Experimental Study on Direct and Indirect Carbonation of Fly Ash from Fluidized Bed Combustion of Lignite

Abstract

The research problem was to determine the possibility of aqueous mineral carbonation using fly ash from lignite fluidized bed combustion. Both direct and indirect routes were used. The innovative nature of the research consisted of conducting experiments at atmospheric pressure and ambient temperature (20 °C). The synthetic gas mixture with composition analogical to the flue gas (nitrogen and up to 16 vol.% of carbon dioxide) was used. The experiments proved that almost all CO₂ from the gas was chemically bound at pH > 12. The sequestration capacity of studied fly ash is about 55–76 g CO₂ per 1 kg of ash in the case of the indirect method, and 80–95 g CO₂ per 1 kg of ash for the direct route. These values are similar to those presented in the literature, but unlike most publications, they were obtained under ambient conditions, which can significantly reduce the costs of the process.

1. Introduction

The problem of global warming is currently one of the fundamental challenges facing civilization. The phenomenon of global warming is widely associated with anthropogenic carbon dioxide emissions. A number of international agreements, including the famous Kyoto Agreement (2005) and the Paris Agreement (2015), aim to reduce CO₂ emissions. To coordinate these efforts, the Intergovernmental Panel on Climate Change (IPCC) was established, which regularly publishes reports on the current state of the climate, knowledge about climate issues and CO₂ emissions, and efforts to reduce these emissions [1]. It is believed that growing CO₂ emissions threaten to lead to undesired temperature increases in the future.

IPCC reports analyze the sources and volumes of CO₂ emissions and propose a number of actions to reduce the emissions. Methods for reducing CO₂ emissions into the environment include a set of measures known as CCS (Carbon Capture and Storage), which involves capturing CO₂ from exhaust gases (e.g., from power plants) and storing them [2]. CO₂ capture includes a range of technologies, the most common of which is the CO₂ absorption in alkaline solutions (amines, and less frequently, ammonia). Storage can be implemented in selected geological layers (e.g., depleted oil or natural gas deposits) or in sediment layers on the ocean floor. Yet another method is the chemical binding of CO₂ in alkaline minerals, called CO₂ mineralization or mineral carbonation [3]. In general, carbonation involves the reaction of CO₂ with alkaline metal oxides to produce carbonates according to the following general notation [2]:

MeO + CO₂ = MeCO₃

where Me indicates divalent metal, e.g., calcium, magnesium, or iron. The issue of mineral carbonation, particularly of waste fly ash from coal combustion, is the subject of this research.

The concept of mineral carbonation is discussed by Lackner et al. [4]. The mineralization reaction is exothermic, so costs will primarily be related to raw material preparation (mining, crushing, and transporting) and ensuring better kinetics. Two types of processes are discussed: direct carbonation in a gas–solid system at high temperature and pressure, and carbonation in an aqueous solution, where Mg and/or Ca are extracted into the solution, usually also at elevated temperature and CO₂ pressure. In addition to the methods mentioned above, in situ technology is discussed, where gas is injected into geological deposits containing alkaline minerals capable of binding CO₂ [5]. The aqueous carbonation can also be carried out by a direct (one-step process) or indirect (two-step process) route [6,7]. Carbonation technology at the industrial scale is limited by the high energy input required to prepare the mineral raw materials and to conduct the process at high CO₂ pressures to accelerate the kinetics.

Natural sources of alkaline oxides include minerals containing mainly calcium and magnesium silicates, such as wollastonite, serpentine, olivine and talc. However, the appropriate minerals must be mined, crushed, and transported to the intended location. The significant costs of these operations must be taken into account. Moreover, deposits of many minerals are depleting. This prompts a search for waste materials capable of chemically binding CO₂. Fly ash is produced in coal-fired power plants, i.e., in locations where significant amounts of CO₂ are released, eliminating the problem of transport [6]. The ash mineralization process is usually carried out in an aqueous environment, under high pressure and at elevated temperature, which improves the reaction kinetics, which is usually too weak [7]. In addition to fly ash, other mineral wastes are also considered, e.g., steel slag and cement kiln dust [8].

The CO₂ sequestration capacity of alkaline fly ash is 230 g CO₂ per 1 kg of ash, according to [9]. This fly ash contains large amounts of CaO and MgO. According to [8], the capacity of steel slag and cement kiln dust amounted to 105–135 g CO₂ per 1 kg of the dust. Wehrung et al. [10] tested various wastes and obtained different capacities: 144.5 g CO₂/1 kg for bag filter residues; 101 g/kg for bottom ash; 93 g/kg for boiler electrofilter fly ash and 167 g/kg for other wastes. The amount of material needed to capture 1 ton of carbon dioxide is smaller in the case of common minerals, available in large quantities in nature, than in the case of fly ashes, as can be shown in Figure 1. This demand was determined based on the stoichiometric content of alkali oxides, using modified Steinour formula for theoretical CO₂ sequestration capacity presented in [11]. The figure also shows the theoretical requirement of the two industrial wastes: fly ash from pulverized coal fired boilers with desulfurization products (PFD) and fly ash from lignite fluidized bed combustion (FBC). Both ashes come from Polish power plants. The results are close or slightly lower compared to the above-mentioned literature. For PFD, it is about 8.8 t/t CO₂ (which means 113 g CO₂/1 kg fly ash), and for FBC, 5 t/t CO₂ (201 g CO₂/1 kg of fly ash). It should be noted that the differences may result from the differences in the composition of the minerals and waste dust considered. Fly ash, in contrast to minerals, contains free alkaline compounds. Taking into account only the water-soluble alkaline oxides and hydroxides, determined by the XRD analysis, the theoretical sequestration capacity is 26.14 g CO₂/1 kg fly ash for PFD and 95.4 g CO₂/1 kg fly ash for FBC.

