The calcium-looping process for advancing in the development of both C02 capture and thermochemical energy storage systems.

1 The Calcium-Looping process for advancing in the development of both CO2 capture and thermochemical energy storage systems Author Carlos Ortiz Domínguez Supervisors Prof. Ricardo Chacartegui Ramírez Prof. José Manuel Valverde Millán A Thesis submitted to the Universidad de Sevilla for the degree of DOCTOR OF PHILOSOPHY Programa de Doctorado de Ingeniería Energética, Química y Ambiental Escuela Técnica Superior de Ingeniería Universidad de Sevilla July 2018 2 3 Abstract The Calcium-Looping process is based on the multicycle calcination-carbonation of calcium carbonate, which is one of the most abundant materials in the Earth. The process shows promising results facing to solve two of the main challenges within the future energy scenario: CO2 capture from fossil fuel combustion processes and energy storage in renewable-based plants. This thesis is focused in the assessment of the Calcium-Looping allowing a better understanding of the process integration for both applications. Post-combustion CO2 capture based on Calcium-Looping process consist of reacting the CO2 present in the flue gas exiting the fuel power plant with CaO particles according the carbonation reaction, which produces solid CaCO3 as product. CO2 capture efficiencies over 90% have been proved at MWth pilot-scale. After carbonation, solids are sent to a separated reactor where CaO is regenerated according the endothermic calcination reaction. The heat to carry out the reaction is provided by fuel oxy-fuel combustion to ensure an almost pure CO2 stream exiting the reactor, which after a purification process is ready to be stored or used in another process. The regenerated CaO is sent again to the carbonator reactor for a new CO2 capture cycle. On the other hand, by integrating the Calcium-Looping as thermochemical energy storage system in solar thermal power plants, the process starts performing the calcination reaction from concentrated solar power. CO2 and CaO streams produced are stored separately for later use. When energy is demanded, both components are brought together to the carbonator where the stored energy is released by the carbonation reaction. The carbonation is highly exothermic and therefore a proper integration of the heat released (i.e. for power production) is fundamental for the process efficiency. In spite that both applications are based on the same process, the specific operation conditions required leads to a different behaviour which suggest that the process must be analysed in detail according to each application. This thesis is focused on the development of models and process integration schemes based on lab-scale results under specific conditions selected from the analysis of the potential industrial scale implementation. It involves a multidisciplinary approach combining chemical reactions and process engineering to advance in the response of the challenges posed. This document is structured as follow: the first section introduces the reader to the Calcium- Looping applications, stressing the research opportunities that have motivated the present thesis. This leads to the formulation of the objectives of the work, which are addressed in the presentation and the discussion of the main results of the work. After the conclusions, new research lines to be faced in the near future are summarized. 4 5 Acknowledgements There are many people to whom I would like to thank their collaboration in this thesis. First of all, to my supervisors, Prof. Ricardo Chacartegui and Prof. Jose Manuel Valverde, for giving me opportunity to develop my PhD and for their invaluable support during these three intense years. Also thank Dr. Jose Antonio Becerra for his help in this period. I am fully grateful to Dr. Matteo Romano for giving me the opportunity of spending a fruitful research stay at the Politecnico di Milano. Thanks to Dr. Marco Binotti for his assistance. I gratefully acknowledge Prof. Luis Pérez-Maqueda and his team, Antonio, Pedro, Mónica and Beatriz, for their collaboration within this edifying work environment. I would like to thank especially my colleagues Jesús, Pilar, Elisa, Fran, Reyes, Irene and Nabil for making going to work every day so fun. Thanks to Victor, Alessandro, Davide and Giuseppe for the fruitful collaboration. Thank the department of Electronic and Electromagnetism for hosting me during these years. Thanks to Mª Jesús and Silvia for their help with administrative tasks. On a more personal note, I would like to thank my parents for their support and confidence. Thanks to my brothers. Special thanks to Belen, for her unconditional love. I apologize for all the hours that I have not dedicated to you during the realization of this work. Thanks to Guille, our precious baby. 6 7 List of publications This thesis is presented as a compendium of publications according to Article 9 of the regulations governing the regime doctoral thesis (Agreement 9.1/CG 12/04/19) within the framework regulated by the RD 1393/2007. The thesis is supported by several original publications developed by the author and his co-authors during the thesis period. The author has explicit authorization by all his co-authors to use their content as part of this thesis. A complete copy of the publications is presented in Annex 1-10. Moreover, as part of the work carried out in this thesis, several contributions have been made to national and international congresses as well as the application for two patents. Papers in indexed journals ANNEX 1 Ortiz C, Chacartegui R, Valverde J, Becerra J, Perez-Maqueda L. A new model of the carbonator reactor in the calcium looping technology for post-combustion CO2 capture. FUEL 2015;160:328–38. doi:10.1016/j.fuel.2015.07.095. ANNEX 2 Ortiz C, Chacartegui R, Valverde JM, Becerra JA. A new integration model of the calcium looping technology into coal fired power plants for CO2 capture. Appl Energy 2016;169:408– 20. doi:10.1016/j.apenergy.2016.02.050. ANNEX 3 Chacartegui R, Alovisio A, Ortiz C, Valverde JM, Verda V, Becerra JA. Thermochemical energy storage of concentrated solar power by integration of the calcium looping process and a CO2 power cycle. Appl Energy 2016;173:589–605. doi:10.1016/j.apenergy.2016.04.053. ANNEX 4 Ortiz C, Valverde JM, Chacartegui R. Energy Consumption for CO2 Capture by means of the Calcium Looping Process: A Comparative Analysis using Limestone, Dolomite, and Steel Slag. Energy Technol 2016:1–12. doi:10.1002/ente.201600390. ANNEX 5 Alovisio A, Chacartegui R, Ortiz C, Valverde JM, Verda V. Optimizing the CSP-Calcium Looping integration for Thermochemical Energy Storage. Energy Convers Manag 2017;136:85–98. doi:10.1016/j.enconman.2016.12.093. 8 ANNEX 6 Ortiz C, Chacartegui R, Valverde JM, Alovisio A, Becerra JA. Power cycles integration in concentrated solar power plants with energy storage based on calcium looping. Energy Convers Manag 2017;149:815–29. doi:10.1016/j.enconman.2017.03.029. ANNEX 7 Ortiz C, Valverde JM, Chacartegui R, Benítez-Guerrero M, Perejón A, Romeo LM. The Oxy-CaL process: A novel CO2 capture system by integrating partial oxy-combustion with the Calcium- Looping process. Appl Energy 2017;196:1–17. doi:10.1016/j.apenergy.2017.03.120. ANNEX 8 Bonaventura, D., Chacartegui, R., Valverde, J.M., Becerra, J.A., Ortiz, C., Lizana, J., 2018. Dry carbonate process for CO2 capture and storage: Integration with solar thermal power. Renew. Sustain. Energy Rev. 82, 1796–1812. https://doi.org/10.1016/j.rser.2017.06.061 ANNEX 9 Ortiz C, Valverde JM, Chacartegui R, Perez-Maqueda LA. Carbonation of Limestone Derived CaO for Thermochemical Energy Storage: From Kinetics to Process Integration in Concentrating Solar Plants. ACS Sustain Chem Eng 2018;6:6404–17. doi:10.1021/acssuschemeng.8b00199. ANNEX 10 Ortiz C, Romano MC, Valverde JM, Binotti M, Chacartegui R. Process integration of Calcium- Looping thermochemical energy storage system in concentrating solar power plants. Energy 2018;155:535–51. doi:10.1016/j.energy.2018.04.180. In addition, the author has collaborated in other papers not directly related to the subject of the thesis but developed within the research group framework. For this reason, these publications are not included in the annexes: - Soltero VM, Chacartegui R, Ortiz C, Velázquez R. Evaluation of the potential of natural gas district heating cogeneration in Spain as a tool for decarbonisation of the economy. Energy 2016:1–20. doi:10.1016/j.energy.2016.06.038. - Lizana J, Ortiz C, Soltero VM, Chacartegui R. District heating systems based on lowcarbon energy technologies in Mediterranean areas. Energy 2017;120:397–416. doi:10.1016/j.energy.2016.11.096. - Lizana J, Chacartegui R, Barrios-Padura A, Ortiz C. Advanced low-carbon energy measures based on thermal energy storage in buildings: A review. Renew Sustain Energy Rev 2018;82:3705–49. doi:10.1016/j.rsder.2017.10.093. - Soltero V, Chacartegui R, Ortiz C, Lizana J, Quirosa G. Biomass District Heating Systems Based on Agriculture Residues. Appl Sci 2018;8:476. doi:10.3390/app8040476. 9 - Soltero VM, Chacartegui R, Ortiz C, Velázquez R. Potential of biomass district heating systems in rural areas. Energy 2018;156:132–43. doi:10.1016/j.energy.2018.05.051. Participation in conferences 1) Soltero VM, Chacartegui R, Ortiz C, Velázquez Vila R, Becerra Villanueva JA. Evaluation of the potential for Natural Gas District Heating Cogeneration in Spain as a tool for economy decarbonization. 10th Conf. Sustain. Dev. Energy, Water Environ. Syst., Dubrovnik (Croacia): 2015. 2) Ortiz C, Chacartegui R, Valverde JM, Alovisio A, Becerra JA. Calcium looping process for Thermochemical Energy Storage in CSP plants. Integration assessment of direct and indirect power cycles. 11th Conf. Sustain. Dev. Energy, Water Environ. Syst., 2016, p. 1–29. 3) Lizana J, Chacartegui R, Barrios-Padura Á, Ortiz C, Vilches A. Evaluation of thermal energy storage technologies for heating, cooling and hot water applications road to zero energy buildings. 11th Conf. Sustain. Dev. Energy, Water Environ. Syst., Lisbon: 2016. 4) Lizana J, Serrano A, Ortiz C, Becerra Villanueva JA, Chacartegui R. Energy assessment methodology towards low-carbon energy schools applied to Mediterranean buildings. Conf. Sustain. Dev. Energy, Water Environ. Syst., Dubrovnik , Croatia: 2017. 5) Ortiz C, Fernandez R, Chacartegui R, Valverde JM, Becerra JA. Almacenamiento termoquímico en plantas CSP basado en calcium-looping: retos y oportunidades. XVI Congr. Ibérico y XII Congr. Iberoam. Energía Sol., Madrid: 2018. 6) Soltero VM, Chacartegui R, Ortiz C, Quirosa G. Renewable biomass low temperature heating networks for rural locations. ECOS 2018 - 31ST International. Conference on Efficiency, cost, optimization, simulation and Environmental impact of Energy Systems, Guimaraes (Portugal): 2018. 7) Ortiz C, Valverde JM, Chacartegui R, Romeo LM, Perez-Maqueda LA. The mOxy- CaL process: integration of membrane separation, partial oxy-combustion and Calcium Looping for CO2 capture. 21st Conf. Process Integr. Model. Optim. Energy Sav. Pollut. Reduct., 2018. 8) Fernandez R, Ortiz C, Chacartegui R, Valverde JM, Becerra JA. Thermochemical Energy Storage for enhancing the dispatchability in PV plants. 13th Conf. Sustain. Dev. Energy, Water Environ. Syst. - SDEWES Conf., 2018. 16 𝑊󰇗𝑆𝑇 power produced in the steam turbine cycle, MWe 𝜌𝑔 density of gas phase, g/m3 𝑊󰇗𝑃 power consumed in the steam turbine cycle, MWe 𝜂 overall net efficiency 𝑊󰇗𝑃𝑆𝑂𝐿𝐶𝐴𝐿 Power consumptions for solids transport in the calciner side, MWe 𝜂𝑏𝑜𝑖𝑙𝑒𝑟 boiler efficiency 𝑊󰇗𝑃𝑆𝑂𝐿𝐶𝐴𝑅 Power consumptions for solids transport in the carbonator side, MWe 𝜂𝐶𝐹𝑃𝑃 coal fire power plant efficiency 𝑊󰇗𝐴𝑈𝑋𝑃𝑂𝑊𝐶𝐴 auxiliary power consumptions in the calciner side, MWe 𝜂𝑟𝑒𝑓 reference plant efficiency 𝑊󰇗𝐴𝑈𝑋𝑃𝑂𝑊𝐶𝑅 auxiliary power consumptions in the carbonator side, MWe 𝜂𝑖𝑛𝑡 new global efficiency (CFPP-capture system) 𝜙𝑁 particle fraction in the N cycle Δ𝐻𝑅(𝑇𝑟𝑒𝑎𝑐𝑡) reaction enthalpy at the reactor temperature, kJ/mol 𝜏 average residence time in the carbonator, s Δ𝐻𝑅0 standard reaction enthalpy kJ/mol 𝜌𝐶𝑎𝑂 CaO density, g/m3 ∆𝑆20 carbonation entropy change, J/(mol·K) CaL Calcium-Looping SPECCA Energy consumption per kilogram of CO2 avoided CSP Concentrating Solar Power CAPEX Capital Expenditures TCES Thermochemical Energy Storage PCC Post-Combustion Capture CFPP Coal fired power plants CCS CO2 Capture and Storage CFB Circulating Fluidized Bed TES Thermal Energy Storage PCM Phase Change Material 𝑆𝑀 Solar Multiple 17 1. Introduction CaCO3 is the second most abundant material on Earth after water. The Calcium-Looping (CaL) process is based on the multicyclic calcination-carbonation of CaCO3 according Eq. (1). Calcination is a well-known reaction which is at the basis of the cement industry [1] while the CaL process has many other applications. As early as 1933, the use of CaO-based materials was patented for CO2 capture to enhance H2 production from methane reforming [2]. Later on, the CaL process began to be considered as a potential Thermochemical Energy Storage (TCES) system in the late 1970s [3–5]. More recently, the CaL process has been widely studied for postcombustion CO2 Capture (PCC) in fossil power plants [6–8]. 𝐶𝑎𝐶𝑂3(𝑠)⇄𝐶𝑎𝑂(𝑠)+𝐶𝑂2 ∆𝐻𝑟0=178𝑘𝐽 𝑚𝑜𝑙 (1) The CaL the process is fully aligned with some of the most important challenges in the energy scenario. Thus, the CaL process can be applied in both renewable and fossil-fuel based power plants as energy storage and CO2 capture systems, respectively. In the former case, a promising option is the integration of the CaL process in Concentrating Solar Power (CSP) plants to improve the dispatchability of the system. CSP plants, as well as other plants based on renewable resources, are characterized by a variable energy input depending on the non-continuous nature of the solar resource and, when the solar resource is available, variability of the weather conditions. By integrating the CaL process as TCES system, direct solar radiation is used to carry out the calcination reaction (endothermic), releasing CO2 and CaO as products that are stored separately. Storage conditions and time are flexible and could be accommodated to the energy demand [9]. When energy is needed, the stored products are brought together to carry out the carbonation reaction (exothermic), by which the stored energy is released. In comparison with currently commercial energy storage systems, such as solar molten salts, the CaL process presents several benefits such as its feasibility to work at significantly higher power cycle temperatures, a high energy storage density and the possibility to store energy in the medium-long term. Regarding CO2 capture, the CaL process is highlighted as one with higher potential for postcombustion capture. The standard cycle for CO2 capture from flue gas streams uses lime (CaO), which can be derived from limestone calcination, to produce CaCO3 by carbonation at high temperature (~650ºC). Typical CaL conditions for CO2 capture involve carbonation at relatively low CO2 partial pressure (about 0.15 atm) and calcination at very high temperatures (around 950ºC) under high CO2 partial pressure with short residence times at both stages. Once CO2 is captured in the carbonator and heat from the exothermic reaction is recovered, the almost CO2 free flue gas is released into the atmosphere. Pilot plants (of size on the order of 1-2 MWth) have 18 recently demonstrated the achievement of CO2 capture efficiencies around 90% [10,11] whereas model simulations predict an efficiency penalty on power generation around 3-9% when scaling up the technology to a commercial level [12]. Because of the huge potential scope for the CaL process within the future energy system, in both renewable and non-renewable power plants, and in new or retrofitted facilities, a fast technology development with a relevant cost reduction is expected. Moreover, it must be highlighted the very favorable characteristics of the CaO precursors such as limestone or dolomite: low cost, wide availability and non-toxicity [13–15], which are necessary conditions for the massive integration on a large scale of any CO2 capture or energy storage technology. 