Applied Demonstration of a kW-scale solid oxide fuel cell-calciner for power generation and production of calcined materials

(cid:129) A kW-scale solid oxide fuel cell-in- tegrated calciner was designed and constructed. (cid:129) SOFC-integrated calciner generates up to 2 kW el power. (cid:129) The high-grade heat generated is suf-ﬁ cient for full calcination of magne- site. Carbonate looping (CaL) has been shown to be less energy-intensive when compared to mature carbon capture technologies. Further reduction in the e ﬃ ciency penalties can be achieved by employing a more e ﬃ cient source of heat for the calcination process, instead of oxy-fuel combustion. In this study, a kW-scale solid oxide fuel cell (SOFC)-integrated calciner was designed and developed to evaluate the technical feasibility of simultaneously generating power and driving the calcination process using the high-grade heat of the anode o ﬀ -gas. Such a system can be integrated with CaL systems, or employed as a negative-emission technology, where the calcines are used to capture CO 2 from the atmosphere. The demonstration unit consisted of a planar SOFC stack, oper- ating at 750°C, and a combined afterburner/calciner to combust hydrogen slip from the anode o ﬀ -gas, and thermally decompose magnesite, dolomite, and limestone. The demonstrator generated up to 2 kW el,DC power, achieved a temperature in the range of 530 – 550°C at the inlet of the afterburner, and up to 678°C in the calciner, which was su ﬃ cient to demonstrate full calcination of magnesite, and partial calcination of dolomite. However, in order to achieve the temperature required for calcination of limestone, further scale-up and heat integration are needed. These results con ﬁ rmed technical feasibility of the SOFC-calciner concept for production of calcined materials either for the market or for direct air capture (DAC).


Introduction
The recent report by the Intergovernmental Panel on Climate Change (IPCC) has stressed the necessity of limiting global average warming to 1.5°C by 2100, compared to that of the pre-industrial level, to meet global climate ambitions and minimise the severe consequences of climate change [1]. Although following the 1.5°C emission pathway is not a geophysical impossibility yet, it requires substantial efforts to https://doi.org/10.1016/j.apenergy.2019.113731 Received 23 January 2019; Received in revised form 4 July 2019; Accepted 11 August 2019 potentially achieved by reducing the required amount of oxygen for the process, or eliminating the ASU. This can be accomplished by employing chemical looping [21,22], or using an alternative source of heat to drive the calcination process, including utilising heat-carrier sorbents that transfer heat from external sources [23,24], heat pipes [25,26], and heat transfer walls [27,28]. Alternatively, Ziock and Harrison [29] proposed that solid oxide fuel cells (SOFCs) could potentially be used to simultaneously generate electricity, and the high-grade heat required for the calcination process. In addition to the available high-grade heat, the anode off-gas of SOFCs (after combustion of tail gas) comprises mainly steam and CO 2 if the fuel contains hydrocarbons, that can be directly used to drive calcination and produce a pure CO 2 stream, once steam is condensed [30].
On the other hand, the SOFC-integrated calcination process can also be regarded as a negative-emission technology (NET), if the calcined materials can be used for direct air capture (DAC), in a relevant time frame. It has recently been demonstrated that dry and hydrated lime can achieve a high level of carbonation (70-75%) in a time-scale of weeks, when exposed to ambient air [31]. This time-scale is attractive from an engineering perspective, particularly when compared with the leading NETs such as bioenergy with CCS (BECCS) [32,33], and afforestation [34,35]. Hanak et al. [36] performed a techno-economic analysis on a 25 MW el,DC SOFC-integrated calciner for simultaneous electricity generation and DAC using limestone, dolomite, and magnesite. The system was estimated to operate at a net thermal efficiency of 43.7 (magnesite) − 47.7% LHV (limestone), and achieved a levelised cost of electricity of 50 £/MW el h, which is competitive with other low-carbon power generation technologies. Moreover, Hanak and Manovic [37] demonstrated the feasibility of such a system under uncertainties in market conditions, where there is no economic incentive for DAC, and the system can be used to produce lime to be sold in the market. Techno-economic analyses have indicated the SOFC-integrated calciner is a promising and efficient decarbonisation concept; however, there is need to further develop and demonstrate such systems.