However, another problem is the ratio of the amount of CO₂ produced to the absorption capacity of the generated fly ash. Back et al. [9] estimated the possibility of absorption of only 2% of CO₂ for a lignite-fired power plant. A similar percentage of CO₂ uptake was obtained in [12] using coal fly ash. Uibu et al. [13] found that for oil-shale combustion, sorption of 10–11% of the produced CO₂ was possible. This leads to the conclusion that only a small part of the CO₂ produced in coal-fired power plants can be neutralized by the ash produced there. On the one hand, ash is a locally available waste, making its use for CO₂ capture easy. Thus, the vast majority of CO₂ captured from exhaust gases must be managed differently. In this situation, economic considerations must determine the strategy.

At this point, it is worth briefly discussing the differences in ash composition. In the case described in [13], i.e., the oil-shale combustion, the rather unusual ash contained up to 30% of free Ca–Mg oxides, which is higher than in the case of coal combustion. Moreover, the amount of ash produced in relation to the amount of CO₂ emitted is also higher than, for example, in the case of lignite combustion. Even the ash from lignite combustion in [9] had a higher alkali content than the ashes used in the studies presented here. The alkali content is much higher in the ash from lignite combustion (FBC) than that from hard coal (PFD); this difference practically excludes PFD from use as an effective CO₂ sorbent. However, fly ash from pulverized coal-fired power plants (PFD) is a valuable additive to cements. It contains significant amounts of finely divided silica, which undergoes a pozzolanic reaction during cement setting, improving the strength of the cement material and reducing water penetration [14]. Due to the decreasing consumption of hard coal in power plants, the availability of PFD ash is decreasing, and it is practically entirely used in the cement industry. On the other hand, lignite ash has less use in the production of construction materials. It contains significant amounts of gypsum, which poses problems in its use as a cement additive [15,16,17]. This type of ash is a waste with limited application, therefore its use for CO₂ storage seems rational. An additional benefit of using coal fly ash in the carbonation process is the possibility of reducing the potential leakage of heavy and toxic metals contained in waste, as stated in [18,19].

The process of CO₂ mineralization in aqueous solutions using natural minerals presents two unfavorable characteristics. The first is the relatively poor kinetics of both the dissolution of alkaline components and the reaction of Mg and Ca ions with weak carbonic acid. To improve the kinetics of these phenomena, the obvious remedy is to increase the process temperature. However, the dissolution of carbon dioxide in aqueous solutions, generally poor, decreases drastically with the increase in temperature. Hence, it is necessary to use high pressures (usually of the order of several MPa) to absorb CO₂ and ensure its relatively high purity (without inert and acidic gases such as N₂, SO₂, and NO_x) [7,8,20]. Operating the system at high pressures and temperatures significantly increases costs.

Various chemicals can be used to accelerate the limiting step—kinetics of mineral dissolution [21,22]. Alkali oxides, primarily calcium and magnesium, bind CO₂ to form carbonates. Advantages of the carbonation process include safe and permanent immobilization of CO₂ and the possibility to further utilize the product thus obtained (e.g., as a building material). In the case of the use of minerals, the limiting step in the rate of sequestration is the dissolution process. The dissolution rate of some minerals, determined from the literature data [22,23,24,25], are shown in Figure 2. The figure also shows the dissolution rate of both ashes, calculated based on solubility data presented in [26,27]. Industrial wastes have the advantage of much higher kinetics of the sequestration process. The much higher dissolution rate of waste compared to the dissolution rate of minerals makes waste a very attractive material for CO₂ binding.

For practical application, the economics of the process, i.e., costs of absorbing CO₂ in individual minerals and alkaline waste, are very important. These costs are usually very high, especially for natural minerals, which must be mined, crushed, and transported to the point of use. In the work of Lackner et al. [4], the high costs of CO₂ mineralization in natural, fossil alkaline minerals were already mentioned. Metz et al. [2] estimated these costs at USD 50–100 per ton of CO₂ (prices here and in subsequent citations always refer to the year of publication). Huijgen et al. [28] estimated the cost of binding 1 ton of CO₂ in cement at USD 22 and in concrete at USD 37. Huijgen [29] estimated the costs of CO₂ absorption in wollastonite at USD 102 per ton and in steel slag at EUR 77. Sanna et al. [5] also mention high costs, emphasizing the potential of this technology for CO₂ binding on an industrial scale. It can be argued that other methods of storing CO₂ in liquefied form—in geological layers, in disused mines, or in depleted natural gas and oil deposits beneath the ocean floor—are cheaper than mineralization, despite the significant challenges associated with implementing these operations. However, CO₂ mineralization, regardless of the raw materials used, is a safer method: CO₂ cannot escape from stable chemical compounds, which theoretically could happen in specific situations during very long-term storage of liquid carbon dioxide in deep geological layers.

Despite the prohibitive costs of the operation so far, mineralization is slowly beginning to be used on a larger scale than just the laboratory. Wehrung et al. [10] state that “In recent decades, mineral carbonation has evolved from a promising waste management strategy into a commercially viable solution”. Carbon8 Systems claims [30] that a demonstration of a mobile container for CO₂ capture and the production of construction semi-finished products is now available. The first pilot plant was built in The Netherlands in 2020. This project has the potential to capture 22 Mt of CO₂ in Europe. Hills et al. [23] present a mobile “CO2ntainer” installation that uses cement plant waste to store CO₂ released on-site. The system, implemented at two cement plants (Canada and the UK), uses a total of several thousand tons of waste annually to produce construction aggregate. The scope of CO₂ mineralization will probably increase significantly in the coming years, with the use of industrial waste being the first priority.