1.1. Calcium-Looping process for post-combustion CO2 capture Finding competitive and sustainable solutions for capturing and sequestering the CO2 released from fossil fuel combustion processes is a key issue to mitigate global warming [16]. Despite the necessary increase in energy production share from renewable sources, power plants based on natural gas or coal currently generate most of the energy consumed in the world [17]. Currently, 76.5% of the electricity generation in the world is produced by non-renewable sources [18]. It is expected that the commercial deployment of CO2 capture and storage (CCS) technologies plays a main role within the portfolio of urgent measures to limit global warming below 2°C from preindustrial levels [19,20]. On this decarbonization purpose, future fuel-based power plants must be near to CO2 emissions free. Coal fired power plants (CFPP) are prime candidates to be retrofitted with CCS. The main R&D challenge for the viability of CFPPs and other fossil fuelbased facilities is to capture CO2 by means of feasible, sustainable and affordable technologies while, at the same time, the penalties of integrating these technologies on power production and efficiency are minimized. For a real deployment of CO2 capture technologies, it must be reduced both the energy consumption associated to the CO2 capture process and the CO2 capture system size. The former is directly related to the electricity production cost. Typical values reported on energy consumption per kilogram of CO2 avoided (SPECCA) are around 4 MJ/kg CO2 [21] using the already mature amines based technology. On the other hand, the CO2 capture system size affects to capital expenditures (CAPEX). Post-combustion capture (PCC) refers to CO2 removal from the flue gas of fossil fuel-based power plants, which can be accomplished by using chemical solvents, solid sorbents or electrochemical processes. Despite post-combustion capture processes are being widely investigated in the last years, Boundary Dam (100MWe) in Canada and Petra Nova (240MWe) in USA are currently the only commercial CFPP that applies CCS by using a chemical absorption process based on amines. 19 In chemical absorption, the solvent (typically an amine solution such as MEA) binds chemically with the CO2. Amine absorption and stripping consists of passing the post-combustion flue gas through an aqueous amine solvent, which absorbs CO2 by chemical reaction [22]. Then, the solvent loaded with CO2 (the “rich” solvent) is heated up above typically 120 ℃ in the regenerator reactor wherein the CO2-amine chemical reaction is reversed to release nearly pure CO2 and regenerate the amine. The so-called “lean” solvent is recycled back to the absorber to restart the process while the released relatively pure CO2 is compressed to a suitable pressure for an efficient transportation and storage [23]. After regeneration, the solvent is cooled to be reused [24]. Aminebased PCC can efficiently remove around 90% of the CO2 emissions. A main issue of systems based on amine absorption is the large amount of heat required to regenerate the solvent [25–27]. This heat, which is usually obtained from the steam cycle, penalizes significantly (8-12%) the power plant efficiency. Moreover, the commercial expansion of amine-based systems for CO2 capture at large scale is hindered by amines toxicity [28], degradation [29] and equipment corrosion [30]. Thus, there is a need for developing novel largescale capture technologies with reduced energy penalty and capture cost, which should rely on environmentally friendly and abundantly available cheap materials. Another option for PCC is based on membrane separation, which uses the pressure difference between the flue gas and the removed CO2. The membrane technology is generally useful to treat high-pressure gases [31,32] in spite of which a large number of researches have adapted it for post-combustion capture [33–35]. Regarding efficiency penalty associated to membranes use for PCC, it is estimated in the range of 4.9-8.5% [31]. Membrane separation is a promising solution to reduce the costs of PCC. However, the maximum pressure ratio attainable by feed compression and/or permeate vacuum is limited to approximately 10, due to cost and energy considerations [33]. The development of dry CO2 capture processes based on cheap materials operating at relatively low temperatures, which would require relatively low energy for sorbent regeneration, is considered as a promising pathway to advance in the deployment of CO2 capture technologies [36,37]. Na2CO3 is the sorbent employed in the Dry Carbonate Process (DCP) early proposed in [38,39] and currently being demonstrated at the pilot-scale stage [40]. An important advantage of the DCP is that sorbent regeneration would be carried out at 100-200ºC and therefore the process could be efficiently assisted by medium temperature solar thermal power. Another proposed technology for PCC is the use of electrochemical processes in Molten Carbonate fuel cells. Some studies show that electricity generation in the fuel cell partially compensates the penalty on the original cycle in wastewater treatment plants [41] and power plants [42–44]. 20 The CaL process is at the root of a promising 2nd generation technology for post-combustion CO2 capture at coal fired power plants. The potential application of the CaL process is increasingly gaining momentum in the last years mainly due to the possibility of implementing it in already existing power plants and the relatively low penalty on plant performance as compared to other already mature capture technologies such as conventional amine-based capture systems (penalty around 8-12 %) [7,13,45]. Figure 1 shows a conceptual scheme of the CaL process applied for post-combustion CO2 capture. The CaL process consists of the capture of CO2 in the post-combustion flue gas by carbonation in a high temperature fluidized bed reactor of CaO solids derived from the calcination of natural limestone (CaCO3). The process would be carried out in two interconnected circulating fluidized bed (CFB) reactors, both operated under atmospheric pressure at gas velocities of about 5 m/s [46,47] wherein gas-solid contact and heat/mass transfer are promoted. Another CaO precursor could be natural dolomite. According to recent lab-scale studies [48], dolomite-derived CaO would exhibit a greater capture capacity than limestone-derived CaO at CaL conditions for CO2 capture. On the other hand, several synthetic sorbents have been developed with enhanced capture behavior [49,50]. Nevertheless, the use of these synthetic sorbents usually leads to an important increase of the cost of the system operation [12]. Taking into account that flue gases exiting from coal-fired power plants generally contain a mole fraction of CO2 in the range 10-15% [51,52], optimum carbonation temperatures are around 650°C in order to achieve significant efficiencies of CO2 capture (around 80-90%) in short residence times, typically on the order of minutes [53]. 21 Figure 1: Conceptual scheme of the CaL process for post-combustion CO2 capture. Modified from [54] Once CO2 is captured in the carbonator and heat from the exothermic reaction is recovered, the decarbonized flue gas is vented into the atmosphere. The carbonated solids are calcined in a second fluidized bed reactor (calciner) at higher temperatures, which regenerates CaO for a new cycle and produces a highly concentrated CO2 stream ready for compression and sequestration. CaO is regenerated in the calciner by calcination under necessarily high CO2 partial pressure, which requires the use of rather high temperatures (> 930ºC) for efficient calcination in a short residence time [10,55–57]. The endothermic calcination reaction and the drop of temperatures between the sorbent streams entering and leaving the calciner involves the need of providing a high-energy input to the calciner. The most feasible method to generate the heat required in the calciner and avoid CO2 dilution at the same time is the burning of fuel by oxy-combustion [10,14,55–57]. Thus, calcination is carried out in an environment containing between 70% and 90% vol. concentration of CO2 at atmospheric pressure [55,58–61]. After calcination, CO2 is retrieved from the calciner for compression, transport and geological storage or other uses. Under these operating conditions, calcination temperatures above 930ºC are necessary to regenerate the sorbent in short residence times [62–64]. These conditions promote sintering of CaO which compromises the multicyclic CO2 capture capacity [65]. CaO deactivation is one of main issues to be addressed in the CaL process [66]. It is well-known that carbonation occurs in two consecutive phases. The first stage takes place on the free surface HEAT EXCHANGER BOILER COAL AIR Q STEAM POWER CYCLE ELECTRIC POWER flue gas CARBONATOR CaO (s) + CO2 (g) CaCO3 (s) ΔHr0 =-178 kJ/mol Q SECOUNDARY STEAM POWER CYCLE CALCINER CaCO3 (s) CaO (s) + CO2 (g) ΔHr0 =+178 kJ/mol COAL O2 (ASU) PURGE CaO/CaCO3 pure CO2 COMPRESSION STORAGE REUTILIZATION clean gas CaL CFPP 22 of the solid through the nucleation and growth of a CaCO3 layer and it is governed by the kinetics of the reaction between CaO and CO2. The end of the fast stage takes place when a 40–60 nm thickness product layer is formed, which makes inaccessible a large fraction of CaO in the interior of the particles [67]. Then it takes places a second stage controlled by the solid-state diffusion 𝐶𝑂32− and 𝑂2− ions through the product layer, which is a much slower process. TGA tests carried out under calcination at high CO2 partial pressure show that solid-state diffusion plays an important role in the total CaO conversion since capture in the fast reaction controlled regime is greatly reduced with the number of cycles [48,63]. By using limestone under typical CaL conditions for post-combustion CO2 capture, the multicyclic CaO conversion decays with the cycle number up to reach a residual value 0.07-0.08 for residence times of about 5 min [68,69]. CaO deactivation is partially compensated by the periodic introduction of a fresh limestone makeup in the calciner while the purged CaO is particularly well suited for cement production [70,71]. CaO deactivation is further enhanced by irreversible CaO sulphation and ashes due to insitu coal oxy-combustion [6,63,66,72]. To reduce the make-up flow, labscale tests suggest that a possibility to enhance the CO2 capture efficiency could be the recarbonation of the partially carbonated particles in another reactor before the calcination step by using the pure CO2 stream exiting the calciner [73,74]. Regarding the CFPP-CaL integration, several models have been proposed in the literature aimed at predicting the efficiency penalty that arises from integrating the CaL technology in commercial power plants [75]. These models consider integration of CO2 capture into coal fired power plants (CFPP) [6,14,76,77], lignite fired power plants (LFPP) [56], natural gas combined cycle power plants (NGCC) [78] and integrated gasification combined cycle power plants (IGCC) [79–82]. Depending on the power plant type, most results show efficiency penalties in the range 3-9% [83], which is lower than in the MEA-based process. Minimizing energy penalties is a key point for the industrial competitiveness of the CaL process. A number of process schemes and configurations to enhance the energy integration have been proposed in the literature [83]. CaL integration studies are generally built upon equilibrium models for the calciner reactor [6,76,84]. As regards the carbonator, previous works have considered equilibrium models [79,85], semi-predictive models using the Kunii-Levenspiel formulation for the fluidized bed hydrodynamics [6,64] or an average conversion model [84]. Several carbonator reactor models have been already proposed to predict the CO2 capture efficiency under CaL conditions for CO2 capture [47,77,86,87]. These models have been inferred from thermogravimetric analysis (TGA) results in which calcination is usually carried out under low CO2 partial pressure, which does not conform to the expected operation conditions in the real system. Experimental results show that 23 the multicyclic capture capacity of limestone is strongly influenced by the conditions under which calcination is performed [58,88,89]. There are several pilot-scale facilities testing the CaL process for post combustion CO2 capture. The larger plants are La Pereda (1.7 MWth) [10,90], Darmstadt (1 MWth) [11] and Hualien (1.9 MWth) [91], with the rest of pilot plants in the range of 1-120 KWth [75]. Most of the reactors are based on Fluidized Bed technology, either Circulating or Bubbling Fluidized Bed. 1.1.1. Challenges in the road to the deployment of CO2 capture technologies Each one of the above reviewed PCC technologies show specific advantages but also challenges to overcome at their different R&D development stages. Remarkably, PCC is considered as the most appropriate technique to be applied in the short-term for its relatively easy integration in existing fossil fuel power plants [26]. PCC integration penalizes power plant performance, and this hampers indirectly the global CO2 emissions reduction. Main drawbacks that hinder the deployment of the CaL process are the high cost of the full CCS chain and the high efficiency penalty imposed on the power plant. As discussed above, diverse alternatives have been analyzed for mitigating the efficiency penalty. A main inconvenient is that sorbent regeneration requires a large amount of heat at high temperatures (900–950°C), which must be properly recovered to increase the process efficiency. Improving heat integration in the CaL process to reduce energy requirements requires from the two main strategies: recovering energy in the CaL loop and modifying the CaL configuration to avoid (or reduce) oxyfuel combustion in the calciner [12]. CaO deactivation is also an important issue, leading to a large amount of inert solids being recirculated through the plant well above the stochiometric needs. Moreover, there are engineering challenges related to circulating solids at high temperature. Further barriers are the financing of CO2 transport infrastructure, legal and regulatory frameworks and insurance for safe permanent CO2 storage or utilization [36]. A part of this introduction to CO2 capture was published in [92], where the author of the thesis was in charge of the review sections. The full manuscript including complete information about a number of CO2 capture technologies is presented in Annex 8. 24 1.1.2 Recently finished and on-going projects An intense R&D activity is being carried out to assess the feasibility of the CaL process for postcombustion CO2 capture. Some of the recently finished and on-going relevant projects are: 1. SCARLET: Scale-up of Calcium Carbonate Looping Technology for Efficient CO2 Capture from Power and Industrial Plants (2014-2017) The SCARLET project 1 , funded by the European Union’s Horizon 2020 research and innovation programme and recently finished, aimed at validating the feasibility of the CaL technology in an upgraded 1 MWth pilot plant. Successful operation of the pilot plant and the optimization of the process parameters will provide essential information for scaling-up the process to a future 20 MWth pilot plant. The project provided a techno-economic and environmental assessment of the CaL process potential for CO2 capture from power plants as well cement and steel production plants. Several works have been published within this project [11,93]. 2. CEMCAP: CO2 capture from cement production (2015-2018) CEMCAP is a project funded by the European Union’s Horizon 2020 addressing CO2 capture from cement production by integrating the CaL process. The primary objective of CEMCAP is to prepare the ground for large-scale implementation of CO2 capture in the cement industry 2 . CEMCAP aims to leverage to Technology Reediness Level (TRL) 6 for cement plants with a targeted capture rate of 90% and develop techno-economic decision-basis for CO2 capture implementation in the cement industry. 3. CLEANKER: Clean clinker production by Calcium Looping process (2017-2021) This project is funded by the European Union’s Horizon 2020 research and innovation programme. In the same line that CEMCAP project, CLEANKER project aims at demonstrating at TRL7 the CaL concept in a configuration highly integrated with the cement production process, making use of entrained flow reactors. The core activity of the project is the design, construction and operation of a CaL demonstration system in the cement plant operated by Buzzi Unicem sited in Vernasca (Piacenza, Italy) 3 . 4. FLEXICAL: Development of flexible coal power plants with CO2 capture by Calcium Looping (2016-2019) FLEXICAL project, funded by the European Research Fund for Coal and Steel, aims to develop novel Calcium looping systems able to respond to a wide range of flue gas load 1 http://www.project-scarlet.eu 2 https://www.sintef.no/projectweb/cemcap/ 3 http://www.cleanker.eu/home-page-it 25 changes 4 . For this, two process options are tested at pilot-scale: a highly load flexible plant concept and an energy storage using CaO/CaCO3 silos. A recent publication about this project can be found in [94]. 5. LEILAC: Low Emissions Intensity Lime & Cement (2016-2020) This project is funded by the European Union’s Horizon 2020 research and innovation programme. The LEILAC project is based on a technology called Direct Separation, which aims to enable the efficient capture of the unavoidable process emissions from lime and cement production 5 . 