous assumptions considered the basic approach of direct injection into the anaerobic digester, setting a 1 L of H₂/L_r d as the injection rate when considering the high injection case. Using gas recirculation and other configurations can reduce mass transfer limitations, thereby allowing a higher injection value [17,73]. Pan et al. [95] reported a 72% increase in the mass transfer improvements associated with the change in reactor configuration by introducing a draft tube to allow flow circulation. Another factor to consider is the higher bottom pressure of large-scale reactors, which favors mass transfer by increasing hydrogen solubility thanks to the higher concentrations attained at the gas–liquid interface. However, this benefit may not necessarily translate into a higher mass transfer as explained by Jensen et al. [96], who reported that the effect of a reduced bubble size due to pressure (if the superficial bubble area is not increased) may offset the previous advantage. Operating under higher temperatures reduces gas solubility but also increases the reaction rates of the biological systems, thus explaining the higher injection rates used by different authors when working under thermophilic conditions [71,74]. However, when considering co-digestion with other wastes, this is not always the case, with mesophilic systems reporting higher biogas production values under certain conditions than those under the thermophilic regime [97], because the increase in the conversion rate does not necessarily lead to a higher biogas production.

The high injection rates applied under thermophilic digestion come with the identical drawback of high investment costs and excessive electricity demands.

Table 5. Results derived from the simulation when digester runs under thermophilic conditions. H₂ injection flows are 5 and 30 L of H₂/L_r d.

Parameter	Low H2 Injection Case	High H2 Injection Case
Specific H2 injection rate (L H2/Lr d)	5	30
Hydrogen flow (m3 STP/d)	29,400	176,400
Electrolyzer size (MW)	5.3	31.6
Electrolyzer Price (Millions €)	5	13.9
Oxygen produced from water electrolysis (m3/h)	612	3675
Methane production from CO2-biomethanation (m3/h)	306	1840
Electrolyzer heat recovery (kW)	840	5056
Digester’s thermal demand, summer period (kW)	475 ± 88
Digester’s thermal demand, winter period (kW)	622 ± 115
Auxiliary fuel required during winter period to cover digester’s thermal demand (kW)	0

presents the main parameters of the scenario analyzed under both low and high injection cases, considering thermophilic digestion. Oxygen derived from the water electrolysis process can be used in the WAS treatment, thus reducing energy demand by 12% and 75% under the low and high injection cases, respectively. This benefit becomes insignificant when taking into account the power demand of the electrolyzer (see Table 5). The higher thermal demand of the thermophilic digester can be supplied by the extra heat obtained from the electrolyzer (as heat recovery) even during the winter period. However, once again, this benefit is seamless based on the high electricity consumption of this equipment.

An equivalent amount of CO₂ can be converted into methane, making this technology appealing because it enables the digester to function as a CO₂ utilization unit rather than merely a biogas upgrading system. The annual amount of CO₂ that can be transformed is 4160 t CO₂ for the low injection case and 30,500 t CO₂ for the high injection case. Figure 5a shows the scheme representing the introduction of a CO₂ stream along with the electricity generated by the CHP engine when methane is valorized to produce heat and electricity.

The energy recovery (as electricity) of the present strategy is 22%, and the water demand of the electrolyzer accounts for approximately 100 m³/d and 640 m³/d for the low and high injection cases studied, respectively. One additional disadvantage of the hydrogen-based economy is the high water demand required for producing this valuable gas. If the water use of the isolated electrolysis step is considered, then the impact of the amount of water consumed for hydrogen production, compared with the amount of water available, may be cataloged as negligible following the criteria of Beswick et al. [98]. However, this is not the case here, where a significant amount of water is required daily to allow CO₂ conversion. Despite this drawback, H₂ produced from water electrolysis powered by either renewable energy or nuclear energy has the lowest impact when compared with steam reforming or when electricity is derived from an electrical grid with a high share of carbon-producing emission technologies [99].

Given the economy of scale, the balance is significantly improved (see Figure 6); however, this outcome is only achieved if no price is set for the energy demanded by the electrolyzer. The best case considered here would translate into a zero value for the NPV parameter when a price as low as 11 €/MWh is introduced into the balance sheet. The introduction of revenues linked to CO₂ conversion (50 €/t CO₂) does not cause a significant improvement to afford the electricity market price for hydrogen production (zero value of NPV is obtained for an electricity price of 17 €/MWh). Increasing the injection rate (high injection case) to 30 L of H₂/L_r d results in a better economic balance due to the advantage from the economy of scale associated with the electrolyzer capital investment, but the fact that the price of electricity is higher than that of methane makes the whole idea of storing energy in this form unfeasible unless the price of electricity for producing hydrogen is set to zero.