This paper presents the results of CO₂ mineralization in fly ash from lignite combustion in fluidized bed furnaces. Both the direct and indirect methods were studied. Unlike most publications, atmospheric pressure, near-ambient temperature (20 °C), and a working gas composition similar to that of power plant exhaust gases (a mixture of nitrogen and carbon dioxide, containing up to 16% CO₂) were used. All CO₂ contained in the gas introduced in the form of bubbles was chemically bound at pH > 12. Pure CO₂, high pressure, and elevated temperature of the solution were not used. Despite this, satisfactory process capacity and kinetics were achieved, as presented in the following chapters.

2. Materials and Methods

2.1. Characteristics of Industrial Waste

The fly ash from fluidized bed combustion (FBC) of lignite, collected from the Turów power plant in Poland, was used. The main properties of the ash are presented in

Table 1. Main characteristics of fluidized ash studied.

Chemical Composition, wt.%
SiO2	CaO	MgO	Al2O3	Fe2O3	Na2O	K2O	SO3	P2O5	TiO2
27.0	29.1	2.03	20.2	4.54	1.27	1.01	8.75	0.2	1.68
Phase Composition, wt.%
SiO2	CaSO4	CaO	Ca(OH)2	CaCO3	α-Fe2O3	NaAlSi3O8	Rozenite	Jarosite	Amorph.
1.9	12.4	12.0	0.2	6.4	1.6	1.5	0.6	0.4	63.0
Particle Size Distribution, µm					Surface and Porosity
D0.1	D0.5	D0.9	D[3,2]	D[4,3]	BET, m2 g−1	Micropore Area, m2 g−1	Total Pore Volume, mm3 g−1	Micropore Volume, mm3 g−1	Average Pore Size, nm
3.96	24.64	87.36	8.94	36.79	6.664	0.588	37.1	0.260	14.24

. The chemical composition and phase composition of the solid ash were determined using Thermo iCAP 6500 Duo ICP Spectrometer (Thermo Fisher Scientific, Waltham, MA, USA) and Empyrean X-ray Diffractometer (PANalytical, Almelo, The Netherlands), respectively. BET surface area as well as total and micropore volumes were examined using ASAP 2020 physisorption analyzer (Micromeritics, Norcross, GA, USA). The particle size distribution was measured with a Malvern Mastersizer 2000 (Malvern Instruments Limited, Malvern, UK laser diffraction particle size analyzer using the wet dispersion method in propan-2-ol.

A chemical analysis showed that FBC ash contained 29.1 wt.% of calcium (expressed in terms of CaO). Approximately 41.8 wt.% of calcium in FBC ash occurs as free oxides, whereas magnesium is present solely as a component of the vitreous phase. Solubility studies confirmed [26] that the amount of magnesium released into the solution is too low to have any measurable impact on the process of carbonation. Therefore, it is calcium oxide that is the main reactant in the carbonation process using the ash, and the carbonation process can be described by the following overall reactions:

CaO + H₂O → Ca(OH)₂

Ca(OH)₂ + CO₂ → CaCO₃ + H₂O

A laser diffraction analysis found that the median particle size was about 25 µm, and the size of the largest particles did not exceed 200 µm. Thus, the dust was finely comminuted. This average particle size and surface area are typical for fly ashes [6]. Increased finesses and consequently high surface area enhance the dissolution of waste and the reaction with CO₂. The structure of ash particles shown in Figure 3 is typical for ash from fluidized bed boilers, when the combustion temperature of about 750–950 °C is significantly lower than that in conventional pulverized coal fire boilers. The particles have irregular shape and a porous and uneven surface, in opposition to the glassy spherical shape of particles from conventional boilers. Such surface structure should also enhance dust dissolution and the sequestration process.

2.2. Sample Pre-Treatment

The FBC was used in its original form in the direct carbonation studies. Prior to the indirect carbonation studies, the solid fly ash was dissolved in deionized water at room temperature for a time sufficient to obtain a solution of the required concentration. The ash dissolution process was carried out in a 5-Litre reactor made of borosilicate glass (QVF PILOT-TEC GmbH, Jena, Germany) equipped with a propeller mixer with a speed controller, as can be seen in Figure 4 (#4—dissolution tank). The resulting suspension was filtered, then clear solution was directed to the second reactor (#7 in Figure 4). The concentrations of ions were determined in the filtrate by complexometric titration using EDTA (Ca⁺²) and by a spectrophotometer DR2800 (Hach Company, Loveland, CO, USA) using the Hach Lange cuvette test (SO₄⁻², LCK353).

The indirect carbonation process was studied for three different calcium ion concentration values. The maximum concentration of Ca⁺² in the input solution of 0.0536 mol·dm⁻³, which corresponds to a solution saturated by calcium ion, was obtained by dissolving the studied fly ash in water, for a waste-to-solvent weight ratio of 1:20. The dissolution process was carried out for 30 min. The concentration of SO₄⁻² was about 1200 mg·dm⁻³ (equal to 0.0125 mol·dm⁻³), which means that about 38% of free CaO and 27% of CaSO₄ contained in the FBC dissolved in the water.

The so received solution was diluted in order to obtain the lower values of the Ca⁺² concentration. At the same time, an appropriate amount of sodium sulfate was added to maintain a constant concentration of SO₄⁻² in the inlet solution (approximately 0.0125 mol·dm⁻³). In the second step of indirect carbonation, the sulfate ion affects the precipitation of calcite [28]. It was found that the lower SO₄⁻² concentration is advantageous for the precipitation of CaCO₃. On the other hand, it was found in [31] that the presence of Na₂SO₄ enhances the rate of CO₂ absorption in an aqueous Ca(OH)₂ solution. However, the Na₂SO₄ concentrations used in [31], ranging from 0.5 to 7.5 wt.% (corresponding to 0.035–0.528 mol·dm⁻³), were much higher than the sulfate ion concentration fixed in the present work, equal to 0.0125 mol·dm⁻³, with only a fraction, up to 60%, coming from the sodium sulfate added to the solution. It should be noted here that after dissolving the ash containing CaSO₄, sulfate ions will always be present in the solution, but its concentration will be low due to the poor solubility of CaSO₄ in water. Furthermore, the concentration of SO₄⁻² can be controlled to a certain extent during the ash dissolution stage (by changing the ash-to-water ratio and/or the dissolution time). The influence of sulfate ions on the carbonation process was not investigated in the present work. In order to eliminate the influence of this parameter, the studies of indirect carbonation process were conducted for the same concentration of sulfate ions. The concentration of 0.0125 mol·dm⁻³ is the minimum SO₄⁻² concentration for which the maximum calcium ion concentration was obtained in the dissolution studies [27].