In this study a kW-scale SOFC-integrated calciner was designed and demonstrated, and the technical and operational challenges were evaluated. The demonstrator was used to calcine a series of carbonates, including limestone, magnesite, and dolomite. The extent of calcination was assessed using thermogravimetric analysis (TGA). The morphological variations and surface elemental analysis were characterised using scanning electron microscopy (SEM) with an energy dispersive X-ray (EDX) spectrometer. The SOFC-integrated calciner can be simply integrated with the CaL process to further reduce the energy penalty of the system, or utilised as a combined heat and power system to generate electricity and produce metal oxide, such as lime and magnesia, to be sold in the market, or used in a DAC application (NET mode).

Concept description
A conceptual schematic of the SOFC-integrated calcination process is presented in Fig. 1. The system comprises three main components, Fig. 1. Conceptual schematic representation of SOFC-integrated calcination process with CO 2 removal from the air (DAC), using carbonate materials. namely SOFC stack, afterburner, and calciner. The main principle of the process is to use a SOFC stack to generate electricity, and high-grade heat to drive the calcination process. Although SOFCs can be directly fed with hydrogen and/or carbon monoxide, they are also capable of internal reforming when fed by natural gas or syngas [38]. However, internal reforming of hydrocarbons may result in carbon deposition and coke formation [39], which can cover the active side of the anode and block the microchannels [40]. To avoid coke formation within SOFCs, the reforming process can be performed externally. In general, external reforming is utilised for large-scale stationary applications, while internal reforming is employed for small-scale and portable applications to reduce the size and minimise the complexity of the systems [41]. It is recommended by the supplier (Ningbo SOFCMAN, China) that the stack used in this study needs to be operated with external reforming. Apart from mitigating carbon deposition, this also helps in avoiding feeding large amounts of steam into the stack, and increases the hydrogen partial pressure at the anode inlet to achieve a higher cell voltage [42]. When reformed gas is fed into the stack, the reformate products of methane (carbon monoxide and hydrogen), can be electrochemically oxidised on the anode side and generate electricity. However, the rate of electrochemical oxidation of carbon monoxide is 2-5 times lower than that of hydrogen. Thus, carbon monoxide mainly reacts with steam, produced due to oxidation of hydrogen or present in the feed stream, to form carbon dioxide and hydrogen through the water gas shift reaction (CO + H 2 O ⇌ CO 2 + H 2 ), and the electrochemical oxidation of carbon monoxide is negligible [43]. Therefore, since the primary aim of this work was to demonstrate the feasibility of electricity generation and high-grade heat utilisation for calcination processes, and also to avoid promoting carbon deposition and complexity of the experimental rig, no external reformer was used and the stack was directly fed with hydrogen.
SOFCs typically operate at 600-1000°C [44]. It is, therefore, expected that both anode and cathode off-gases leaving the stack are in the same temperature range. At commercial scale, the high-grade heat of the calciner and cathode off-gas can be used to preheat the air and fuel streams before feeding to the stack, as was demonstrated in our previous techno-economic studies [36,37]. However, in this demonstrator, the calciner off-gas was not utilised due to the design complexity of heat recovery at relatively small scales, i.e., relatively low flow-rates. In addition, in the SOFC stack used in this study, cathode offgas is released from the side of the stack into the hot-box (and subsequently into the atmosphere); thus, this heat could not be utilised, Fig.  S1 (Supplementary Information). Therefore, preheating of gas streams was performed using heating tapes. With regard to the anode off-gas, if the heat loss between the stack and afterburner is minimised, this stream could potentially experience a temperature rise upon combustion of the unused hydrogen, prior to the calciner.
After completion of the calcination processes, the calcined materials can be potentially transferred to a carbonator for decarbonation of flue gases (CaL) [16], sold in the market as commodities [37], or exposed to the ambient air to re-carbonate in a DAC application (NET) [31]. In addition, the concentrated CO 2 stream can be further conditioned for storage in geological formations [45], or utilised as a commercial product [46].

Material
Limestone (Longcliffe Ltd., UK), magnesite (Grecian Magnesite, Greece), and dolomite (Lhoist, UK) were of industrial grade, with a particle size of 100-500 µm. Nitrogen and hydrogen were supplied by BOC (UK) with a purity of higher than 99.999%. Air was provided using an air compressor.