1.2 Calcium-Looping as thermochemical energy storage process The commercial expansion of renewable energy technologies is an urgent need to limit global warming below 2.0°C by 2100 and pursue 1.5°C above pre-industrial levels as agreed at Paris COP21 Conference [20]. Among renewable energy technologies, Concentrated Solar Power (CSP) has a great potential for a full commercial expansion [95]. However, for renewable energies to achieve full autonomy from fossil fuels and to increase their feasibility, a main hurdle to overcome is their inherent variability in production. Thus, efficient and low cost energy storage stands as one of the main technological challenge for large deployment of renewable energies [96–98]. Moreover, large-scale energy storage is essential for a global system with high penetration of solar energy in order to increase the electric grid flexibility and to avoid risks derived from transient peaks [99]. Regarding the necessary dispatchability of renewable power facilities, CSP plants show several advantages over solar PV and wind due to the relatively low cost and feasible integration of thermal energy storage technologies in large-scale facilities compared to battery storage systems [100–102]. Efficient, cheap and non-toxic thermal energy storage (TES) is a key issue for Concentrating Solar Power (CSP) plants to provide a significant share of electricity generation. In recent years, a number of thermal storage technologies for medium to high temperature CSP systems have been developed based upon three main concepts: sensible thermal energy storage (TES), latent heat storage and thermochemical energy storage (TCES) [103,104]. Sensible heat storage systems are the most mature technologies [105] and they involve the use of different materials with high heat capacity such as water [103], molten salts [106–109], mineral oils [110] or ceramic materials [111]. Commercial plants of considerable size (above 100 MWt) do already 4 https://www.flexical.eu/ 5 https://www.project-leilac.eu/ 32 temperatures under superheated steam are considerably decreased down to ~700-750°C [163], which could be a further advantage for a quick integration of the CaL technology in CSP plants. Furthermore, the reactivity of CaO regenerated by calcination under superheated steam is notably enhanced [164], which would expectedly serve to boost the efficiency of the CSP-CaL integration. In addition, superheated steam for calcination in the solar reactor would be directly available from already possible in-situ generation of superheated steam in solar receivers [165]. Nevertheless, a higher energy penalty is expected by carrying out the calcination under superheated steam because of the solar energy consumption to bring the water (after separation) up to the calcination conditions. Another possibility to reduce the calcination temperature is by reducing the CO2 partial pressure introducing Helium in the calciner, which allows that calcination occurs at 725ºC [166] because of the high thermal conductivity and CO2 diffusivity in He. Nevertheless, Helium/CO2 separation brings a new challenge that must be further analyzed. Remarkably, the reduction of calcination temperature could highly reduce the reradiation losses in the solar receiver. While the particle receiver and high-temperature solids conveying represent a design challenge for the development of the CSP-CaL integration, many other equipment of the plant are currently widely developed on an industrial scale, mainly from the cement and lime industry [1,7,70]. This is the case of the cyclonic gas-solid heat exchangers [167], which could be used in the CSP-CaL integration to preheat the CaCO3 entering the calciner. Efficient gas-solid contact and heat/mass transfer in the calciner and carbonator reactors could be ensured by the use of circulating fluidized-beds (CFB), which are typically operated under the fast fluidization regime with gas velocities of the order of 5-10 m/s [46,168]. CFB reactors are characterized by their contrasted efficiency and durability and their suitability to the CaL process is already proven in CO2 capture pilot plants [10,11]. Separation of the calcination products prior to storage is feasible using commercially available devices such as gas coolers and cyclones to separate the solids from the CO2. The CO2 released in the calcination reaction must be compressed at high pressure to minimize the storage volume [169–171]. This brings about a remarkable loss of efficiency in the global power cycle. In this regard, the option of performing intercooling compression to minimize the energy penalty must be assessed. One of the main drawbacks of the CaL process is the progressive loss of activity towards carbonation in short residence times of the regenerated CaO as the number of cycles increases depending on the carbonation/calcination operating conditions [12,68,172]. Thus, only a fraction X of the total flow of CaO entering the carbonator reacts to produce CaCO3, while 1-X remains as unreacted CaO. The average CaO conversion X is thus a critical material property for the simulations. CaO conversion is highly dependent on the carbonation-calcination conditions as well as on the CaO precursor [173]. The CaL process applied to post-combustion CO2 capture, 33 where most of the CaL research has been focused on in recent years, involves carbonation under low CO2 partial pressure whereas calcination is carried out under high CO2 concentration at temperatures around 950ºC. On the other hand, the CO2 concentration in the gas stream entering the carbonator is imposed (15% v/v of CO2 in the flue gas from a coal fired power plant). These conditions lead to a severe drop of CaO conversion with the number of cycles, reaching a residual value of just around X=0.07-0.08 [48]. However, CaL conditions for TCES in CSP must not be necessarily the same as those employed for CO2 capture. Since the concentration of CO2 entering the carbonator is not imposed, carbonation conditions can be chosen to minimize the negative impact of CaO deactivation while at the same time the thermoelectric efficiency is enhanced. 1.2.4 Research lines and on-going projects Thermochemical energy storage, solar calcination and high-temperature solar receivers are gaining the attention of many research projects. As described above, the CSP-CaL process has a number of advantages which are fully aligned with the current research lines about CSP plants, namely: i) increasing the plant efficiency; ii) LCOE reduction; iii) increasing the dispatchability in CSP plants and iv) improving the sustainability and the environmental impact [174]. Several projects (recently finished or on-going) related with the CSP-CaL integration are: 1. SOCRATCES: Solar calcium-looping integration for thermo-chemical energy storage (2018-2020) The SOCRATCES project, funded by the European Union’s Horizon 2020 research and innovation programme, aims at demonstrating the suitability of this integration by erecting a pilotscale plant that will reduce the core risks of scaling up the technology as well as allow further understanding of the process integration. The longer-term goal is enabling highly competitive and sustainable CSP plants. The SOCRATCES concept stems from laboratory results Technology Readiness Level (TRL) of 4 to test the concept in a relevant environment (TRL5). Main expected results are: i) prototype demonstration of capacity for energy storage; ii) successful carbonator design with possibility to scale-up and iii) successful calcination at prototype scale by means of flash calcination technology. 2. Carbon Dioxide Shuttling Thermochemical Storage Using Strontium (2014-2015) The project was awarded within the ELEMENTS (Efficiently Leveraging Equilibrium Mechanisms for Engineering New Thermochemical Storage) program, supported by the SunShot Initiative (DOE). The project uses safe, and non-corrosive strontium-based carbonates and high temperatures from concentrated sunlight to break chemical bonds and store energy during the day time. The use of strontium carbonate allows the system to be charged via the reversible 34 decarbonation reaction at solar energy input temperatures of 1300°C, allowing discharge of the stored solar energy at very high temperatures of over 1000°C. The project explored sintering inhibitors. Relatively stable materials supported by Yttria- Stabilized Zirconia (YSZ) or SrZO3 were identified as the leading candidates. 3. Regenerative Carbonate-Based Thermochemical Energy Storage System for Concentrating Solar Power (2014-2016) Within the ELEMENTS program, the Regenerative Carbonate-Based Thermochemical Energy Storage System for Concentrating Solar Power project aims to develop a TCES for CSP based on reversible gas-solids reactions of carbonate and silicate sorbent-based at bench-scale. The modified calcium carbonate sorbents under investigation by SRI for solar TCES performed hundreds of cycles in simultaneous TGA/DSC equipment to measure both mass and heat change as the material undergoes repeatedly high temperature carbonation/decarbonation reactions. Results showed that the highly-refined and tailored reinforced CaO sorbent developed exhibits a very stable multicycle behavior with a stable sorbent capacity around 0.3, although maybe conditioned by the expense of a considerable increase of material costs and technical issues in the Ca-based synthetic material fabrication. 4. Demonstration of High-Temperature Calcium-Based Thermochemical Storage System for use with Concentrating Solar Power Facilities (2014-2018) This project, funded by the APOLLO program (DOE), aims to demonstrate a novel hightemperature calcium-based thermochemical storage system for use with CSP facilities, which is directly related to the CSP-CaL system described throughout the present thesis. The system uses a highly refined and tailored reinforced CaO sorbent undergoing a reversible carbonation reaction in a parallel-plate heat exchanger reactor. Main objectives of the project are [175]: i) to optimize the CaO sorbent composition and to demonstrate the required high capacity and negligible degradation over a large number of cycles at temperatures around 750ºC; ii) to design, build and test bench-scale heat exchanger reactor system to demonstrate the sorbent capacity; iii) to carry out a techno-economic evaluation to optimize heat exchanger reactor system and balance of plant design; iv) to build and validate heat exchanger reactor design multiphysics mathematical models and v) to design, build and demonstrate field scale systems (1MWh). 5. CSP2: Concentrated solar power in particles (2011-2015) The CSP2 project, funded by the European Union’s Horizon 2020 research and innovation programme, proposes to use dense gas-particle suspensions (around 50% of solids) in tubes irradiated by CSP as HTF. The dense phase is composed of a granular solid that resists temperatures up to 750ºC, which overcomes the limit imposed by molten salts systems, with a 35 70-90 % efficiency. It was developed and operated successfully a pilot solar receiver with 100- 150 kW thermal capacity placed on the focus of the 1 MW solar furnace of the CNRS. The solar loop was tested for 8 hours, with a flow rate of about 0.7 to 1.8 tonnes per hour. The solar receiver consisted of 16 1m long vertical tubes in which the dense gas-particle suspension (DPS) flowed vertically upward. 6. TCSPower: Thermochemical Energy Storage for CSP Plants (2011-2015) The overall objective of the TCSPower project, funded by the European Union’s Horizon 2020 research and innovation programme, is to build a new, efficient, reliable and economic thermochemical energy storage (TCS) for concentrated solar power plants based on the decomposition of calcium hydroxide, which allows storing energy in a temperature range between 450 and 550°C. For the higher temperatures CSP plants the redox reaction of manganese oxide is evaluated. A simulation tool for the design of TCS reactors with improved heat and mass transfer characteristics is used to identify suitable reactor concepts for the hydroxide and the redox reaction system. Both concepts are being experimentally evaluated at laboratory scale. Additionally, an up-scaling to 10kW will be realized for the more promising reaction system to evaluate the performance of a pilot-scale TCS reactor. 7. SOLPART: High temperature Solar-Heated Reactors for Industrials Production of Reactive Particulates (2016-2019) This project is also funded by the European Union’s Horizon 2020 research and innovation programme. The main objective of the SOLPART project is to develop, at the pilot scale, a high temperature (950°C) 24h/day solar process suitable for supply totally or partially the thermal energy requirement for CaCO3 calcination by high temperature solar heat. The system will operate at 950°C and will include a 30 kWth solar reactor producing 30 kg/h CaO and a 16h hot CaO storage [176]. The project develops and merges three advanced technologies: high temperature solar reactor, transport of high-temperature solid materials and high temperature thermal storage. 8. NEXT-CSP: High Temperature concentrated solar thermal power plan with particle receiver and direct thermal storage (2016-2020). This project is also funded by the European Union’s Horizon 2020 research and innovation programme. The project proposes a fluidized particle-in-tube concept. It aims to further develop the technology to make it rapidly cost-efficient and ready to be introduced in the market. A cost reduction by 38% is expected with respect to current CSP electricity cost. A 4-MWth tubular solar receiver able to heat particles up to 800°C will be constructed and tested. The rest of the loop, a two-tank particle heat storage and a particle-to-pressurized air heat exchanger coupled to a 1.2 MWe gas turbine, will also be tested [177]. 36 1.3 Motivation of the thesis and research opportunities The motivation of the thesis is to advance in the knowledge on the CaL process for both applications: post-combustion CO2 capture (PCC) and thermochemical energy storage (TCES). As discussed throughout the previous sections, the CaL process has potential advantages regarding the state-of-the-art of the technology, i.e. amines-based CO2 capture system and molten salt-based CSP plants, respectively. The CaL process is already widely studied for CO2 capture in fossil fuel power plants, as TCES system in renewable power plants and to enhance H2 production from methane reforming. Either one of these applications requires particular reaction conditions to which the sorbent performance (reaction kinetics and multicycle conversion) is extremely sensitive [12]. Therefore, specific models based on the conditions of any application are needed. For post-combustion CO2 capture, promising preliminary results have led to the proliferation of a large number of studies in recent years about the CaL integration in non-renewable power plants being remarkable the successful operation of 1-2 MW pilot plants for a large number of hours, such as La Pereda [10] or Darmstadt [93]. However, several challenges must be solved for a largescale deployment of the CaL technology. Nowadays, research is mainly focused on enhancing the CaO multicyclic activity, improving the reactors design, improving heat integration in the plant and reducing both CAPEX and OPEX to increase the process profitability. A proper understanding of the reaction mechanisms is a key point for designing the reactors, whose operating conditions (pressure, temperature, size, gas velocity, etc.) will determine the CO2 capture efficiency. It requires specific accurate models. The multicyclic calcination/carbonation behavior of CaO used in models found in literature [47,77,86,87] has usually been inferred from thermogravimetric analysis (TGA) results in which calcination is often carried out under low CO2 partial pressure due to technical limitations related to insufficiently fast heating/cooling rates of conventional ovens. Under these conditions, most of the carbonation reaction in short residence times occurs on the surface of the solids through a reaction controlled fast phase whereas the subsequent solid-state diffusion-controlled phase is comparatively negligible. Recent TGA studies demonstrate that carbonation on the surface of the particles in the reaction controlled fast phase is extraordinarily hampered when the solids are regenerated under high CO2 pressure in accordance to realistic conditions. On the other hand, carbonation in the subsequent solid-state diffusion phase gains a substantial relevance [173,178]. The integration of CaL process as post-combustion CO2 capture system in CFPP would involve energy penalty ranges between 4 and 7% points over the reference plant efficiency, which is less 37 than using MEA-based systems. Nevertheless, CaO deactivation with the number of cycles remains as a shortcoming of the CaL system, which causes an energy consumption that could be reduced