2.3. The Indirect Carbonation Process Studies

The studies on the indirect carbonation process were performed in a laboratory installation presented in Figure 4. A photo of this installation is shown in Figure S1 in the Supplementary Materials. The reactor in which the precipitation occurs (#7 in Figure 4), with a capacity of about 1.5 dm³, was made of borosilicate glass (Normag, Ilmenau, Germany) and equipped with a heating jacket, a draft tube with 4 baffles, and a propeller mixer with a speed controller. The reactor was also equipped with a system for dosing the gas mixture (N₂/CO₂) and dispersing the gases in the liquid, which was a Ca⁺²-rich clear solution from the dissolution tank (#4 in Figure 4). The composition of the gas mixture was similar to that of the flue gas (9–16 vol.% CO₂). A polypropylene tube with an internal diameter of 4 mm was inserted into the steel shaft holding the agitator, and the end of the tube was brought out just below the agitator blades. A dispersing device in the form of a low-porosity spherical stone, normally used for aerating an aquarium, was installed at the end of the tube. Dosing the gases using a glass sinter did not work, as the sinter clogged up due to the very intense precipitation of carbonate particles, making it impossible to continue the process. The phenomenon of pore clogging did not occur when an aeration stone was used.

Experimental Procedure

The indirect carbonation studies were performed at ambient pressure and a temperature of 20 °C, in a reactor operating in a semi-periodic manner (continuous gas dosing). The empty reactor was washed with a stream of nitrogen. Then, the reactor was filled with a previously prepared solution after dissolving the FBC, further dosing a stream of pure nitrogen to the reactor at a specified flow rate. Once the set temperature was reached, a carbon dioxide stream with a given flow rate was turned on. During the measurement, the pH (Oakton pH Meter, accuracy ± 0.01) and temperature of the solution, the flow rate of the gases (N₂ and CO₂ separately) and the CO₂ concentration in the outlet gas stream were continuously recorded. Periodically, samples of the suspension (approx. 20–25 mL) were taken from the reactor and then filtered. The following concentrations were determined in the obtained filtrate: Ca⁺² ion by complexometric titration using EDTA solution (average relative error ± 2%), SO₄⁻² ion (LCK353 test, standard deviation 10.4 mg·dm⁻³), and total amount of dissolved CO₂ (LCK388 test, standard deviation 20 mg·dm⁻³) using the KORONA DR 2800 spectrophotometer. The total amount of dissolved CO₂ is understood as the sum of carbonates and carbon dioxide that are dissolved in the solution, calculated as CO₂. The gas flow rate was measured using an Aalborg Mass Flow Meter GFM17 (Aalborg Instruments, Aalborg, Denmark) with an accuracy of ±1.5% of the full range (1 dm³·min⁻¹). The Servomex Gas Analyser Series 1400 (Servomex Company, Sugar Land, TX, USA) was used to measure the CO₂ concentration in the gas leaving the reactor, with a measurement accuracy of 1% of the full scale (25% CO₂). After the measurement, the reactor was emptied. One part of the suspension was filtered, and the obtained solution was analyzed. The remainder of the suspension was used to analyze a particle size distribution of the resulting product.

The process parameters used in the indirect carbonation studies are presented in

Table 2. Process parameters used in carbonation studies.

	Indirect Route	Direct Route
Stirrer speed, min−1	300, 600, 1100	300, 600, 1100
Concentration of CO2, vol.%	9.65, 12.76, 15.75	9.64, 12.69, 15.83
Ca+2 initial concentration, mol·dm−3	0.0212, 0.0363, 0.0536	n.a.
Ash-to-water weight ratio	n.a.	1:100, 1:50, 1:20

. The total flow rate of the gas supplied to the reactor was constant at approximately 680 cm³·min⁻¹. For the given process conditions, the experimental tests were carried out twice: 1—until the equilibrium state was reached (establishing pH readings and establishing CO₂ concentration in the gas at the reactor outlet) and 2—until the solution pH dropped rapidly (i.e., to a pH value of approx. 10).

2.4. The Direct Carbonation Process Studies

The studies on the direct carbonation process were performed in a laboratory installation, presented in Figure 5. The main component of the plant, the reactor in which fly ash dissolution and calcite precipitation take place (#5 in Figure 5), together with the gas dosing system, was the same as in the indirect carbonation studies. The direct carbonation studies were performed at ambient pressure and a temperature of 20 °C for gas mixtures with a composition similar to that of the flue gas. The experimental procedure used for both carbonation routes was also very similar. The difference is that in the case of direct carbonation, the reactor was filled with deionized water and, once the set temperature was reached, the CO₂ flow was simultaneously turned on and the fly ash was fed into the reactor.

The process parameters used in the direct carbonation studies are presented in Table 2. The total flow rate of the gas supplied to the reactor was constant and was approximately 680 cm³·min⁻¹, similar to the indirect carbonation studies. For the given process conditions, the experimental tests were carried out three times: 1—until the equilibrium state was reached (constant pH readings and CO₂ concentration in the gas at the reactor outlet), 2—until the solution pH dropped rapidly (i.e., to a pH value of approx. 10), and 3—to a pH of approximately 7.