Experimental setup
The schematic of the SOFC-integrated calciner demonstrator is presented in Fig. 2. The SOFC stack (Ningbo SOFCMAN, China) comprised 30 planar anode-supported cells, and had a configuration of Ni-YSZ/YSZ/CGO/LSCF-CGO, Fig. S1 (Supplementary Information). The cell size was 14 × 14 cm, with an active area of 150 cm 2 . A 3.5 kW furnace (Ningbo SOFCMAN, China) was used to keep the stack at the desired temperature in the initiation step. The furnace was automatically switched off during the demonstration. An external pressure load of~11.5 N cm −2 was applied to the stack and maintained throughout the test to ensure the stack was sealed properly. It should be noted that temperature variation during the test can affect the load. Thus, the pressure load needed to be monitored and readjusted to 11.5 N cm −2 , as necessary. During operation, a DC electronic load (4.8 kW, Elektro-Automatik, Germany) was used to dissipate generated electricity. In addition, the stack voltage was measured using a multimeter (Testo, UK). 1.2 kW and 2.1 kW (0.9 kW + 1.2 kW) heating tapes (OMEGA, UK) were initially used to preheat hydrogen and air, respectively, before introduction to the stack. It should be noted that during the stable operation step, the feed gases were also heated by the exothermic reaction in the stack, which reduced the required electric power input for preheating. Thus, it is important to ensure a highly responsive temperature control system is in place to avoid overheating of the feed pipes. Prior to the installation of heating tapes, the pipe was wrapped with mica tape (COGEBI) to avoid any potential electric discharge or short circuit during operation. Moreover, to avoid any potential short circuit between the cathode and anode sides of the stack, nylon tubing was used upstream of the pre-heater in the hydrogen line to limit continuity. The flow rates of entering air, nitrogen, and hydrogen into the demonstrator were measured and controlled using mass flow controllers (Alicat, UK). The reported flow rates are based on the operating conditions of each stream. An in-house afterburner/calciner (H × W × D of 650 × 500 × 500 cm) was designed and used to burn hydrogen slip from the anode off-gas, and further calcine the carbonate materials. The furnace was constructed from polished stainless steel and lined internally with 50 mm of ridged ceramic fibre board. Considering 60-80% hydrogen utilisation within the stack, associated with a Wobbe number of~3.4-7.8, and feed hydrogen flow rates of up to 100 L min −1 , a 5 kW up-fired radiant plaque burner with ceramic surface was used and situated at the base of the furnace, in order to provide an even distribution of temperature within the calciner trays. The calciner was a fixed bed, situated at the top of the furnace with the same dimensions (W × D of 500 × 500 cm), and consisted of two perforated trays located at 25 and 52 cm from the gas inlet of the burner, where the carbonate materials are placed. A flue damper (FD) was installed in the flue to adjust the amount of entering air into the afterburner. The composition of the off-gas was measured using a Fourier Transform Infrared analyser (FTIR, Protea, model FTPA-002). The safety system comprised a flame arrestor, located just before the hydrogen pre-heater, and an in-house shut-off system, including an automatic shut-off valve (ASV) in the hydrogen line, that was activated at the absence of flame in the afterburner, and/or once the temperature of the hotbox exceeded 950°C. In addition, a hydrogen detector (Testo 316-EX, UK) was used to identify any potential leakage. The developed demonstrator rig is presented in Fig. S2 (Supplementary Information).
Prior to reducing the stack, the hotbox was preheated to 750°C, at a rate of 1-3°C min −1 . Further, the anode side was purged with 4 L min −1 preheated nitrogen at 600°C for 10 min. The stack was then reduced by introducing 4 L min −1 preheated hydrogen at 600°C, and 12 L min −1 preheated air at 500°C into the anode and cathode sides, respectively. The reduction continued until the open circuit voltage (OCV) of the stack reached a plateau. Further, the demonstration was carried out by increasing the hydrogen and air flow rates to 18-43 L min −1 , and 54-150 L min −1 , respectively. Upon completion of the main tests, the stack was cooled to ambient temperature, at a rate of 1-2°C min −1 , and under 4 L min −1 of hydrogen flow.

Thermogravimetric Analysis (TGA)
The degree of calcination of the samples was measured using a thermogravimetric analyser (Pyris 1, Perkin Elmer). In each test, 20-30 mg of the material was heated from ambient temperature to 800°C (magnesite) or 900°C (limestone and dolomite) with a heating rate of 30°C/min under a nitrogen flow rate of 20 mL/min, and maintained at 800°C or 900°C for a further 5 min.