if CaO rapid deactivation is mitigated. An important effect of CaO deactivation is that a larger equipment size is required because of the massive presence of non-reacting solids in the system. Improving the CaO multicycle conversion and process integration could reduce the equipment size and therefore the capital cost, which currently hinder the deployment of CO2 capture technologies. Thus, further knowledge is needed to: a) Evaluate the CaL process efficiency under realistic conditions, similar to those that would be obtained when applying the CaL process in a power plant at a commercial scale. It includes the development of models for the reactions to estimate the capture efficiency taking into account the carbonation in the solid-state diffusion carbonation stage. Results from these novel models can affect considerably the optimal conditions that until now have been considered for the operation of the CaL process, namely, calcination and carbonation temperatures of 900ºC and 650ºC respectively with a particles residence time in both reactors of around 5 minutes. b) Develop new integration schemes. Thus, if the reactor design changes by considering more realistic conditions, the carbon capture process scheme should also vary to be adapted to the optimal operating conditions. Therefore, not only novel reactor models are needed but also new CFPP-Cal integration process schemes. c) Moreover, from the analysis of the penalties in the power plant performance when incorporating the CCS system, synergies between main CO2 capture technologies, namely, oxy-combustion, membranes and post-combustion (either CaL or amines) can be detected, which makes interesting the analysis of hybrid processes. This requires the study of each one of these processes to identify the optimum integration and operation conditions to improve the global process. To this end, lab-scale test and modelling are needed to identify the synergies and impact. Regarding the integration of the CaL process in CSP plants, Section 1.2 discusses the huge potential of the CaL process as TCES system. The number of Concentrating Solar Power (CSP) plants is notably increasing throughout the world in the last 10 years, reaching a total installed capacity worldwide 5 GWe [121]. However, for a massive deployment of CSP facilities several challenges need to be addressed. The CSP-CaL process is fully aligned with the priority research lines in CSP technology, namely, enhancing dispatchability, increasing plant efficiency and cost reduction [174]. Dispatchability can be improved taking advantage of the high storage energy 38 density of the CaCO3/CaO system. Because the carbonation reaction occurs at temperatures as high as 890ºC in a pure CO2 atmosphere at ambient pressure (and even higher by increasing the reactor pressure) a higher temperature of heat introduced in the thermodynamic cycle and therefore a higher cycle efficiency than in molten salt-based plants can be achieved. Moreover, cost and sustainability of CSP plants can be enhanced due to low cost, wide availability and nontoxicity of natural CaO precursors (i.e. limestone or dolomite). The interesting opportunities for the integration of the CaL process as TCES system require new knowledge focused on: d) Analysis of the feasibility of the process equipment and systems to be integrated at industrial scale. e) Reactors modelling under TCES realistic operation conditions. CaL process conditions for thermochemical energy storage are not identical to those in CO2 capture. Thus, optimum reactors temperature and pressure or atmosphere composition varies from one application to another and therefore kinetics, sorbent multicyclic conversion, and physical properties must be analyzed from specific lab-scale tests. f) Process schemes to optimize the energy integration. A large number of feasible schemes for integrating the CaL process in CSP plants are possible. This includes configurations considering low temperature tanks for long-term energy storage but also high temperature storage in configurations similar to commercial ones with daily storage. g) Power cycles integration assessment. The CaL process allows operating at higher temperatures than in current CSP plants. To take advantage of this, a proper selection and configuration of the power cycle must be addressed. h) Finally, it is remarkable that most of the current research on integrating the CaL process for both CO2 capture and TCES considers limestone as CaO precursor because of the important advantages of using this natural compound. However, the use of other materials with interesting characteristic should be analyzed. These research opportunities are addressed through the present thesis. The next section describes in detail specific objectives according to the research lines described above. 39 2. Objectives In the previous section the context, framework and opportunity for the development of this thesis has been presented. Following, the specific objectives addressed along the thesis are described. The present thesis aims to contribute to the CFPP-CaL integration research with the following objectives: 1) Developing models based on lab-scale tests using realistic process conditions. From the analysis of full-scale plant operation, a set of operation conditions should be defined to be later reproduced at lab-scale. Results of lab-scale tests can be used in the development of a novel carbonator model which considers realistic CaO regeneration conditions. The new model would allow a more accurate evaluation and prediction of the carbonator performance, which is relevant for the development of process configurations and reactors design for the integration in real power plants. This objective intends to respond to the research opportunity (a) discussed in the previous section. 2) Improving the process scheme to reduce the energy consumption in realistic CaL operation. CO2 capture efficiency and heat integration within the CFPP-CaL plant should be designed by considering the important role of the diffusion-controlled carbonation phase, which becomes relevant when CaO regeneration is carried out under high CO2 partial pressure as is the case in the CaL process for CO2 capture. This objective intends to respond to the research opportunity (b) discussed in the previous section. 3) Evaluating the performance of several natural CaO precursors, such as limestone, dolomite, and steel slag. Even though there is a large number of synthetic sorbents that are being currently analyzed within the CaL process, the present thesis is focused on the study of sustainable, environmental-friendly, low cost and widely available precursors as required to make possible the CaL large scale deployment. Experimental results on their multicycle capture capacity behavior at realistic CaL conditions are needed to evaluate the feasibility of using these CaO precursors. Furthermore, the energy penalty that arises from the integration of the CaL process into a coal fired power plant using these CaO precursors should be addressed. This objective intends to respond to the research opportunity (h) discussed in the previous section. 4) Developing a new hybrid process based on the CaL technology which could improve state-of-the-art configurations. Hybrid processes, which take advantage of synergies 40 between different CO2 capture systems (such as CaL, oxy-combustion or membranes technology), could reduce the energy consumption and the equipment size associated to integrate CO2 capture in fossil-based power plants. This objective intends to respond to the research opportunity (c) discussed in the previous section. Regarding the development of the CaL process as TCES system, the present thesis aims to contribute to the CSP-CaL integration research through the following objectives: 5) Developing models based on realistic operation and multicyclic behavior of CaO within TCES schemes. Thus, a specific carbonation model for TCES should consider optimal integration conditions of the CaL process within CSP facilities, such as reducing the receiver temperature and increasing, as much as possible, the carbonation temperature. Moreover, carbonation under high pressure is an interesting option to improve the power cycle efficiency in TCES schemes, which requires the analysis of how the carbonation (kinetics, sorbent deactivation, etc.) is affected by increasing the reactor pressure. This objective intends to respond to the research opportunity (e) discussed in the previous section. 6) Performing a critical assessment about the advantages and challenges of developing the CaL process for high temperature thermochemical energy storage. This objective intends to respond to the research opportunity (d) discussed in the previous section. 7) Gaining knowledge about the possibilities of integrating the CaL process in both new and operating CSP plants. Power production from the stored energy by means of a Brayton CO2 closed cycle could improve both the reliability and the power plant efficiency regarding previously proposed schemes, which are based on air open schemes. This objective intends to respond to the research opportunity (d) discussed in the previous section. 8) Exploring the integration with the TCES core system of alternative direct and indirect cycles (steam turbine, closed Brayton CO2 and indirect-supercritical CO2) for relevant CSP-CaL integration conditions. This objective intends to respond to the research opportunity (g) discussed in the previous section. 9) Enhancing the CSP-CaL plant efficiency by improving heat integration. An interesting possibility of the CaL process is to store energy at ambient temperature for long term energy storage. However, due to the high temperatures in both the calciner and carbonator reactors, a low temperature storage involves a large streams temperature change along the entire cycle, which makes crucial an optimized heat integration to achieve an adequate system efficiency. This objective intends to respond to the research opportunity (f) discussed in the previous section. 41 10) Another possibility is to store the solids at high temperature. In this case the process scheme is simplified by requiring fewer heat exchangers due to a lower temperature difference between the reactors and storage. On the other hand, long term storage cannot be considered because of the temperature losses. However, the CSP-CaL system could be operated under a solar multiple (defined as the ratio of the solar thermal power to the power block design thermal input) in a similar way than in currently CSP plants. The development of specific schemes for this configuration could reach a compromise between the overall plant efficiency and heat integration complexity. This objective intends to respond to the research opportunity (f) discussed in the previous section. 11) Assessing the technologies to be used within the CSP-CaL scheme, with special focus on solar receiver and solids management. This objective intends to respond to the research opportunity (d) discussed in the previous section. 48 As the reaction evolves, the carbonation rate is determined as a function of conversion degree X, and reaction temperature T and pressure P. Experimental results on the time evolution of conversion show the typical sigmoidal shape of autocatalytic processes and are well fitted by a Prout−Tompkins model function [189] modified by introducing as conversion limit the CaO conversion at the end of the reaction controlled phase: 𝑑𝑋 𝑑𝑡=𝑋 (1−𝑋 𝑋𝑘) 𝑟(𝑇,𝑃) ↔𝑋(𝑡)=𝑋𝑘 1+𝑒−𝑟(𝑡−𝑡0) (19) The new reaction kinetics expression shows a good agreement with experimental data and previous works [158]. As may be seen in Figure 7a, a rather good agreement can be found between experiments and theory by only adjusting the pre-factor 𝑎2 in Eq. 18 as a free parameter in the theoretical curve. In view of these results, and even though technical limitations prevented us from carrying out carbonation experiments at pressures greater than atmospheric, Eq. 18 can be used to estimate the reaction rate under CO2 at pressurized conditions. Table 2 summarizes the values used for the reaction enthalpies, entropies, and activation energies that will be employed for the theoretical reaction rate in the modelling analysis ahead. Table 2: CO2 partial pressure at equilibrium, reaction rate and values of enthalpy-entropy changes in the chemical decomposition and desorption stages, and activation energies [183]. ∆𝐻𝑟0 180 kJ/mol 𝐸2 20 kJ/mol ∆𝐻10 160 kJ/mol ∆𝑆𝑟0 0.16 kJ/(mol·K) ∆𝐻𝑑0 20 kJ/mol ∆𝑆10 0.068 kJ/(mol·K) 𝐸𝑑 20 kJ/mol ∆𝑆𝑑0 0.092 kJ/(mol·K) 𝐸1 180 kJ/mol 𝑎2 1160 (1/s) 49 Figure 7: Reaction rate obtained from experimental tests (using different TG analyzers as indicated) and theoretically predicted (Eq. 10, with a2=1160 1/s) for carbonation under pure CO2 at atmospheric pressure as a function of the temperature [183]. Using the proposed model together with experimental thermogravimetric analysis results found in the literature [141,173,190], a carbonator model, which was previously employed to analyze the CaL process for CO2 capture [87], has been adapted to analyze the carbonation behavior after a long number of cycles in the industrial process for its integration in CSP plants. As discussed above in the kinetics study, CaO conversion for a single particle in a certain residence time is dependent on carbonation temperature. Moreover, CaO conversion after several cycles decays to a residual value, which is also dependent on process conditions. By means of a mass balance, the maximum average conversion of the CaO particles in the carbonator can be expressed according Eq. (3). Unlike in the CO2 capture application, after calcination at 725ºC, most of the carbonation in short residence times on the regenerated CaO skeleton occurs in the fast stage due to the high CO2 concentration and carbonation temperature [190]. Thus, carbonation in the diffusion-controlled phase can be neglected in this case and therefore 𝑋𝑟=𝑋𝑘 . Accordingly, the present carbonator model assumes that carbonation occurs at a given rate until it reaches the maximum carbonation allowed in the fast carbonation stage, after which the particles remain inactive. Thus, only a fraction of particles, 𝑓𝑎 , are active in the carbonator with the capacity to react in the fast stage (Eq. 8). Considering a perfect mixing model, the average conversion of the particles leaving the carbonator (𝑋) can be calculated using the followed equations [87]: 50 𝑋=∫𝑟𝑎𝑣𝑒 𝑡𝐾 0𝑡 (1𝜏)𝑒−𝑡𝜏 𝑑𝑡 1−𝑒−𝑡𝐾 𝜏=𝑋𝑚𝑎𝑥,𝑎𝑣𝑒 𝑓𝑎 ln(1 1−𝑓𝑎) (20) 𝑓𝑐𝑎𝑟𝑏=𝑋 𝑋𝑚𝑎𝑥,𝑎𝑣𝑒 (21) 𝐸𝐶𝑂2=𝐹𝑅 𝐹𝐶𝑂2𝑋=𝑁𝐶𝑎𝑓𝑎𝑟𝑎𝑣𝑒 𝐹𝐶𝑂2 (22) where 𝑓𝑐𝑎𝑟𝑏 is the average carbonation level in the carbonator and 𝑟𝑎𝑣𝑒 is the average reaction rate in the fast carbonation stage, which is calculated from the kinetics model theoretical prediction (Eq. 18). 3.2.2 Modelling results The carbonator model allows us to carry out a sensitivity analysis to assess the effect of pressure, temperature and solids inventory in the carbonator on the average carbonation level (𝑓𝑐𝑎𝑟𝑏). As can be seen in Figure 8, the average carbonation level (𝑓𝑐𝑎𝑟𝑏) is enhanced by increasing the carbonator pressure due to faster reaction kinetics. After a few of seconds in the carbonator, most of the CaO reaches the maximum conversion due to the very fast kinetics achieved in these CaL conditions. This is basically because of the high amount of CO2 entering the carbonator (which exceeds the stoichiometric amount in order to use the effluent excess as heat carrier). According to the kinetics model, by increasing the carbonator temperature, the average carbonation level is slightly enhanced up to a temperature T(r max) is reached from which kinetics is penalized with the consequent curtailment of the average carbonation level. As shown in Figure 8, this effect is mitigated when the carbonator pressure is increased, which is beneficial in practice since working at higher carbonator temperatures will promote the power cycle efficiency. These results were published in [183]. The full manuscript including complete information about the proposed carbonation model for TCES is presented in Annex 9. 51 Figure 8: Carbonator model results. Average carbonation level (fcarb) for several carbonation conditions (P,T). T(rmax) is the temperature at which the reaction rate reaches a maximum (Fig. 3) [183]. 3.3. CaL integration in Coal-Fired Power plants This section summarizes the main results of the work according to objectives (2-3) of the present thesis. A principal feature of the CaL process at CO2 capture conditions is that it produces a large amount of energy and therefore an optimized integration of the systems energy flows is essential 52 for the feasibility of the integration at the commercial level. The proposed integration model aims to analyze the process under realistic conditions for CO2 capture so the model proposed in Section 3.1 is used to estimate the carbonation efficiency. This involves considering the important role of the diffusion-controlled carbonation phase, which becomes relevant when CaO regeneration is carried out under high CO2 partial pressure as it is the case with the CaL process for CO2 capture. These results were published in [191]. 