3. Results and Discussion

3.1. Indirect Carbonation Process

Changes in solution pH during the indirect carbonation process for different values of the stirrer speed are presented in Figure 6a. Figure 6b–d show changes in the concentration of Ca⁺² and SO₄⁻² ions as well as the total amount of CO₂ dissolved in the solution occurring at that time. The lines are drawn only to guide the eye. The effect of the stirrer speed on the precipitation process was studied for a concentration of CO₂ in the stream directed to the reactor of 12.8 vol.%, with an initial concentration of calcium ion in the solution equal to 0.036 mol·dm⁻³. In all cases studied, approximately 77% of the calcium ions in solution originated from free calcium oxide/hydroxide contained in the fly ash. The rest came from CaSO₄. The products of the calcium sulfate dissolution reaction do not react with components of CO₂/H₂O system, so the calcium cations formed as a result of this reaction have no sequestration potential [11].

As can be seen in Figure 6a–d, up to the solution pH of approximately 12 (i.e., after of about 7.5–8 min), the concentration of calcium ions systematically decreases, and the total amount of CO₂ dissolved in the solution practically does not change. At the same time, the decrease in the concentration of calcium ions corresponds to almost the total amount of CO₂ introduced to the reactor. This means that 100% of the carbon dioxide directed to the reactor at that time was bound in the form of calcium carbonate. The calcium ion concentration at this time dropped to about 0.016 mol·dm⁻³, so about 55% of the calcium dissolved in the solution reacted with CO₂.

This step is followed by a sharp drop in the pH of the solution and an increase in the concentration of total CO₂ dissolved in the solution, as only part of the carbon dioxide introduced to the reactor has reacted with calcium. At the same time, the Ca⁺² ion concentration still decreases for a short time (calcium carbonate is still formed) and then starts to increase. The latter results from the fact that with increasing acidity of the solution, there is a dissolution of calcium carbonate due to an increase in the concentration of HCO₃⁻ ions [32,33,34]. The minimum determined concentrations of calcium ions in the solution were as follows: 0.0139 mol·dm⁻³, 0.0141 mol·dm⁻³, and 0.0142 mol·dm⁻³ for the stirrer speeds of 300, 600, and 1100 min⁻¹, respectively, at a solution pH of approximately 9–9.5. These values of concentration correspond with over a 60% conversion of calcium, and the corresponding degree of CO₂ bonding was between 83 and 95%. At the same time, the concentration of sulfate ions in the solution both during and after the precipitation process is practically the same as in the input solution. It can therefore be seen that only calcium carbonate precipitates from the solution. The shape of the curves showing changes in solution concentration and pH over time were similar for all the tests performed.

Analysis of the data presented in Figure 6a–d shows that the stirrer speed has almost no effect on the carbonation process. It can therefore be concluded that the rate of the carbonation process, under the test conditions, depends mainly on the reaction kinetics and not on the mass transport rate. Only for 300 rpm does the calcium ion concentration curve differ slightly from the others. It is likely that this stirrer speed value is too low to achieve perfect mixing of the solution. So, the next tests were performed for a speed of 600 rpm.

For the different values of initial calcium ion concentrations, as can be seen in Figure 7a, the times needed to achieve a solution pH of about 12 were as follows: 4.5, 7.5, and 9 min for the initial concentrations of calcium ions in the solution equal to 0.0212, 0.0363, and 0.0536 mol·dm⁻³, respectively. During this time, the concentration of calcium ions dropped to a value corresponding to approximately 50% calcium conversion. However, the degree of CO₂ bonding was lowest for the lowest initial calcium ion concentration in the solution, at around 80%. For the highest initial calcium ion concentration in the solution, i.e., 0.0536 mol·dm⁻³, the depletion in the amount of calcium exceeded the amount of CO₂ introduced into the reactor. At the same time, a decrease in sulfate ion concentration was observed for this case. This means that in addition to calcium carbonate, calcium sulfate also precipitated from the solution. However, calculations show that almost all CO₂ introduced to the reactor was bound at higher initial calcium concentrations by the time the solution pH reached 12.

The minimum determined Ca⁺² concentrations, at a solution pH of approximately 9–9.5, were 0.00813 mol·dm⁻³, 0.0141 mol·dm⁻³, and 0.0222 mol·dm⁻³, and the initial calcium ion concentrations in the solution were 0.0212, 0.0363, and 0.0536 mol·dm⁻³, respectively. These values of Ca⁺² concentrations correspond to over 60% conversion of calcium ions in the solution, and the corresponding degree of CO₂ bonding was over 80%. However, for the highest value of calcium ion concentration, i.e., 0.0536 mol·dm⁻³, the loss of calcium was higher than the amount of introduced CO₂, and the concentration of sulfate ion was still lower than the initial value.

The effect of CO₂ concentration in the inlet gas directed into the reactor on the solution parameters during the precipitation process is illustrated in Figure 8a–d. The measurements were performed at a stirrer speed of 600 min⁻¹, for an initial concentration of calcium ions of 0.051 mol·dm⁻³. From the data presented in Figure 8a–d, it follows that, as in the previous cases, up to a solution pH of about 12, the concentration of calcium ions decreases steadily, while the concentration of the total amount of CO₂ dissolved in the solution practically does not change. The time taken to reach this pH value depends, of course, on the amount of carbon dioxide introduced into the reactor and ranges from 6 min (15.75 vol.% CO₂) to about 13 min (9.65 vol.% CO₂). During this time, the calcium ion concentration decreased to a value corresponding to about 50% calcium conversion, except in the case of the highest CO₂ concentration (40% calcium conversion). At that time, almost all of the carbon dioxide directed to the reactor was bound in the form of calcium carbonate, but only at the lowest CO₂ concentration, i.e., 9.65 vol.%, and no precipitation of calcium sulfate was observed.