Scanning Electron Microscopy -Energy Dispersive X-ray Spectroscopy (SEM-EDX)
A scanning electron microscope (FEI XL30, Philips) with an energy dispersive X-ray spectrometer (JEOL 7800F) was used to assess the morphology of the samples, and analyse the elemental compositions of their surfaces, respectively. The images were taken at an accelerated voltage of 20 keV. Prior to scanning, the samples were coated with gold nanoparticles to prevent the accumulation of electrostatic charges.

Operating characterisation of SOFC stack
A series of experiments were designed to assess the performance of the SOFC stack with respect to power generation and operational challenges. In the first set of experiments, the reduction of the anode side was carried out at hydrogen and air flow rates of 4 and 12 L min −1 respectively, and continued for nearly three hours until the OCV of the stack reached a plateau of~34.1 V. Further, hydrogen and air flow rates were gradually increased to nearly 35 and 116 L min −1 , respectively, where the performance of the SOFC stack was evaluated. Fig. 3 presents the current-voltage characteristic of the stack, measured in the constant-current mode and at a rate of~0.6-1 A min −1 . At a stack current of 100 A, a stack voltage of 20.5 V was measured, indicating that the stack accomplished 2.05 kW el,DC (0.46 W cm −2 cell −1 ) power generation. It should be highlighted that during stable operation of the demonstrator, its electric power requirement averaged 1.0 kW. Therefore, although the demonstrator generates 2 kW of electricity, the net power was reduced.
Although the local temperatures at the top and bottom of the stack were similar,~750°C, during the reduction process, introduction of the current load led to elevation of the local temperature at the top bỹ 15°C, while no considerable change in temperature was observed at the bottom of the stack. It was also noted that an increase in hydrogen flow rates from 35 to 43 L min −1 further widened the top-to-bottom temperature gradient to 26°C. However, the top-to-bottom temperature difference did not exceed the maximum threshold of 50°C (recommended by supplier), as long as the hydrogen flow rate was kept below 43 L min −1 . The local temperature rise along the stack can be attributed to localised joule heating, arising from the conduction of electric current through the cell stack materials [47].
After completion of the first experiment and cooling down of the stack, it was found that a localised burn zone had formed at the front right side of the air exhaust, Fig. S3 (Supplementary Information). However, since no drop in performance of the stack was experienced during the operation, the experiment was repeated two more times to evaluate the importance of such physical damage on the effectiveness of the stack. The performance of the stack, by means of measured OCV and stack voltage, after formation of the burnt zone (first experiment) is presented in Fig. 4. Condition 1 refers to comparison of OCV of the stack, measured at hydrogen and air flow rates of 4 and 12 L min −1 , respectively, in three successive experiments. Condition 2 refers to measured stack voltage at a stack current of 8 A, and hydrogen and air flow rates of~35 and 116 L min −1 , respectively. It can be seen that there is a slight decline in the OCV of the stack (almost 2.5-3.3%) in each successive experiment. On the other hand, the stack voltage measured in a stack current of 8 A slightly increased in the second experiment, followed by a slight decrease in the third experiment. Therefore, it seems that the effect of the burnt zone on the performance of the stack at both high and low flow rates was almost negligible. However, longer experimental campaigns are needed to evaluate the effect of the burnt zone in the longer term. It was not possible to identify at what stage (namely reduction, operation, or cool down) the burnt zone was formed. However, it can be potentially associated with localised leakage and diffusion of fuel to the cathode side which can cause the formation of localised hot spots and, eventually, partial degradation of the cells. Leakage could occur due to the presence of nucleated and propagated cracks that are formed because of residual stresses during the heating or cooling processes [48,49]. In addition, hot spots can be potentially created by localised high current-density regions due to inhomogeneous consumption of the fuel across the cell [50]. Moreover, cracks can also be caused by physical damage during the assembly, transport or installation of the stack.