3.3.1 Integration scheme development A reference 505 MWe coal fired power plant (CFPP) is used in this work for analyzing the energy penalty arising from the CaL integration [51]. Combustion of 205 tonnes per hour of coal takes place in the steam boiler to generate 1,335.7 MWth with an 88.6% efficiency. For the steam cycle, a 42% thermal to electric steam turbine efficiency is assumed based on results reported for similar plants [77,192]. For the CaL-CFPP integration analysis, mass and energy balances in the process have been calculated by using the commercial software ASPEN PLUSTM. Figure 9 shows the schematic CFPP-CaL integration scheme. Figure 9: CFPP-CaL integration scheme [191]. The CaL process is started by calcination of the CaO precursor in the CFB calciner, where the heat necessary to increase the temperature of the solids to carry out the endothermic calcination reaction is provided by in-situ oxy-combustion of coal. For simulating the oxy-combustion of coal in the calciner reactor a model based on chemical and phase equilibrium through Gibbs’ free energy minimization methods is used. In the model, the energy penalty caused by The Air Unit Separation (ASU) is taken as 200 kWh per tonne of O2 produced according to results from previous reports [86,193] while the energy consumption derived from solids transportation has 53 been set at 20 MJ per tonne of solids [194]. A make-up flow of the CaO precursor is introduced into the calciner to compensate for sorbent deactivation. CO2 capture takes place in the CFB carbonator reactor. A solids stream (𝐹𝑅) composed by CaO enters the carbonator to react with the CO2 loaded flue gas stream (𝐹𝐶𝑂2=10,388 kmol/h) coming from CFPP main boiler. The operating carbonator temperature is set to 650°C and it works at atmospheric pressure. An inventory of solids in the carbonator bed (𝑊𝑠) of 200 tonnes of solids (400 kg/MWel) is considered as a typical value used in carbonator models [47,87,179]. Gas pressure loss across the carbonator reactor, which is operated in the continuous fluidized bed (CFB) regime, is calculated from the Kunii-Levenspiel (K-L) fluid dynamic model [195,196]. The heat produced in the CaL cycle is used in a secondary superheated steam cycle for electricity generation. Steam operational parameters have been chosen from data of similar real power plants [197]. The steam boiler is modelled as: i) a preheater where the sensible heat from the flue gas stream is used; ii) an evaporator, which takes advantage of the energy released in the exothermic carbonation reaction and iii) a steam super-heater, which uses the sensible heat of the CO2 exiting the calciner. A possible strategy to minimize the global energy penalty is to exchange heat between the solids leaving the calciner and the solids entering it that must be heated up to the calcination temperature. Different systems have been proposed to this end such as a mixing valve [198] or a cyclonic preheater [199]. In order to simplify the model, a simple solid–solid heat exchanger is assumed with an approach temperature of 20 ºC. At the carbonator exit, the flue gas (after recovering 70-90% of CO2) is sent to a series of heat exchangers in order to recover sensible heat from the exhaust gases stream before being vented to the atmosphere. 3.3.2 CFPP-CaL simulation results Parametric tests have been carried out in order to evaluate the CaL integration potential in a real power plant. Percentage of penalty points was calculated over the plant efficiency. As may be seen in Figure 10, the efficiency penalty is notably decreased as the solids residence time in the carbonator is increased. Moreover, the lower the flow rate of fresh limestone makeup the larger the reduction of efficiency penalty. Thus, it can be concluded that a lower flow of solids recirculation between the reactors would lead to a remarkable improvement of the integrated cycle efficiency. This is a relevant result to be highlighted that arises from the new carbonator model in which the importance of carbonation of the CaO solids in the diffusion controlled phase is considered. 54 Figure 10: Efficiency penalty for the integration of the CaL technology in a coal combustion power plant as a function of the solids residence time in the carbonator [191]. An interesting parameter to assess the tradeoff between the achieved CO2 capture efficiency and energy penalty is the specific energy consumption per kg of CO2 captured (SPECCA). Thus, SPECCA represents the additional fuel consumption (in MJ) necessary to avoid the emission of 1 kg of CO2 into the atmosphere [56,200,201]. Most of the studies assessing SPECCA for conventional amine based capture systems lead to typical values of about 4 MJ/kg CO2 [21,35,202]. Figure 11 shows the results of the values obtained for SPECCA in the proposed CFPP-CaL integration model to obtain a CO2 capture efficiency of 80% and 90%, respectively, as depending on the solids recirculation flow rate. As can be seen, for a given value of the solids recirculation flow rate (FR/F0 fixed), it is necessary to increase the amount of energy required to obtain a higher capture efficiency since the makeup flow rate of fresh limestone introduced into the system F0 must be increased. Furthermore, it is observed that a lower SPECCA is attained as the solids recirculation flow rate if decreased for a given capture efficiency. Thus, the additional energy consumption necessary to achieve a given capture efficiency is lowered as FR/F0 is diminished. 510152025 4% 5% 6% 7% 246810 FR /FCO2 Efficiency penalty (% points) t(min) F0=0,1 F0=0,05 F0=0,01 55 Figure 11: SPECCA value as a function of the ratio of solids recirculation flow rate to CO2 flow rate (FR/FCO2) to get a fixed CO2 capture efficiency (ECO2) value of 80% (left) and 90% (right) [191]. 3.3.3 Comparison between natural CaO precursors Previous sections are based on limestone as CaO precursor. According to objective (3) alternative Ca-based precursors have been studied. Although there is a large number of synthetic sorbents being currently analyzed within the CaL process, the present thesis is focused on the study of environmental-friendly, low cost and widely available precursors because of the importance of these factors for the CaL large scale deployment. Thus, this section summarizes results achieved by comparing limestone, dolomite and steel slag as CaO precursors over the CFPP-CaL scheme summarized in the previous section. These results were published in [203]. Dolomite (MgCa(CO3)2) is also a cheap and abundant natural CaO precursor, which upon calcination under the presence of CO2 is decomposed into MgO and CaO by a two stages process [48,204,205]. The use of dolomite in the CaL process for CO2 capture would allow reducing the calcination temperature to about 900ºC to attain full calcination of the makeup flow in short residence times [48]. Moreover, the presence of inert MgO grains in the dolomitic lime hinders aggregation and thus sintering of the CaO grains, which mitigates the drop of CO2 capture capacity with the number for cycles [48,206]. In the same line, the use of steel slag as alternative CaO precursor in the CaL process is also gaining the attention of researchers [207,208]. Steel slag is produced in large amounts by the metallurgical industry, remaining an important part as final waste without valorization. By a previous treatment with acetic acid, calcium acetate as CaO precursor is obtained [209]. Results from TGA tests show that decomposition of calcium acetate occurs mainly in three steps [209– 211]: i) an initial dehydration, which occurs from ambient temperature to approximately 250ºC; ii) calcium acetate decomposition from 300ºC to 450ºC to release acetone and CaCO3 and iii) calcination of CaCO3 from 620ºC to 700ºC. According to TGA results [209], the CaO-based 2.58 3.17 3.16 4.19 3.34 4.39 3.38 4.41 ECO2=80% ECO2=90% SPECCA (MJ/KG CO2) FR/FCO2=5 FR/FCO2=10 FR/FCO2=15 FR/FCO2=20 56 sorbent derived from steel slag presents a low deactivation rate at realistic CaL conditions, which would expectedly improve the CO2 capture efficiency of the process. A comparison of capture capacity (CC) in both fast and diffusive-controlled carbonation stages between limestone, dolomite and steel slag is shown in Figure 12. Figure 12: Total CO2 capture capacity, capture capacity in the fast reaction controlled phase (FRP) and solid-sate diffusion controlled phase (SDP) as a function of the cycle number for limestone, dolomite and steel slag [63] [48] [209]. Reproduced from [203]. As can be seen in Figure 12 in the case of limestone derived CaO, the contribution of the diffusion controlled phase to the overall CO2 capture capacity in each cycle is similar to the capture capacity in the fast reaction controlled phase (for 5 min overall carbonation periods). Carbonation in the SDP becomes even more relevant for dolomite and steel slag, which suggests that the capture performance may be further improved if the solids residence time in the carbonator is prolonged beyond a few minutes. SPECCA results plotted in Figure 13 show that increasing the solids residence time leads to a decrease of energy consumption for all the CaO precursors. Figure 13: SPECCA values as a function of solids residence time in the carbonator for different CaO precursors [203]. 57 As shown in Figure 13, under the simulated conditions, a SPECCA as low as 2.2 MJ/Kg CO2 is obtained for dolomite at carbonation residence times of 10 minutes (corresponding to a solids recirculation flow rate of 𝐹𝑅= 40,500 kmol/h) as compared to 2.6 MJ/Kg CO2 for limestone. These results were published in [191,203]. The full manuscripts including further information about the CFPP-CaL integration modelling are presented in Annexes 2 and 4. 3.4. The novel Oxy-CaL process for CO2 capture This section summarizes the main results of the work according to objective (4) of the present thesis. Thus, a novel CO2 capture system (Oxy-CaL) is proposed from the integration of partial oxy-combustion and the CaL process based on the multicycle carbonation/calcination of limestone derived CaO. Basically, Oxy-CaL consists of carrying out a partial oxy-combustion process to produce a flue gas with a CO2 concentration in the range 30-60% vol, which is then sent to the CaL capture process. The new process development is supported by experimental results from TGA analysis on the multicycle conversion of limestone derived CaO under realistic calcination conditions (high temperature and high CO2 concentration). These TGA results are used in the Oxy-CaL integration model to calculate the CO2 capture efficiency from process simulations. The energy penalty arising from the diverse CO2 capture technologies considered (total oxy-combustion, CaL and Oxy-CaL) is analyzed and compared with those of other CO2 capture systems. These results were published in [54]. 3.4.1 Thermogravimetric analysis TGA tests were carried out by Dra. Mónica Benítez and Dr. Antonio Perejón, recognized specialists in TGA techniques, within a fruitful collaboration among our working groups [54]. In these TGA tests, the CO2 concentration in the carbonation environment was varied in the range 15-60% vol to address the effect of an excess of CO2 in the carbonator over the typical vol% in the flue gas at typical combustion conditions (~15%). Moreover, the carbonation temperature was varied in the range 625-680ºC, which affects critically the carbonation kinetics in the solid-state diffusion-controlled stage as will be seen. Figure 14 shows the different sorbent behaviors in both configurations. As may be seen, carbonation in the fast phase is markedly enhanced as the CO2 concentration is increased, whereas, on the contrary, diffusion-controlled carbonation is markedly 64 (~20-40% v/v) is sufficient in the case of partial oxy-combustion, air separation can be carried out by means of membranes, which highly reduce the energy consumption associated with the oxy-combustion. In fact, membranes are economically suitable for O2 purities in the range 25- 50% [214]. Membranes are used in a huge amount of industrial applications among which we find air separation [215]. Polymeric membranes are today the most developed membranes for gas separation although other possibilities are gaining attention in the last years such as ion transport membranes. Polymeric membranes are generally non-porous and separate gases from the solution-diffusion transport mechanism by applying a transmembrane partial pressure difference. In a typical membrane-based gas separation process the feed gas is put into contact with the membrane surface under a certain pressure, which is the driving force of the process. The most permeable gas crosses the membrane and leaves the system as permeate stream (at lower pressure that the feed gas), whereas the remaining gas exits at high pressure as the retentate stream without crossing the membrane. Membranes for gas separation are usually analysed in terms of permeability and selectivity, which highly influence the efficiency and cost of a membrane-based gas separation process. As pointed out by [216], in polymeric membranes there is usually a tradeoff between permeability and selectivity. The O2/N2 separation factor for the best commercially available polymer membranes is 6–8 [217]. In the same work, data can be found for permeability and selectivity of many polymeric membranes. The pressure ratio (PR) between the upstream and downstream membrane sides is a key operating parameter in the separation process. In order to reach a O2 purity 40% it would be enough a pressure ratio across the membrane of 0.1 by using a single-stage process for a membrane selectivity 3 [214]. Nevertheless, it is remarkable that the higher pressure ratio the higher permeability of polymeric membranes [216]. Thus, in order to increase the permeability and therefore reduce the membrane area needed, a higher PR must be imposed. The energy consumption associated to a separation process can be calculated from the adiabatic compression work of a perfect gas, which gives a realistic estimation of energy consumption [218]. The pressure ratio can be imposed either by increasing pressure on the upstream side (i.e. compressing the gas entering the membrane system) or by creating vacuum on the downstream side, which is the selected technique in the present work. Since a vacuum strategy is considered, the pressure must be increased again in the permeate (OEA) before entering the boiler. Thus, energy consumption (𝐸𝑠𝑒𝑝) can be calculated as [219]: 65 𝐸𝑠𝑒𝑝=𝑚𝑝𝛾𝑅𝑇 𝛾−1[((𝑃𝑖𝑛 𝑃𝑝)𝛾−1 𝛾−1)+((𝑃𝑏 𝑃𝑝)𝛾−1 𝛾−1)] (23) where 𝑚𝑝 is the mass of gas permeating though the membrane, 𝛾 is the adiabatic expansion factor of the gas, 𝑇 is the temperature of the system and 𝑃𝑖𝑛, 𝑃𝑝 and 𝑃𝑏 are the pressures at the upstream side, downstream side and boiler inlet, respectively. Membrane separation is simulated within the whole mOxy-CaL process from the energy consumption as given by Eq (1). Thus, by considering a gas pressure at the permeate side of 0.235 bar (similarly to previous works such as [220]), the energy requirement for separation is 90 kwh/tonne O2. A stage cut, defined as the ratio of permeate flowrate to feed flowrate of 0.1 is assumed [221]. The mOxy-CaL process has been modelled by means of ASPEN PLUSTM with the same assumptions than in the Oxy-CaL simulation case presented in the previous section. Table 4 shows the main results. Table 4: Main inputs and results for the mOxy-CaL30 process Coal-Fired Power Plant CaL Parameter Value Parameter Value 𝑚󰇗𝑐𝑜𝑎𝑙 (𝑘𝑔/𝑠) 46.1 𝐹𝑅/𝐹𝐶𝑂2 15 𝑚󰇗𝑎𝑖𝑟 (𝑘𝑔/𝑠) 0 𝐹0/𝐹𝐶𝑂2 0.05 𝑚󰇗𝑂2(𝑥𝑂2=0.95) (𝑘𝑔/𝑠) 0 𝜏 (𝑚𝑖𝑛) 2.79 𝑚󰇗𝑂𝐸𝐴(𝑥𝑂2=0.38) (𝑘𝑔/𝑠) 281.02 𝑇𝑐𝑎𝑙𝑐 (º𝐶) 950 𝛾𝑓𝑔 0.63 𝑇𝑐𝑎𝑟𝑏 (º𝐶) 650 𝐹𝑓𝑔 (𝑘𝑚𝑜𝑙/𝑠) 10.25 𝐸𝐶𝑂2 0.950 𝐹𝐶𝑂2 (𝑘𝑚𝑜𝑙/𝑠) 2.84 𝑊󰇗𝑠𝑒𝑐 (𝑀𝑊) 118.14 𝑣𝐶𝑂2 0.3 𝑊󰇗𝑠𝑜𝑙𝑖𝑑𝑠 (𝑀𝑊) 49.73 𝑣𝑂2 0.025 𝑊󰇗𝐴𝑆𝑈 (𝑀𝑊) 42.13 𝜂𝑏𝑜𝑖𝑙𝑒𝑟 0.89 𝑊󰇗𝑐𝑜𝑚𝑝,𝐶𝑂2 (𝑀𝑊) 70.94 𝜂𝐶𝐹𝑃𝑃 0.362 𝑊󰇗𝑐𝑜𝑚𝑝,𝐹𝐺 (𝑀𝑊) 7.37 𝑛𝑒𝑡 𝑤𝑜𝑟𝑘 (𝑀𝑊) 513.59 𝜂𝑖𝑛𝑡 0.301 𝑃𝑒𝑛𝑎𝑙𝑡𝑦 1.63% 𝑃𝑒𝑛𝑎𝑙𝑡𝑦 6.12% 𝑆𝑃𝐸𝐶𝐶𝐴 (MJ/kg CO2) 0.55 𝑆𝑃𝐸𝐶𝐶𝐴 (MJ/kg CO2) 2.49 𝑃𝑒𝑛𝑎𝑙𝑡𝑦𝑡𝑜𝑡𝑎𝑙 7.75% 𝑆𝑃𝐸𝐶𝐶𝐴𝑡𝑜𝑡𝑎𝑙 (MJ/kg CO2) 3.05 As can be seen in Table 4, a notable reduction of energy consumption for CO2 capture is achieved by means of the mOxy-CaL (3.05 MJ/kg CO2) process in comparison with the Oxy-CaL system (3.56 MJ/kg CO2 achieved in the previous section). This energy reduction is directly attributable to the membrane separation, which saves power to produce O2 enriched air as compared with cryogenic distillation. 