The minimum concentrations of calcium ions determined in the solution were as follows: 0.0199 mol·dm⁻³, 0.0205 mol·dm⁻³, and 0.0161 mol·dm⁻³, and for the carbon dioxide concentrations in the inlet gas, 9.65 vol.%, 12.76 vol.%, and 15.75 vol.%, respectively, at a solution pH of approximately 8.8–9.4. For these values of Ca⁺² concentration, the degree of calcium conversion is 58.8%, 57.7% and 68.1%, respectively. At the same time, almost all of the carbon dioxide directed to the reactor was bound, but precipitation of calcium sulfate was also observed at the two higher CO₂ concentration values.

The average size (median value) of the obtained calcium carbonate particles ranges from 4.5 to 7.9 µm. In all the cases studied, as the pH of the solution drops below about 9 and calcium carbonate dissolves due to the increase in the concentration of HCO₃⁻ ions in the solution, the size of the obtained particles decreases. Figure 9 shows exemplary particle size distributions of calcium carbonate obtained after 10 min of precipitation (D_0.5 = 7.83 µm) and when the process was carried out until equilibrium was established (pH equal to approximately 6.5, D_0.5 = 5.87 µm). The measurements were taken at a stirrer speed of 600 min⁻¹, a gas mixture containing 12.8 vol.% carbon dioxide, and a solution with an initial calcium ion concentration of 0.0536 mol·dm⁻³. Detailed measurement results for the indirect carbonation studies are summarized in Table S1 in the Supplementary Materials.

Studies have shown that, in the indirect carbonation process, it is possible to utilize about 50–55% of the calcium contained in the solution, with complete binding of the carbon dioxide introduced into the reactor if the process is carried out to a solution pH of about 12. Continuing the process to a solution pH of about 9–9.5 allows for over 60% utilization of calcium, with binding of about 80–90% of carbon dioxide. For the highest value of calcium ion concentration in the solution and the highest value of the amount of CO₂ in the inlet gas, calcium sulfate starts to precipitate out of the solution in addition to calcium carbonate. Taking into account that about 77% of the calcium ions in the solution came from free calcium oxide/hydroxide contained in the ash, it can be concluded that about 65–70% of free calcium oxide/hydroxide can be bound as CaCO₃ during the indirect carbonation if the process is carried out to a solution pH of about 12.

The highest concentration of calcium ions in the solution equal to 0.0536 mol·dm⁻³ was achieved when the saturation state was reached during the dissolution process [26]. This means that when the tested fly ash was dissolved in water at a weight ratio of waste-to-solvent of 1:20, only 50% of calcium contained in free CaO, and Ca(OH)₂ was released into the solution. The extraction degree of calcium from free calcium oxides into water was approximately 100% at a waste-to-solvent weight ratio of 1:50, and in this case, the concentration of Ca⁺² in solution was about 0.047 mol·dm⁻³. Thus, in the indirect carbonation process, the maximum achievable degree of utilization of free calcium oxides contained in the ash is approximately 65–70%. Taking this into account, the actual average sequestration capacity of FBC in the indirect carbonation process under ambient conditions is about 62–67 g CO₂/1 kg of fly ash.

3.2. Direct Carbonation Process

During direct carbonation, all process steps, i.e., waste dissolution, CO₂ absorption, CO₂ reaction with CaO, and precipitation, take place simultaneously in one reactor. The final product of direct carbonation is a mixture of precipitated calcium carbonate, water-insoluble mineral/waste components, and alkali metal oxides/hydroxides that have not passed into the reaction medium.

Changes in solution pH during the direct carbonation process study for different values of the stirrer speed are presented in Figure 10a. Figure 10b–d show changes in the concentration of Ca⁺² and SO₄⁻² ions as well as the total amount of CO₂ dissolved in the solution occurring at that time. The lines are drawn only to guide the eye. The effect of the stirrer speed on the precipitation process was studied for a concentration of CO₂ in the stream directed to the reactor of 12.7 vol.%, with a weight ratio of waste (ash)-to-solvent (water) of 1:50.

From an analysis of the data shown in Figure 10a–d, it is apparent that during the first minutes of the process, both the pH of the solution and the calcium ion concentration increase. This means that the rate of ash dissolution and calcium ion extraction into the solution prevails over the rate of CO₂ absorption and calcium carbonate precipitation. Then the concentration of calcium ions decreases, and the pH of the solution slowly drops to about 12. During this time, which was 18.5, 17, and 16 min at a stirrer speed of 300, 600, and 1100 rpm, respectively, the concentration of CO₂ dissolved in water practically does not change, which means that almost all of the absorbed carbon dioxide has reacted with the calcium ions. By measuring the CO₂ concentration in the gas leaving the reactor, it was determined that about 90% of the carbon dioxide directed into the reactor was bound as calcium carbonate. This value corresponds to about 90% conversion of total calcium derived from free CaO and Ca(OH)₂ contained in the initial amount of fly ash.