Calcination of carbonates
The solid lines in Fig. 5 present the thermogravimetric analysis of carbonates, including magnesite, dolomite, and limestone under nitrogen flow, indicating the range of required temperatures for calcination. Therefore, the afterburner/calciner was designed to provide the required temperature, assuming 60-80% fuel consumption in the stack and up to 10% excess air for combustion. In the afterburner/calciner, the carbonates could be placed in two trays located at 25 cm (ST 1 ) and 52 cm (ST 2 ) from the inlet, and their temperatures, TC 1 and TC 2 , respectively, were monitored, Fig. 2. In addition, the temperature of the anode off-gas stream entering the afterburner was indicated by TC AF,I , Fig. 2.
It was seen that, when the stack was fed with hydrogen and air flow rates of~35 and 116 L min −1 , respectively, measured TC 1 and TC 2 were as high as 449°C and 377°C, respectively. This implied that neither tray had sufficient temperature for the calcination of dolomite, limestone and magnesite. In addition, TC AF,I was around 535°C, which was more than 200°C lower than the operation temperature of the stack (~750°C). The hydrogen flow rate was then gradually increased to 42 L min −1 , while the air flow rate was kept constant, in order to slightly increase the hydrogen slip at the anode off-gas. As a result, both TC 1 and TC 2 increased up to 542°C and 426°C, respectively; however, there was no considerable change in TC AF,I (~545°C). This implies that, if a larger stack is used, a higher temperature in the calciner can potentially be experienced. It was noted that further increase in the hydrogen flow rate resulted in a sudden increase in the local temperature at the top of the stack. Thus, the hydrogen flow rate was kept to a maximum of 43 L min −1 , to limit top-to-bottom temperature gradient along the stack to below 50°C. Further, it was noted that the oxygen concentration in the flue was high (~17%), indicating massive uptake of ambient air by the afterburner. Therefore, a flue damper was utilised to reduce the amount of air entering the afterburner. Consequently, the oxygen concentration was lowered to 12%, which led to an increase in TC 1 and TC 2 to as high as 678°C and 464°C, respectively. In addition, almost no change in TC AF,I (~530°C) was observed. The oxygen concentration of~12% was the lowest possible that could be achieved, and any further reduction by closing the flue damper resulted in an 'abrupt cessation of combustion'. Accordingly, 150 g of carbonates was placed in tray 1 (ST 1 , Fig. 2), where the highest temperatures were experienced, to proceed with the calcination process.
The extent of calcination of the carbonates tested in the afterburner/ calciner was assessed using TGA (Fig. 5 with the dashed lines). With regard to magnesite (Fig. 5a), no weight loss was observed in the TGA curve, implying that the sample was fully calcined in the SOFC-integrated calciner (demonstrator). This result is also confirmed by EDX analysis, Table 1, in which both magnesite samples calcined in the demonstrator and TGA had almost the same chemical composition,  with no trace of C element (presence of MgCO 3 ) in either of them. Moreover, the SEM images of both calcined magnesite samples in the TGA (Fig. 6a.2) and the demonstrator (Fig. 6a.3) appeared to be almost identical, and characterised with a less dense and more open structure compared to magnesite (Fig. 6a.1), which is attributed to the lower molar volume of MgO compared to that of MgCO 3 . Fig. 5b (solid line) presents the full thermal decomposition of dolomite to oxides. Although thermal decomposition of dolomite is a complex process and occurs through a series of mechanisms, overall it can be considered as a two-stage process. The first stage (here 350-700°C) can be associated with dehydration, and decomposition of dolomite to CaCO 3 and MgO, while in the second stage (here 700-900°C), CaCO 3 is calcined to CaO [51]. In comparison, TGA data for the dolomite sample tested in the demonstrator (Fig. 5b, dashed line) showed no significant weight loss up to~700°C, indicating that the demonstrator was capable of providing the temperature required to decompose dolomite to CaCO 3 and MgO. Further, although there was a steeper reduction in weight loss for calcined dolomite in the demonstrator (dashed line) compared to the fully calcined dolomite sample in TGA (solid line) over the temperature range of 700-750°C, eventually both curves were almost parallel over 750-900°C, where the calcination of CaCO 3 occurred. This indicates that the demonstrator was incapable of the full thermal decomposition of CaCO 3 . In addition, EDX analysis, Table 1, showed that in comparison to dolomite, the concentration of O element of the calcined sample in the demonstrator was reduced, while the concentration of Mg and Ca increased, implying that calcination occurred to some extent. Moreover, the SEM image of the dolomite fully decomposed in the TGA, featured (Fig. 6b.2) low crystallinity which can be associated with decomposition of dolomitic calcite. On the other hand, the thermally decomposed dolomite in the demonstrator (Fig. 6b.3) showed a different structure, characterised by an amorphous phase, which may be related to poorly crystallised calcite formed due to immediate carbonation of nascent CaO crystals [52], and small MgO grains grown outwards and at the surface of the dolomite phase [53].