66 3.5. Calcium-Looping integration in concentrating power plants This section summarizes the main results of the thesis according to objectives (5,7,9). Concentrated Solar Power (CSP) should play a leading role within the new energy landscape as it lends itself to potentially cheap storage of energy in the form of heat [9,95]. Thus, efficient and affordable thermal energy storage systems must be developed in order to decouple production and demand [97,98], which would allow a deep penetration of solar energy power generation into the grid. Among the possibilities explored for TCES in CSP tower plants at large scale, one of the most promising systems is the CaL process. These results were published in [145,222]. 3.5.1 CSP-CaL plant description Process integration is fundamental for the introduction of the CaL process within the solar power plant. Despite that the CaL process was considered as solar energy storage systems in the late 1970s [3,128] and solar calcination has been tested since the 80s [4,5], process integration schemes have not been proposed until a few years ago. In our scheme, solar energy is used to carry out the calcination at high temperature to achieve sufficiently fast calcination under a pure CO2 atmosphere. According to equilibrium conditions [161] and reaction kinetics, high temperatures are necessary when operating under high CO2 partial pressure (above 900°C) for sufficiently fast calcination to achieve completion in short residence times [10,11,223]. A number of solar calciner prototypes have been already developed based on fluidized beds [5,147], rotary kilns [224,225] and cyclone atmospheric reactors [226]. A potentially cheap and simple option consists of falling particle receivers [227]. Further assessment of solar calciners is shown in Section 3.8.1. Since the CSP-CaL integration scheme proposed in this thesis relies on a closed CO2 cycle, a pure CO2 stream must be retrieved as output from the calciner. Two options are feasible to this end: i) performing calcination under a 100% CO2 atmosphere or ii) under an easily separable gas from CO2. The first choice would make it necessary to operate the calciner under low absolute pressure in order to reduce the calcination temperature and to avoid hampering the reactivity of the regenerated CaO due to excessive sintering as observed when calcination is carried out under high CO2 partial pressure for CO2 capture [88]. Among CSP power technologies, solar tower systems are the most appropriate for CaCO3 calcination because of the required temperature. Thus, the total solar energy entering into the system is used to heat the reactants and to drive the endothermic chemical reaction, which serves to store the solar energy as both sensible and chemical and form. The CaO and CO2 streams released within the calcination are stored separately. Products can be stored at ambient 67 temperature, which is an important advantage of the CSP-CaL integration over current state of the art sensible heat storage using molten salts, whose temperature must be kept above ~200°C to avoid solidification [123]. In order to use reasonably sized CO2 storage volumes a minimum pressure of 75 bar is needed to store CO2 storage under supercritical conditions (considering storage at ambient temperature) [222]. The high compression ratio from calciner to storage conditions (1:75) requires the use of intercooling compression to minimize the efficiency penalty. Note that compression-expansion process of CO2 before and after than storage resembles a Compressed Air Energy Storage (CAES) system [228] and therefore the process incorporates energy storage not just in chemical form but also as sensible heat and mechanical energy. When power is needed, both calcination products are brought together to another reactor (carbonator) where energy is released through carbonation. Pressurized carbonation is desirable for the powercycle direct integration and allows carbonation at high temperatures and at a fast rate [229]. The energy discharge phase takes place in the carbonator, which is a pressurized fluidized bed reactor wherein carbonation takes place under pure CO2 at the highest possible temperature in order to achieve a high thermal to electric efficiency. According to thermodynamic equilibrium, carbonation can be carried out at high temperature (>850ºC) under a pure CO2 atmosphere at ambient pressure [161]. This would allow a highly efficient generation of electricity thus overcoming temperature limits (T~550-600°C) in currently commercial CSP plants with thermal storage in molten salts. Solids exiting the carbonator are passed through another heat exchanger network to preheat the CaO and CO2 streams entering the carbonator. In order to achieve a higher integration efficiency an optimized heat exchanger network is necessary. Solids conveying can be carried out by means of the mature pneumatic technology that guarantees rapid transport of granular solids, routing flexibility, as well as the possibility of streams splitting, and it is suitable to successfully convey high temperature materials. Other possibilities for solids conveying are discussed in Section 3.8.3. Plant performance is determined as an average over a 24 hours period. An average daytime period ∆𝑡𝑠𝑢𝑛 is considered during which the solar thermal power entering the calciner is constant. Different operations in “sun” and “night” modes were considered. A Solar Multiple (SM) equal to 3 was assumed, which means that, when solar energy is available, the CO2 mass flow entering the carbonator side is 1/3 the amount produced in the calciner whereas the 2/3 remaining are sent to a CO2 storage vessel which is discharged during the “night” mode operation (estimated as 2/3 of the day). Therefore, the plant efficiency is determined as a weighted average of the performances in “sun” and “night” modes by means of Eq. 24: 68 𝜂=∫𝑊󰇗𝑛𝑒𝑡 𝑑𝑡 24ℎ ∫𝑄󰇗𝑖𝑛𝑝𝑢𝑡 𝑑𝑡 24ℎ=𝑊󰇗𝑛𝑒𝑡,𝑠𝑢𝑛 ∆𝑡𝑠𝑢𝑛+𝑊󰇗𝑛𝑒𝑡,𝑛𝑖𝑔ℎ𝑡 (24−∆𝑡𝑠𝑢𝑛) 𝑄󰇗𝑖𝑛𝑝𝑢𝑡 ∆𝑡𝑠𝑢𝑛 (24) Where 𝑊󰇗𝑛𝑒𝑡,𝑠𝑢𝑛 and 𝑊󰇗𝑛𝑒𝑡,𝑛𝑖𝑔ℎ𝑡 are the net power produced in “sun” and “night” modes, respectively and 𝑄󰇗𝑖𝑛𝑝𝑢𝑡 is the net solar power entering the calciner. 3.5.2 Process scheme: base case Regarding the integration of the power cycle in the carbonator zone, previous works have proposed the use of an air stream as heat transfer fluid in an open Brayton cycle (Figure 19) [194]. According to this scheme, the CO2 stream entering the carbonator is wrongly assumed to react completely with the CaO solids to produce CaCO3. Thus, it is assumed that a pure air stream can exit the carbonator to enter the gas turbine for power production in an open Brayton cycle. The outflowing air from the turbine passes through a heat exchanger network, releasing sensible heat further used to preheat the solids directed into the carbonator. Figure 19: Air/CO2 open cycle integration in the carbonator zone. Originally proposed by Edwards et al. [194]. Reproduced from [145] However, reaction equilibrium poses a fundamental limitation to this scheme since the reaction will reach equilibrium and carbonation will stop as soon as the CO2 partial pressure in the carbonator reactor is decreased to the equilibrium partial pressure as depending on the carbonator temperature. Thus, the effluent gas from the carbonator to be sent to exhaust cannot be free of CO2. For example, when a 50% v/v CO2 stream is introduced into a carbonator at atmospheric pressure (P=1 bar) and T=800ºC, the minimum CO2 concentration in the gas stream exiting the carbonator is around 21% [161]. 69 A possible solution to avoid the inconveniences of an open Brayton cycle is to use a closed CO2 Brayton cycle [222]. In this configuration, solids in the carbonator (CaO) are fluidized by a pure CO2 gas flow with a molar rate well above the stoichiometric value. The CO2 fraction not participating in the reaction is employed to remove heat from the carbonation and is delivered to a gas turbine for the power cycle. This section shows the global integration model and process design to transform an air-open power cycle into a closed-CO2 power cycle as a first step for improving the CSP-CaL integration. Figure 20 shows a first possible approach for the proposed integration. As can be seen, solids entering the calciner are preheated using the sensible heat released by the hot streams leaving this reactor in a gas-solid heat exchanger (HXB in Figure 20) and in a solid-solid heat exchanger (HXA). The CO2 power cycle is a closed and regenerative cycle in which the heat removed by the reactants in the carbonator is recovered in a solid-gas heat exchanger (HXF). The additional CO2 storage vessel (CO2 inventory storage) is included in the scheme as an inventory control strategy for the gas turbine. On the other hand, the residual heat from the solids at the carbonator output is extracted to pre-heat the CO2 stream entering the carbonator by means of another gas-solid heat exchanger (HXE). As shown in Figure 20, part of the power needed in the compression stage of the Brayton cycle is provided by the expansion of the pressurized CO2 coming from the storage vessel [222]. CaO CO2 CaCO3 CO2 CaCO3 CaO CO2 Turbine C T Solids (CaCO3 + inactive CaO) CaO CO2 HXA HXB HXE HXF Cooler CO2 storage pCO2 storage CO2 inventory storage C T Inventory Pressurized Fluidized Bed CARBONATOR Tcarb pcarb Solar receiver - CALCINER CaO storage CaCO3 (+inactive CaO) storage s7 s3 c1 s5 s4 s6 s8 c2g1 g2 g3 s1s2 c3 c4 g4 g5 g6 g7 g8 g9 g10 g11 g12 Figure 20: Base case for the CO2 closed Brayton cycle integration in the CSP-CaL plant. Further information about the stream data can be found in [145]. The CSP-CaL base case was simulated by using the main parameters showed in Table 5. Further information about the simulation is available in Annex 5. 70 Table 5: Main parameters used for simulating the scheme proposed in Figure 20 [145] Net absorbed solar flux in calciner 100 MW Approach temperature CO2 cooler 10 °C Thermal dispersions in carbonator 10% Intercoolings in CO2 storage compression 5 Calciner temperature 900 ºC Intercoolings in CO2 cycle compression 4 Ambient temperature 20 ºC CO2 storage temperature 25 ºC Average CaO conversion 0.20 CO2 storage pressure 75 bar Carbonator temperature (T) 875 ºC Solid phase conveying energy consumption 20 MJ/ton Carbonator pressure (P) 6 bar Daylight hours 12h Turbine outlet pressure (TIT) 0.2 bar Solar Multiple 2 Approach temperature solidsolid HX 20 °C Isentropic efficiencies (compression/expansion) 0.89 Approach temperature solidgas HX 15 °C Figure 21 shows a contour plot of the system efficiency as a function of both carbonator pressure and turbine outlet pressure, which serves to infer the optimum pressure choice. Figure 21: Efficiency as a function of carbonator and turbine outlet pressure. Dashed white lines indicate iso-efficiency cycle values [145]. As shown in Figure 21, maximum efficiency occurs at pressure ratios around 40-50. However, it is important to note that for pressure ratios over 30 the efficiency does not change considerably. From pressure ratios of 30 up to 70, efficiency increases less than 0.5%. Since such high values of pressure ratios are difficult to achieve in practice, a pressure ratio of 30 represents a good tradeoff. Moreover, higher efficiency can be achieved by decreasing the turbine outlet pressure. 71 3.5.3 Improving the CSP-CaL heat integration From the base case (Figure 20), a CO2 regenerator (HXG) is incorporated in the heat exchanger train in order to heat up a fraction of the gas stream entering the HXE exchanger (and then in the carbonator) using the sensible heat of a portion of the CO2 turbine output flow. The two CO2 streams, which are separately heated through heat exchangers HXF and HXG, are rejoined to evolve in the closed loop power-cycle. By regulating the CO2 split ratios (i.e. the fractions of the CO2 stream respectively sent to HXG and HXF), the internal overall heat exchange can be optimized by means of a pinch-analysis [230] of the carbonator side. The pinch-analysis were mainly carried out by Alessandro Alovisio and Professor Ricardo Chacartegui [145,222] with the collaboration of the author of the present thesis. Full information about the pinch analysis developed is available in annex 5. As result of the pinch-analysis, the energy-optimized process scheme is shown in Figure 22. Remarkably, this energy-optimized scheme is realized by considering low-temperature energy storage conditions. Temperature in the storage vessels highly conditions the configuration of the process scheme and the efficiency of it. Due to the high temperatures in both calciner and carbonator reactors, a low temperature storage involves a large streams temperature change along the entire cycle which makes crucial an optimized heat integration to achieve an adequate system efficiency. Another possibility, which is investigated throughout the Section 3.7, is to storage the solids at high temperature. In this case the process scheme is simplified by requiring fewer heat exchangers due to a lower temperature difference between the reactors and the storage. CaO CO2 CaCO3 CO2 CaO HXA HXB HXF CO2 storage pCO2 storage CO2 inventory storage C T Inventory Tcarb pcarb Solar receiver - CALCINER CaO storage CaCO3 (+inactive CaO) storage CO2 Turbine C T HXE Cooler HXG Pressurized Fluidized Bed CARBONATOR HXI CaCO3 (+ inactive CaO) CaO CO2 s1 s3 s2 s4 s5 s6 s7 s8 s9 c1 c2 c3 c4 g1 g2 g3 g4 g5 g6 g7 g8 g9 g10 g11 g12 g13 g14 g15 g16 g17 g18 g19 Figure 22: Plant diagram of the energy-optimized integration layout for the CSP-CaL plant with low-temperature storage. Full information about the stream data is given in [145]. 72 According to the pinch analysis results, the proposed final plant configuration is equipped with a solid-solid heat exchanger (HXA), four gas-solid heat exchangers (HXB, HXF, HXE, HXI) and with a gas-gas regenerator (HXG). The CO2 stream from storage (produced in the calciner side operation) and the CO2 stream coming from the power loop are mixed, flowing through a heat exchangers train (HXG and HXI) which optimize heat recovery at low temperature. On the other hand, the CO2 stream flowing out from the turbine is divided into two sub-streams through HXF and HXG to preheat the CaO stream and a fraction of CO2 entering the carbonator, respectively. In HXE, the high-temperature sensible heat from the CaCO3 stream is used in the final stage of CO2 stream preheating above the pinch, which serves to maximize the gas temperature at the carbonator inlet and therefore the cycle performance. Table 6 shows a comparison of the main data according to an energy balance for each configuration. At a fixed CaO conversion of 0.2, the global net efficiency increases of about 5% with respect to the base case. Table 6: Energy balance of for the two CSP-CaL configurations [145]. Parameter Layout 1 (Figure 20) Layout 2 (Figure 22) Solar thermal power (MWth) - 12h- 100 100 Heat exchangers Thermal Power (MWth) HXA 120.0 120.0 HXB 23.2 23.2 HXF 26.8 46.0 HXG - 78.6 HXI - 53.7 HXE 70.4 15.5 Auxiliary cooler 3.9 6.9 Power outlet (MWe) compressor calciner (storage) 7.0 7.0 compressor carbonator (power cycle) 16.7 14.0 Solids conveying (average) 1.8 1.8 Power inlet (MWe) turbine (storage) 1.2 2.2 main turbine (power cycle) 38.6 37.3 Global net efficiency 0.356 0.403 The high overall efficiencies achieved indicate a potential interest for the integration of the CaL process in Concentrating Solar Power Plants, although major technological challenges related to the design of the solar receiver and of the high temperature solid handling devices will need to be faced. 73 These results were published in [145,222]. The full manuscripts including further information about the CSP-CaL integration modelling are presented in Annex 3 and 5. 3.6. Indirect integration of power cycles This section summarizes the main results of the work according to objective (8) of the present thesis. Previous results are based on the integration of a CO2 closed Brayton cycle with direct integration in which the heat transfer fluid used in the carbonator is sent directly into a gas turbine. Main results of the analyses carried out for several power cycle configurations with the main goal of increasing the performance of the overall system integration are summarized below. Possible integration schemes are proposed in which power production is carried out indirectly, by means of a steam reheat Rankine cycle or a supercritical carbon dioxide Brayton cycle. These results were published in [231]. First, the indirect integration of a reheat Rankine cycle is considered, which is the typical power cycle in commercial CSP tower plants [165]. In this configuration, the preheat water of the Rankine cycle is heated by the exhaust CO2 stream from the carbonator in a heat recovery steam generator (HRSG) until super-heated status is reached. Maximum efficiency (around 35.5%) was obtained at 875ºC operating under atmospheric pressure. From a sensitivity analysis it was shown that the global integration efficiency is enhanced up to 37.5% by increasing the live steam and the reheat temperature. An indirect integration of a recompression supercritical CO2 (sCO2) Brayton cycle was also proposed, which led to a global efficiency close to 32% that was limited by the large amount of non-used energy linked to the cooling process before the compression stage. The analysis showed that hot and cold sides in the regenerator are strongly unbalanced, what allows predicting a certain improvement after optimization. Moreover, the low global efficiency value achieved is a consequence of the moderated turbomachinery efficiency values assumed since this technology is yet under development. The integration of a Combined Cycle (CC) in the CSP-CaL process was also analyzed in [231]. The CC is considered by a hybrid direct-indirect power cycle integration with the CSP-CaL system (Figure 23). The CO2 stream exiting the carbonator is expanded in a gas turbine as a previous step for transferring heat to steam cycle through a HRSG. Results showed that the highest performance is obtained by decreasing the outlet turbine pressure up to reach a maximum value of 40.4% for operation under an inlet/outlet turbine pressure ratio of 3.6/1. 