The high injection case can result in a payback period of 8 years, provided the price of electricity increases to 37.8 c€/kWh, as long as the electrical cost of generating H₂ is not factored into the equation. Figure 6b also shows the electricity price that allows a payback period of 8 years if the investment cost of the electrolyzer is reduced by 10% and 20%, respectively. This same exercise was conducted under the best-case scenario, assuming a 20% decrease in electrolyzer investment costs. The acceptable price of electricity to produce hydrogen is shown in Figure 7.

The estimated values were determined by analyzing payback periods ranging from 8 to 15 years and considering the cases where sludge was dewatered and when sludge was dried. In this assessment, it was initially assumed that no revenues were earned from transforming CO₂, whereas the other scenarios considered potential earnings of €50 and €100 per ton of CO₂ for the same activity. As can be seen from this figure, neither case can achieve profits at the current electricity price under the Spanish scenario (76.3 €/MWh). Figure 7b shows better results (sludge drying case) due to the savings associated with the dryer fuel demand. The greater amount of methane available for producing electricity in the plant also increases the volume of hot combustion gases, thus supplying the heat required for drying sludge.

3.3. Challenges and Opportunities for Implementing CO₂-Biomethanation into WWTPs

The strategy of transforming conventional anaerobic digesters into either biogas upgrading units or CO₂-conversion units to become part of an energy storage system presents numerous limitations, including the high cost of the electrolyzer and electricity prices. Additionally, other technical challenges are worth noting. Hydrogen is a small molecule with high diffusion and energetic content, resulting in a high flame propagation speed [100], which makes the handling of this fuel extremely challenging. Metal embrittlement is a significant concern when dealing with hydrogen in daily plant activities. The small size of the molecules facilitates diffusion into metal structures, causing cracks and fractures [101,102]. Safety considerations are usually associated with materials used for transporting and storing hydrogen and issues related to the daily operations of the personnel involved in handling this fuel since any leakage goes undetected by humans because it is a colorless, odorless, and tasteless flammable gas [103]; thus, specific protocols need to be implemented in conjunction with the installation of H₂-sensors for detecting gas leakage and the specification of shut down protocols and isolation of areas when leakages are detected [104]. The operators of WWTPs have experience working with methane, which is a flammable gas. Nevertheless, the properties of this gas differ significantly from those of hydrogen, as methane has a higher ignition temperature, a lower flame speed, and a narrower range of flammability limits compared with hydrogen. Therefore, special training is required for operators working in zones with a higher flammability risk, which translates into higher labor and maintenance costs, thereby reducing the feasibility of the present approach even further.

Despite these challenges, the conversion of extra-renewable electricity into a chemical compound using existing biological reactors seems feasible due to the advantages of working with mixed microflora, the lack of investment required for constructing large-scale reactors, and the availability of trained personnel. Thus, the strategy leverages the experience gained from several years of reactor design and the operation of industrial reactors. Although other reactor configurations, such as gas lift, anaerobic filters, and membrane reactors, are also available, the upfront cost of building a new large-scale reactor and acquiring the electrolyzer, along with all the necessary auxiliary equipment, would annul any economic benefit.

4. Conclusions

The biological conversion of CO₂ into methane offers the possibility of either upgrading biogas or transforming a conventional anaerobic digester into a CO₂ utilization unit when additional captured CO₂ is introduced into the system. Although the proposal may appear environmentally friendly, it entails excessive energy demands for operating the electrolyzer and involves high investments, negating any potential profitability. The present study assessed the conversion of CO₂ under mesophilic and thermophilic conditions by assuming direct hydrogen injection into a digester operating in a conventional WWTP. Introducing a water electrolysis unit enables heat recovery, which can be used to cover the digester’s thermal demand. Oxygen derived from the electrolyzer can be used as a supplement to the air stream required for the WAS treatment system. However, the benefits obtained in this case are modest, resulting in a 3% reduction in airflow and covering the digester’s thermal demand during the winter period, if assuming a hydrogen injection rate of 1 L of H₂/L_r d (first scenario). The increase in the hydrogen injection rate (5 and 30 L of H₂/L_r d) was evaluated in the second operational mode, where the digester runs under thermophilic conditions. The economy of scale in hydrogen production favored this approach but required establishing lower electricity prices for this specific operation. The system only attains profitability if a price between 14 and 30 €/MWh is set and additional revenues are obtained from CO₂-biomethanation. Setting the standard price of electricity for hydrogen production resulted in a negative NPV for any of the cases analyzed.