After this time, the pH of the solution drops rapidly, the concentration of total CO₂ in the solution increases, and the concentration of calcium ions decreases further for a short time and then begins to increase. At this stage of the process, the increase in Ca⁺² concentration is due, among other things, to the fact that with the increase in the acidity of the solution, calcium carbonate dissolves due to the increase in the concentration of HCO₃⁻ ions. The minimum concentrations of calcium ions determined in the solution were as follows: 0.00805 mol·dm⁻³, 0.00803 mol·dm⁻³, and 0.0089 mol·dm⁻³, and for the stirrer speeds of 300, 600, and 1100 min⁻¹, respectively, at a solution pH of about 10–10.8. The times to reach the minimum observed Ca⁺² ion concentration were approximately 20, 19.5, and 18 min, respectively. During this time, about 95% of calcium contained in free CaO and Ca(OH)₂ reacted with CO₂ to form calcium carbonate, while the CO₂ bounding degree was over 85%. The concentration of sulfate ions increases steadily during the carbonation process (see Figure 10c) due to the continuous dissolution of CaSO₄, regardless of the value of the stirrer speed. The maximum value of the SO₄⁻² concentration depends only on the process running time. The slight effect of stirrer speed on the carbonation process observed in the direct carbonation studies proves that the free calcium oxides contained in the FBC dissolve fast enough not to slow down the CO₂ fixation process.

Changes in the pH of the solution during the direct carbonation process for different concentrations of carbon dioxide in the inlet gas are shown in Figure 11a. The corresponding changes in the concentrations of Ca⁺² and SO₄⁻² ions and the total amount of CO₂ dissolved in the solution are shown in Figure 11b–d. The studies were performed at a stirrer speed of 600 min⁻¹, for a weight ratio of waste-to-solvent of 1:50. It can be seen from Figure 11a–d that, as in the previous case, during the first minutes of the process, both the pH of the solution and the calcium ion concentration increase. In the next stage, the calcium ion concentration decreases, and the pH of the solution slowly drops to a value of about 12. Up to this point, the concentration of CO₂ dissolved in water practically does not change. The time to reach this pH value depends, of course, on the amount of carbon dioxide introduced into the reactor and ranged from 14 min (15.83 vol.% CO₂) to 22.5 min (9.64 vol.% CO₂). Calculations show that, within this time, more than 90% of the CO₂ directed to the reactor is bound as calcium carbonate, and the degree of calcium conversion is about 90% in relation to the calcium from free CaO and Ca(OH)₂ contained in the ash.

Then the pH of the solution drops rapidly, the concentration of total CO₂ in the solution increases, and the concentration of calcium ion decreases for a short time and then begins to increase. The minimum determined concentrations of calcium ions in the solution were as follows: 0.00946 mol·dm⁻³, 0.00803 mol·dm⁻³, and 0.00910 mol·dm⁻³ for the concentrations of carbon dioxide in the inlet gas of 9.64 vol.%, 12.69 vol.% and 15.83 vol.%, respectively, at a solution pH of about 9.7–10.7. The times to reach a solution pH of about 10 and the minimum observed concentration of Ca⁺² were about 25, 19.5 and 16 min for carbon dioxide concentrations in the inlet gas of 9.64 vol.%, 12.69 vol.% and 15.83 vol.%, respectively. After this time, the degree of calcium conversion is almost 100%, in relation to calcium from free CaO and Ca(OH)₂ contained in the ash. The corresponding degree of CO₂ bonding is about 90%. As in the previous case, the concentration of sulfate ions increases continuously during the carbonation process (see Figure 11c), regardless of the CO₂ concentration in the stream directed to the reactor. The maximum value of SO₄⁻² ion concentration depends only on the time of the process.

The influence of the waste-to-solvent weight ratio on the direct carbonation process was studied for a stirrer speed of 600 min⁻¹ and a CO₂ concentration in the inlet gas of 12.7 vol.%. The results are shown in Figure 12a–d. As can be seen from these figures, the course of the process is similar to the above-described cases. The time in which the rate of dissolution exceeds the rate of CO₂ absorption combined with precipitation of calcium carbonate, which is manifested by an increase in pH and Ca⁺² concentration in the solution, increases from 3.5 min to 10 min with the increase in the weight ratio of waste-to-solvent from 1:100 to 1:20. The times required for the pH of the solution to drop to approximately 12 were 8, 17 and 40.5 min, for waste-to-solvent weight ratios of 1:100, 1:50 and 1:20, respectively. During this time, over 90% of the CO₂ fed to the reactor is bound as calcium carbonate, and the calcium conversion rate is about 85% relative to the calcium derived from free CaO and Ca(OH)₂ contained in the ash. The minimum determined concentration of calcium ion in the solution was as follows: 0.00496 mol·dm⁻³, 0.00803 mol·dm⁻³, and 0.01662 mol·dm⁻³ for the weight ratios of waste-to-solvent of 1:100, 1:50 and 1:20, respectively, at a solution pH of about 9–10.4. The corresponding times to reach these values were about 9.5, 19.5 and 46.5 min. After these times, the calcium conversion degrees are 98.1%, 97.1% and 98.5%, respectively, in relation to the calcium from free CaO and Ca(OH)₂ contained in the waste. The corresponding degrees of CO₂ bonding are as follows: 89.6%, 87.2% and 92.4%, respectively.

As in the previous cases, the sulfate ion concentration increases steadily during the carbonation process (see Figure 12c), for each waste-to-solvent weight ratio tested. In this case, however, the maximum value of the SO₄⁻² concentration depends not only on the duration of the process, but also on the amount of ash introduced into the water. The sulfate ion concentrations at the end of the direct carbonation process were approximately 0.0073, 0.0113, and 0.0157 mol·dm⁻³ for waste-to-solvent weight ratios of 1:100, 1:50 and 1:20, respectively. This means that the amounts of CaSO₄ dissolved during the process were 80%, 63% and 35%, respectively.

The average size (median value) of the obtained direct carbonation product particles ranged from about 18 to about 20 µm. The analysis of the particle size distribution shows that the population of carbonized fly ash is not uniform. The resulting product is a mixture of particles of the order of a few µm and particles of approximately 20–30 µm. It is likely that the smaller particles are precipitated calcium carbonate and the larger ones are unreacted ash particles. At the same time, the median values fall between the sizes of calcium carbonate obtained in the indirect carbonation studies (4.5–7.9 µm) and the average size of FBC particles (24.64 µm). Figure 13 shows exemplary particle size distributions of the direct carbonation product obtained after 19.5 min of the process operation (D_0.5 = 20.19 µm) and when the process was carried out until equilibrium was established (pH equal to approximately 6.8, D_0.5 = 18.32 µm). The measurements were taken at a stirrer speed of 600 min⁻¹, a gas mixture containing 12.7 vol.% carbon dioxide, and a waste-to-solvent weight ratio of 1:50.