80 development. In recent years, research on solar particle receivers has been notably intensified to increase both efficiency and dispatchability in CSP plants. Heat transfer losses is a critical issue for the solar receiver design, especially when high temperatures in the receiver are considered as would be the case of solar calcination under CO2. Radiation losses are mainly dependent on the receiver temperature although they may be mitigated by a proper design [233]. Conduction losses can be reduced by improving thermal insulation. Thus, increasing the solar absorptance, decreasing the thermal emittance and/or reducing conduction/convective heat losses can yield an important enhancement of thermal efficiency in the receiver [125]. In general, solar particle receivers can be classified by the way particles are irradiated by solar energy, either in a direct or indirect configuration, using in the second case tubes or other receptors to convey and heat the particles. Selection of the solar particle receiver to carry out the calcination of CaCO3 particles is dependent on a multitude of factors, mainly related to how the reaction is produced (kinetics, temperature, atmosphere composition, particle size) but also dependent on solar receiver characteristics (thermal emittance, solar absorptance, etc.) and physical properties of CaCO3 and CaO particles entering the solar calciner relevant to solar. Thus, CaCO3 has a poor solar absorptance [4] and therefore indirect heating in tubes may be more suitable as absorptance is not a determinant factor in this case [125]. Important factors that must be faced for the development of solar particle receivers are: i) enough residence time of the particles to reach the target temperature and allow for reaction completion; ii) adequate particle size; iii) proper management of elevated gas flow-rates used for particles transport (specially in fluidized beds); iv) possible deposition of particles in the optical window causing damage; v) continue operation preferable for scaling-up; vi) minimization of high thermal gradients and non-homogeneous conversion; and vi) particles attrition and agglomeration. Particles residence time and mass flow are crucial parameters when designing the receiver for solar calcination. In the CSP-CaL integration, solar radiation is used to heat the particles up to calcination temperature. The residence time to achieve full calcination must be long enough for full calcination to be achieved. The temperature at which the particles enter the solar receiver is dependent on the CSP-CaL scheme. Since carbonation does not occur completely, typically a mixture of CaCO3 and CaO solids enters the solar calciner. 3.8.1.1. Falling particle receivers Falling particle receivers are based on direct heating of solid particles that are falling while irradiated by a beam of concentrated solar energy flux. According to [125], the most interesting features of falling particle receivers are: i) unlimited maximum solar irradiance (>2000 kW/m2); ii) maximum temperature >1000ºC; iii) thermal efficiency between 50-90% and iv) estimated cost 81 of about 125$/kWt, which could be lower than molten salt receivers [125,234]. On the other hand, particle attrition is a critical issue that can be promoted in falling particle receivers. Figure 27 shows a conceptual scheme of the falling particle receiver system where a curtain of particles (produced by a slot at the base of a hopper) is being irradiated within a cavity receiver. This technology seems to be scalable to 10-100 MWe power tower systems [125]. Figure 27: Falling particle receiver system. Reproduced from [235]. A number of previous works have analyzed the main characteristics of falling particle solar receivers [125]. Siegel et al. developed a 2MWt on-sun test of a falling particle receiver, which is useful for validating computational models. The performance model indicated that an 80 W/cm2 average flux will be required to achieve 900°C given an inlet temperature of 600°C [236]. When the particle mass flow is increased the output temperature is decreased albeit thermal efficiency is improved. A possibility to increase the particles residence time is to recirculate the hot particles into the curtain [235,237]. Another possible strategy to raise the particles outlet temperature is to hinder the flow of falling particles by using chevron-like meshes. Ho et al. [238] studied the performance of a 1 MWth high-temperature falling particle receiver in an on-sun test with continuous recirculation. Ceramic particles were used in the test were and peak particle temperatures reached over 700 °C. They observed significant differences between the temperature in the center of the receiver and along the sides due to a non-uniform irradiance distribution. At an average irradiance of 211 kW/m2, the particles temperature increase was 57.1 °C per meter of drop length, and the thermal efficiency was ~65% [238]. 82 3.8.1.2. Centrifugal particle receivers The rotary kiln is a well-known technology for calcination in the cement industry. Since 1895 it has been the central part of all modern clinker producing installations [1]. The rotary kiln consists of a steel tube with a length to diameter ratio of in the range 10-38:1. Some characteristics of conventional rotary kilns are a maximum working temperature up to 2000ºC and scalability (> 200m long) able to produce around 3600 tonnes/day [1]. Usually, rotary kilns are coupled with particles preheaters where the gas effluent from the calciner at 1000ºC preheats the solids entering the kiln up to 700-800ºC. This would be interesting for the CSP-CaL integration. Nevertheless, for its use as solar receiver, the rotary kiln must be properly adapted to be heated at hightemperature by concentrated solar irradiation. Main advantages of rotary kilns for solar calcination are [149]: i) adjustable residence time of particles by controlling the tube rotational speed and tilting angle; ii) high temperatures achievable in the cavity; iii) solids conversion in continuous mode and in a controlled atmosphere; iv) the rotating tube enables uniform radial heating of the tube; vi) good heat transfer thanks to the direct contact of the rotating bed of particles with the hot tube walls; vii) possible operation in co-current or countercurrent flow, allowing solids preheating by the exhaust gas. A calciner based on a solar rotary kiln technology was proposed in 1980 by Flamant, Badie et al. [4,5]. The reactor is almost horizontal, with an inclination of 5º to facilitate particles transport. Both works highlighted the relevance of the radiative properties of the solar reactor. A thermal efficiency of 0.1-0.3 was achieved with a maximum calcination degree of 0.6 by using a 2 kWt solar furnace reactor with a total absorptance of about 0.9-1 due to the cavity effect. Excellent transfer coefficients between the bed and the internal wall were reported. Meier et al. developed two 10 kWth solar reactors to study calcination, one directly irradiated [224] and one indirectly irradiated [225]. For directly and indirectly heated solar reactors, maximum conversion efficiencies of 20% and 38%, respectively, were achieved for solar flux inputs of 2000 kW/m2. The typical reactor efficiency achieved from several tests reached 13% while the maximum efficiency reached 20% [224]. Main reactor energy losses were due to re-radiation (16.5%), conduction (9.5%), and sensible heat of products (10%), whereas the remaining losses (50.5%) were mainly due to convection heat losses and un-recovered calcined powder. It is expected that a further improved reactor design incorporating a recovering thermal energy system, and reduced heat losses by convection and conduction would lead to efficiencies up to 83% [224]. The indirectly heated reactor proposed by Meier at al. [225] consists of a tilted cylindrical steel drum with ceramic insulation (Figure 28). The indirectly heated reactor design seems especially interesting for the CSP-CaL integration since CO2 losses are avoided. In this case, an adequate gas conveying system is needed to ensure a proper gas circulation along the receiver. Typical 83 reactor overall efficiency varies between 25.2 and 34.8% depending on the rotational speed (8– 18 rpm), mass flow rate (2–8 kg/h), and particle size (2–3 mm) [225]. Figure 28: Schematics of the indirect heating multitube rotary kiln. Reproduced from [225]. Main reactor energy losses are caused by re-radiation (13.6%), conduction (20.5%), and sensible heat of products (20.2%), whereas the remaining losses (14.2%) were mainly due to loss of material and convection heat losses. In this case, an improved reactor design incorporating a recovering thermal energy system, and a reduced heat loss by convection and conduction would raise efficiencies up to 70%. 3.8.1.3. Fluidized bed receiver and comparison Fluidized Bed Reactors (FBRs) present higher thermal inertia than rotary kilns. Therefore continuous operation is improved in the former as the thermal regime can be maintained in the case of short fluctuations of solar radiation [4]. Both rotary kilns and FBRs have the potential for rapid deployment since the technologies on which they are based are already mature in the cement industry and FB-boiler manufacturing, respectively. Flamant, Badie et al. [4,5] proposed a FBR to carry out the CaCO3 decomposition. The total absorptance of the fluidized bed depends both on particle’s emittance and bed characteristics (porosity, particles size, fluidization regime) [4]. Salatino et al. [239] analysed the thermal performance of a bubbling fluidized bed of sand particles to store energy from CSP in sensible form. To evaluate the hydrodynamics of the process a near-2D fluidized bed prototype was constructed [240]. The use of compartmented dense gas fluidized beds was proposed by Chirone et al. [241] (Figure 29). The compartments are: i) the receiver, located around the cavity through which concentrating solar rays enter by means of a Beam-Down technology; ii) the heat exchanger section, where the sensible heat of solids is released to a steam cycle; and iii) storage. Dynamic simulation results showed an overall thermal efficiency (defined as the ratio of the 84 power transferred to the steam cycle to the incident radiative power) up to 71%. This study was the base for the construction of a 100kWt solar FBR prototype within the SOLTESS project [241]. Figure 29: Conceptual scheme of a compartmented dense gas fluidized beds. Reproduced from [241]. Since temperatures achievable in this pilot-scale plant were around 900ºC, it could be suitable to attain calcination of limestone, which could be used instead of sand thus storing energy both in sensible and thermochemical form as proposed in [242]. Finally, a comparison between the main receiver technologies applicable to CSP-CaL integration is shown in Table 7. Table 7: Receiver technology comparison. Particle receiver technology Falling particle receiver Rotary kiln FBR CSP-CaL integration issues Benefits -High solar irradiance - High thermal efficiency - Highly tested in the last years (MWth scale) -Tmax >1000ºC - Good scalability - Geometry similar than solar tower - Improved continuous operation mode - co-current or countercurrent flow -Core technology for calcination in cement industry -Tmax >1000ºC -Well-known preheaters coupling -adjustable residence time of particles - Improved continuous operation mode -good heat transfers - co-current or countercurrent flow - high thermal inertia improving the thermal regimen under short variations of solar radiation - Well-known technology at industry - high transfer coefficients - Direct and indirect heating configuration Challenges - Window integration -Increasing the residence time (especially problematic for particles >200um) - Particle attrition - Scalability must be addressed because the focalized solar heating - Window integration -Thermal losses in rotary kiln mainly occurs by conduction improve thermal insulation -total absorptance of the fluidized bed depends both of particle’s emittance and bed configuration - Beam-down technology development - Energy consumption for fluidization -Gas-solid separation Restrictions - Low solar absorptance of CaCO3 particles for direct irradiation - Direct heating is not possible because avoiding CO2 losses - Horizontal design which makes difficult is integration in solar tower at large scale. - Geometry incompatibilities with solar towers Tested for calcination No [4,5] [224] [225] [243] [4,5] 85 3.8.2 Heat exchangers An optimized heat integration is crucial for achieving high efficiencies in the CSP-CaL scheme due to the high temperature difference between the calciner, storage vessels and carbonator. Because of the system characteristics, a gas-gas heat exchanger as regenerator would be needed by considering a CO2 close-loop for power production [145]. Thus, CO2 exiting the turbine is used to preheat the CO2 stream entering the carbonator. The higher temperature of CO2 arriving at the carbonator the higher amount of carbonation energy used for power production, with the consequent increase in plant efficiency. On the other hand, the CO2 exiting the calciner at high temperature (>900ºC) must be cooled previously to be compressed and stored if storage at ambient temperature is the option chosen. To take advantage of the hot CO2 stream a gas-gas heat exchanger could be used as heat recovery steam generator (HRGS) within a secondary steam cycle, which could produce enough power to carry out CO2 compression [153]. Another possibility is to use solid-gas heat exchangers to release energy from the hot CO2 stream to preheat the CaCO3 particles entering the calciner [222]. Similarly, solid-gas preheaters allow increasing the CO2 temperature entering the carbonator -from the storage in which CO2 is at low temperatureusing heat from the CaCO3 particles exiting the fluidized bed reactor after carbonation. Finally, solid-solid heat exchangers would further improve heat integration. Gas–solid heat exchange can be carried out in either the open or closed configuration. Direct contact within an open configuration is a well-known technology [1]. Solids heating could be performed in a suspension preheater where gas and solids enter into contact sequentially in risers and are separated by cyclones, as commonly used in cement plants for raw material preheating [167]. In suspension preheaters raw particles are maintained in suspension by the hot gas from the calciner. Figure 30 shows a schematic of the suspension preheater. Suspension preheaters usually have between four and six cyclone stages, which are arranged one above the other in a 50 –120 m height tower [1]. The exhaust gases from the calciner flow through the cyclone stages from bottom to top. Particles are added to the exhaust gas before the uppermost cyclone stage. They are separated from the gas in the cyclones and re-join it before the next cyclone stage [1]. By using this solid-gas heat exchanger, a same exit temperature of the gas and solids streams are achieved from a co-flow arrangement with a considerably high contact surface. Due to high level of maturity of this technology, it seems an optimum option to carry out both the CaCO3 particles preheating in the calciner side and particles cooling in the carbonator side. 86 Figure 30: Suspension preheater; b) Grate preheater. Reproduced from [87,113] The grate preheater is another well-known technique in the cement industry (Figure 30: -b). Particles are fed onto a horizontal grate which travels through a closed tunnel. A fan draws the exhaust gas from the calciner into the top of the preheater. The temperature of the gas could be reduced from >1000ºC to 150ºC while solids would be heated up to 700-800ºC [1]. Another possible solid-gas direct heat exchanger was proposed by Shimizu et al. [244]. This heat exchanger is an axial flow cyclone in which a quasi-counter-flow heat exchange is realized to achieve an effective recovery of