The calcination extent of limestone is presented in Fig. 5c. It can be seen that the TGA data for limestone sample tested in the demonstrator (dashed line) are very similar to those of fully calcined sample (solid line), with only~11% difference in weight loss at 900°C. This implies that calcination of limestone in the demonstrator was very limited. On the other hand, although it was expected that the EDX analysis of limestone tested in the demonstrator would be more similar to the limestone chemical composition, instead, it was found that there was more similarity to the fully calcined limestone. Moreover, comparing the SEM images of limestone ( Fig. 6c.1) and fully calcined limestone (Fig. 6c.2), with the limestone sample tested in the demonstrator (Fig. 6c.3), it can be seen that there are some small CaO grains around unreacted CaCO 3 grains, which may relate to the occurrence of limestone calcination at low temperatures, but with slow kinetics [54]. In addition, the smaller grain size of CaO produced in the demonstrator compared to that of the fully calcined sample is attributed to thermal decomposition at lower temperature [55]. The developed SOFC-integrated calciner demonstrated the technical feasibility of full and partial calcination of magnesite and dolomite, respectively. However, due to its relatively low cost, marked demand for lime, and the utilisation in carbonate looping cycles, calcination of limestone is still more attractive. Therefore, further improvements are needed to ensure the required temperature for the calcination of limestone in the calciner can be achieved (starting from~840°C for steam-rich streams [30]). The first improvement may be achieved by increasing the anode off-gas temperature entering the afterburner. As was noted, although the operating temperature of the stack was 750°C (recommended by supplier), the anode off-gas temperature at the inlet of the afterburner was almost 200°C lower, despite the connecting pipe being well-insulated. This is mainly attributed to the heat losses at small scale, due to low flow rates. It was mentioned previously that some SOFCs can operate at up to 1000°C [44] and, therefore, another solution can be to use a stack that can operate at higher temperature. However, this may lead to an increase in the degradation rate of stacks over the long term. In addition, scaling up the process means that larger SOFC stacks, operating at higher fuel flow rates should be utilised, which in turn would minimise the temperature difference between the stack exit and afterburner inlet. Moreover, once the system is scaled up, the feed gas stream can be preheated using recovered waste heat, which results in an increased net power being generated. The second improvement can be implemented by optimising the excess air fed to the afterburner. Therefore, the current afterburner can be replaced by an oxy-fuel boiler with flue gas recirculation to minimise the heat losses, required in order to produce concentrated CO 2 ready for compression and sequestration.

Conclusions
A kW-scale solid oxide fuel cell (SOFC)-integrated calciner was designed and constructed to explore the technical feasibility of simultaneous power generation and calcination of carbonates. Such a system not only can be regarded as an alternative for conventional calciners, where the required heat for the thermal decomposition of carbonates is provided by oxy-fuel combustion, but also, can be used as a negativeemission technology for electricity generation, if the produced calcined materials are utilised for direct air capture. The SOFC stack was successfully tested and generated up to 2 kW el,DC power. However, it was found that the temperature of the anode off-gas stream entering the afterburner/calciner was around 530-550°C, at least 200°C lower than the operating temperature of the stack (~750°C). In addition, the maximum temperature achieved in the calciner, after combustion of the hydrogen slip of the anode off-gas, was up to 678°C. Consequently, it was demonstrated that the SOFC-integrated calciner was capable of full and partial calcination of magnesite and dolomite, respectively. However, the extent of calcination of limestone was very low. Further improvement of such a system, by means of providing the temperature required for the calcination of limestone in a steam-rich stream, can be achieved with a scaled-up system, using larger SOFC stacks, operating at higher temperatures (up to 1000°C). In addition, replacing the current air-fired afterburner with an oxy-fuel boiler with flue gas recirculation, which is required in order to produce concentrated CO 2 stream, and optimisation of oxygen excess ratio can address the heat losses experienced in the demonstrator and enable reaching the temperature needed for calcination of limestone.