Detailed measurement results of the direct carbonation studies are summarized in Table S2 in the Supplementary Materials.

Studies have shown that, in the direct carbonation process, it is possible to utilize about 90% of the free calcium oxide contained in the studied FBC, with almost complete binding of the carbon dioxide introduced into the reactor (90–95% of CO₂) if the process is carried out to a solution pH of about 12. Continuing the process to a pH of about 10 allowed achieving a free calcium oxide conversion rate of about 94–97%, but the carbon dioxide binding rate was lower at about 85–90% in relation to the total amount of CO₂ introduced into the reactor. In the next stage of the process, when the pH drops below 10, an increase in the concentration of Ca⁺² ions is observed. This increase is caused, among other things, by the dissolution of calcium carbonate, due to the increase in the concentration of the HCO₃⁻ ion at a decreasing pH value of the solution. It is known from the literature that in the CO₂/H₂O system in the range of 6.5 < pH < 10.5, mainly HCO₃⁻ ions are present, and CaCO₃ is rarely produced [34]. The rising calcium ion concentration is also caused by further dissolution of ash, as evidenced by the continuous increase in the concentration of the SO₄⁻² ion. In all the cases studied, the concentration of sulfate ions increased during the process. The maximum value of SO₄⁻² ion concentration depends on the time of the process and the amount of ash introduced into the reactor. Of course, throughout the direct carbonation process, calcium oxides are also dissolving, and the carbon dioxide binding reaction is constantly taking place.

Although CaSO₄ does not have the potential to sequester CO₂, its presence in the ash studied appears to have a positive effect on the rate of this process. Calcium ions released during the dissolution of anhydrite increase the ionic activity product of calcium carbonate, which definitely accelerates the precipitation process. At the same time, as shown in [31], SO₄⁻² ions enhance the kinetics of CO₂ absorption. It should be emphasized that the presence of CaSO₄ in FBC does not prevent its use in cement production. The aim of the research presented in this paper was to determine the possibility of mineral carbonation using the tested FBC to remove free calcium oxide. The results show that in the direct carbonation process, it is possible to remove almost the entire amount of free calcium oxide. In the case of indirect carbonation, free CaO is removed during the ash dissolution process, which must be carried out properly to achieve the desired effect. The clear solution obtained in the first stage can then be used to bind CO₂ and obtain pure CaCO₃ in this way.

The actual average value of CO₂ sequestration capacity obtained in the studies is 86 g CO₂ per 1 kg of ash for direct carbonation if the process is carried out to a solution pH of about 12, and about 93 g CO₂ per 1 kg of ash if the process is carried out to a solution pH of about 10. These values are similar to those presented in the literature [7,35], but it is worth emphasizing that the studies presented in this paper were conducted under ambient conditions (20 °C, 0.1 MPa), while most of the described studies were conducted at elevated CO₂ pressure [19,36].

In direct carbonation studies, a slight influence of the stirrer speed on the process and the related changes in pH (Figure 10a) and concentrations in the solution (Figure 10b–d) were observed. The solubility studies [26] also showed a slight effect of mixing on the FBC dissolution kinetics. We can therefore say that water-soluble components of the ash (free calcium oxides and calcium sulfate) dissolve fast enough not to slow down the CO₂ fixation process.

4. Conclusions

Mineral carbonation is a promising technology of CO₂ capture, currently in the commercial implementation phase. Using industrial waste to capture carbon dioxide allows for lower sequestration costs and sustainable waste management [28,29]. The results of carbonation kinetics studies presented in this paper demonstrate that fly ash from lignite combustion in fluidized bed boilers can effectively bind CO₂, creating products with new properties that can be used.

The direct and indirect aqueous carbonation routes were considered. The main advantage of direct carbonation is the simpler technological scheme and the possibility to make better use of sorbents. The advantage of the indirect process is the possibility of obtaining a valuable commercial product, which is precipitated calcium carbonate. But the end-product from the direct carbonation process can also be used to fabricate construction materials, mainly due to the low content of free CaO. In addition, costs of the indirect carbonation process will be, probably, much higher than those of the direct process.

Based on the research carried out, the following final conclusions were formulated:

• The studies were conducted using a gas mixture reflecting the typical composition of coal-fired boiler exhaust gases at atmospheric pressure and ambient temperature (20 °C). The CO₂ absorption efficiencies obtained were similar to those reported in the literature. However, it should be noted that most studies on carbon dioxide mineralization employed higher CO₂ concentrations, gas pressures, and process temperatures.
• The indirect process achieved lower yields (from 55.5 ± 7.0 to 76.1 ± 7.1 g CO₂ per kg of ash) than the direct process (from 79.9 ± 5.6 up to 95.1 ± 7.3 g/kg).
• Carrying out the process under ambient conditions allows for achieving satisfactory efficiency and better economics compared to methods using higher CO₂ concentrations, pressure, and temperature often presented in the literature.
• The direct carbonation can be used to effectively remove free calcium oxide contained in FBC, enabling the use of this ash in the production of building materials.
• Mineralization is a method that combines the separation of CO₂ from flue gases and its safe storage. Despite its limited storage capacity, its use, even on a limited scale, seems economically viable, especially when industrial waste is used.
• The research results will be used to develop a comprehensive model of the carbon dioxide mineralization process, which will enable the economics of the project to be assessed.