the heat carried by particles. By using this design an outlet gas temperature much higher than the particle outlet temperature could be attained. Fluidized-bed designs are also a possibility to perform the solid-gas heating with high heat transfer coefficients (up to ~600 W/m2-K) but with higher parasitic power consumption and heat loss associated to particles fluidization [245]. In the case of the CSP-CaL integration in which CaO particles are preheated before entering the carbonator with the CO2 exhaust stream [145,194], an indirect gas-solid heat exchanger is needed in the carbonator side to avoid direct contact between CaO and CO2, which could lead to partial carbonation with the consequent reduction in the carbonation heat transfer to the power cycle. Indirect solid gas heat exchange could be performed by using multiple heat transfer plates conveniently spaced to allow the flow of material to be heated inside [246]. Another hightemperature solid-gas heat exchanger was proposed by Al-Ansary et al. [247] in which particles circulate on the shell side through tubes arrangements while the CO2 passes inside the tubes. Moving packed-bed heat exchangers implementing shell-and-tube and finned shell-and-tube designs were investigated by Ho et al. [245]. Tests showed that the particle-side heat transfer (a) (b) 87 coefficient could reach values up to ~100 W/m2-K. Regarding the solid-solid heat exchanger, a possible system would use two solid-gas indirect heat exchangers with one intermediate heattransfer fluid recirculated within the bulk of both solids. Because of the high temperature needs, liquid-metals could be used as fluid transfer to this end. 3.8.3 Solids conveying and gas separation Available technologies to transport high temperature particles up-down the receiver include: mine hoist, bucket elevator, pocket elevator, screw conveyor, OLDS elevator, pneumatic conveyors, conveyor belts, cleated conveyor belts, metallic belted conveyors, masses elevators, bucket wheels, linear induction motor powered elevators, and electromagnetic field conveyors [235]. De la Verne [248] analyzed a basic skip at the end of a wire rope which showed an overall efficiency of 0.748. The OLDS elevator employs a circular casing rotating around a stationary screw or helix with a small parasitic power consumption [235]. The type of solar receiver technology determines the method employed for solid conveyance. Some commercial technology applicable to the CSP-CaL integration are: - OLDs elevator, which efficiency of is around 8% [235]. - Conventional Bucket lift. A commercial design for a bucket lift with 120 m transport height at ~2.4x105 kg/hour (260 tons/hour) and 200°C with a 171 kWe motor results in a 50% efficiency [235]. - Mine hoist with an insulated container (skip). Mine hoist consists of a skip at the end of a wire rope which is wrapped around a large motor-powered drum. The material is loaded into the skip at the bottom of a shaft, the skip is raised, and the material is discharged at the top. Mine hoist efficiency is largely determined by the weight of material hoisted and the speed at which it is hoisted. De la Vergne [248] shows an overall lift efficiency of 0.748 for two skips hoisting in balance on a single drum. - Pocket elevator. This system consists of a series of specially shaped containers (pockets) that travel along a loop of steel cable between two pulleys. The solids fall into the pockets at the bottom pulley and is discharged at the top pulley when the pocket is inverted [235]. Pocket elevators are reported to handle materials up to 500 °C and 930 °C. - Belt conveyors. Belt conveyors are cost effective only when there is a relatively high production rate and the transport distance is significant [248]. Many types of feeders exist for belt conveyors, but the most popular is the vibrating feeder. It is not recommended 88 to use belt conveyors to transport dry concentrate. In this case the recommended option is using screw conveyors. - Screw conveyors. This technology has been considered as suitable for Calcium-Looping applications [249]. In a commercial-scale, with large volumes of solids circulating through the calciner side, the solid loads will vary due to changes in the solar irradiation making the screw conveyors mechanically inefficient [75]. Screw conveyors were used in the Carina European project [93,250] to fed CaCO3 into the calciner. The flow rate of solids through the calciner was properly controlled by using the screw conveyors but the temperature range is a challenge for this system [250]. Since CO2 and CaO are stored separately, solid-gas separation after calcination is fundamental within the CSP-CaL integration. The same occurs in the power production side, where CO2 exiting the carbonator and headed to the turbine has to be free of particles. The most common method to separate particles from a gas stream is by means of cyclones, which are based on centrifugal separation. Thus, particles are forced out against the outer wall and then eliminated through an aperture at the bottom of the unit. Cyclones allow a high separation efficiency (95-100%), they are robust and allow operating under high-temperature with low energy requirement. However, the use of cyclones for small size particles [14], typically lower than 20 µm, is limited. A post-cyclone was proposed by Ray et al. [69] to improve the efficiency when working with fines. The post cyclone consists of a cylindrical annular shell located on top of the vortex finder. Both experiments and simulations indicate a 50% decrease in emission of <5 µm size particles. 89 4. Discussion After the summary and discussion of main results achieved along the present thesis, this section further discusses about the fulfillment of the objectives set out in Section 2. 4.1. Calcium-Looping for CO2 capture The multicycle CaO behavior, and therefore the calcination and carbonation reactions, are highly dependent on the process conditions, mainly temperature, pressure, particle size and atmosphere composition in the reactors. Carbonator models formulated until now for simulating postcombustion CO2 capture are based on lab-scale multicyclic CaO conversion results carried out by regenerating CaO under low CO2 partial pressure. From the analysis of realistic lab-scale results, the novel carbonator model proposed in this thesis incorporates in the calculation of the capture efficiency the conversion in the diffusion-controlled stage of carbonation, which was found as relevant according TGA tests at realistic operation conditions. In this way, objective 1 proposed at the beginning of the thesis is tackled, allowing a more accurate evaluation and prediction of carbonator’s performance. Simulations show that the capture efficiency is improved by increasing the residence time in the carbonator, which is a novel find. Thus, for typical inventories of solids in the carbonator used in previous models, the molar flow rate of CaO recirculated can be substantially decreased while a high capture efficiency is obtained. This reduces energy penalty in the operation of power plants since the amount of heat required for calcination would be lowered and the energy necessary for material transport would be also reduced. Taking into account the important effect of solids residence time in the carbonator on the CO2 capture efficiency, a novel CFPP-CaL process scheme has been developed. This satisfies objective 2 proposed to adapt the integration schemes to the conditions that improve the process performance considering industrial scale operation. Simulations over the proposed CFPP-CaL scheme show that the integration energy penalty is significantly reduced as the solids residence time is increased within the practical operation range of a few minutes, reaching values as low as 4 and 7% points over the reference plant efficiency. 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J Chem Eng JAPAN 1998;31:281–4. doi:10.1252/jcej.31.281. [159] Peng Q, Yang X, Ding J, Wei X, Yang J. Design of new molten salt thermal energy storage material for solar thermal power plant. Appl Energy 2013;112:682–9. doi:10.1016/j.apenergy.2012.10.048. [160] Ho CK. Advances in central receivers for concentrating solar applications. Sol Energy 2017. doi:10.1016/j.solener.2017.03.048. 112 Annexes (complete copy of publications) ANNEX 1: Ortiz C, Chacartegui R, Valverde J, Becerra J, Perez-Maqueda L. A new model of the carbonator reactor in the calcium looping technology for post-combustion CO2 capture. FUEL 2015;160:328–38. doi:10.1016/j.fuel.2015.07.095. A new model of the carbonator reactor in the calcium looping technology for post-combustion CO 2 capture C. Ortiz a , R. Chacartegui a, ⇑ , J.M. Valverde b , J.A. Becerra a , L.A. Perez-Maqueda c a Energy Engineering Department, University of Seville, Camino de los Descubrimientos s/n, 41092 Sevilla, Spain b Faculty of Physics, University of Seville, Avenida Reina Mercedes s/n, 41012 Sevilla, Spain c Instituto de Ciencia de Materiales de Sevilla (C.S.I.C. – Univ. Sevilla), Americo Vespucio 49, 41092 Sevilla, Spain highlights A new carbonator model is proposed taking into account realistic CaL conditions. Carbonation in the diffusion stage is considered to predict the capture efficiency. High capture efficiency may be achieved by operating with high particles residence time. New model allows accurate modelling for extended residence’s time. Results are in good agreement with data from pilot-scale tests. article info Article history: Received 19 May 2015 Received in revised form 28 July 2015 Accepted 29 July 2015 Available online 4 August 2015 Keywords: CO 2 capture Ca-looping Limestone calcination Diffusion phase Calcium looping Carbon dioxide capture abstract The Ca-Looping (CaL) process is considered as a promising technology for CO 2 post-combustion capture in power generation plants yielding a minor penalty on plant performance as compared with other capture technologies such as conventional amine-based capture systems. This manuscript presents a new carbonator reactor model based on lab-scale multicyclic CaO conversion results, which take into account realistic CaO regeneration conditions that necessarily involve calcination under high CO 2 partial pressure and high temperature. Under these conditions, CaO conversion in the diffusion controlled stage is a relevant contribution to the carbonation degree during typical residence times. The main novelty of the model proposed in the present work is the consideration of the capture efficiency in the diffusion controlled phase of carbonation. It is demonstrated that increasing the residence time by a few minutes in the carbonator yields a significant improvement of the capture efficiency. Model predictions are shown to agree with experimental results retrieved from pilot-scale tests. The new model allows a more accurate evaluation and prediction of carbonator’s performance over a wider range of residence times. The results obtained may be relevant for the optimization of CaL operation parameters to be used in real power plants. Ó2015 Elsevier Ltd. All rights reserved. 1. Introduction Taking into account the current energy system and its expected evolution, capture and permanent geological storage of CO 2 from fossil fuel power plants is considered as a necessary technology to be developed at commercial scale in the short term for mitigation of global warming. The IPCC [1] has evaluated existing and emerging technologies of CO 2 capture and storage (CCS) and concludes that, in most scenarios, CCS should contribute by a 15– 55% to the cumulative global mitigation effort until 2100. Amongst the CO 2 capture technologies potentially foreseeable for commercial development, the Ca-Looping (CaL) process is highlighted as one with good prospects for post-combustion capture, mainly due to the possibility of being implemented in already existing power plants and the minor penalty on plant performance as compared to other capture technologies such as conventional amine-based capture systems [2,3]. A further advantage of the CaL technology is the low price, wide availability and harmlessness towards the environment of the sorbent precursor raw material, namely natural limestone [2,4,5]. The CaL technology for CO 2 post-combustion capture is based on the reversible chemical reaction of carbonation–calcination of CaO. The basic cycle involves CO 2 capture from the flue gas stream of a power plant using lime (CaO) derived from calcination of natural limestone: http://dx.doi.org/10.1016/j.fuel.2015.07.095 0016-2361/Ó2015 Elsevier Ltd. All rights reserved. ⇑ Corresponding author. Tel.: +34 954487615. E-mail address: [email protected] (R. Chacartegui). Fuel 160 (2015) 328–338 Contents lists available at ScienceDirect Fuel journal homepage: www.elsevier.com/locate/fuel CaO ðsÞ þCO 2 ¢CaCO 3ðsÞ D H 0 r ¼178 kJ mol ð1Þ Flue gases from coal-fired power plants generally contain a mole fraction of CO 2 in the range of 10–15% [6,7], whereas typical residence times in the carbonator reactor would be on the order of minutes. Taking into account these constraints, optimum carbonation temperatures are around 650 °C for a quick enough reaction kinetics and low value of the equilibrium CO 2 concentration in order to achieve significant efficiencies of CO 2 capture (around 80–90%) [8]. Efficient gas–solid contact and heat/mass transfer would be ensured by the use of circulating fluidized-bed reactors (CBF). CFBs are typically operated at atmospheric pressure under the fast fluidization regime, with gas velocities of the order of 5–10 m/s [9,10]. The partially carbonated particles are driven into a second fluidized bed reactor (calciner), where CaO is regenerated by calcination under high temperatures and necessarily high CO 2 concentration [11–14]. Once CO 2 is captured in the carbonator and heat from the exothermic reaction is recovered, the flue gas almost free of CO 2 is released into the atmosphere. After a residence time in the carbonator of a few minutes, the carbonated solids are circulated into the calciner. Coal is burned into the calciner reactor using pure O 2 (oxy-combustion) to increase the temperature up to the value required for the endothermic calcination reaction to occur fast and avoid CO 2 dilution at the same time. Thus, CO 2 is retrieved from the calciner for compression, transport and geological storage or other uses. The regenerated sorbent produced in the calciner is returned to the carbonator for a new cycle. In order to predict the efficiency of CO 2 capture by integrating the CaL technology on existing power plants, a number of models have being proposed in the literature [15–18]. Besides, several pilot-medium scale plants (up to 1.7 MWth) demonstrate a high CO 2 capture efficiency, which raises hopes on a successful scale-up of the technology [13,14,19–21]. Interestingly, the CaL technology is also suitable to be integrated with biomass combustion as shown in a pilot-scale plant (300 kWth), which would allow producing power with negative CO 2 emissions [20]. The multicyclic calcination/carbonation behavior of CaO used in current carbonator models has been inferred from thermogravimetric analysis (TGA) results in which calcination is usually carried out under low CO 2 partial pressure mainly due to technical difficulties. Under these conditions, most of the carbonation reaction in short residence times occurs on the surface of the CaO particles through a kinetically controlled fast phase whereas the subsequent solid-state diffusion controlled phase is comparatively negligible as will be seen in detail below. The present manuscript proposes a new carbonator model that does consider the relevant effect of realistic calcination conditions, necessarily involving calcination under high CO 2 pressure, on the multicyclic CaO conversion behavior. In order to stress the relevant role of calcination conditions on the sorbent performance, the next section is devoted to briefly review recent TGA results that will be employed in the formulation of the new carbonator model. We follow up with the development of the carbonator model. Finally, capture efficiency results from the new model are compared with predictions by models reported in the literature and pilot-scale results. Nomenclature Acarbonator section, m 2 CCO2 average CO 2 concentration, mol/m 3 CCO2;eq equilibrium concentration of CO 2 , mol/m 3 Deff intrinsic diffusion constant, m 3 /(mol s) D eff equivalent diffusion constant, m 3 /(mol s) ECO2 carbon capture efficiency ECO2;eq maximum carbon capture efficiency ECO2;Dcarbon capture efficiency in the diffusion regime ECO2;Kcarbon capture efficiency in the kinetic regime f0inlet molar fraction of CO 2 favolumetric fraction of CaO that reacts in the kinetic reaction regime feequilibrium molar fraction of CO 2 F0mole flow of fresh makeup limestone, mol/s FCO2 mole flow of CO 2 in flue gas entering the carbonator FCO2;cal mole flow of CO 2 originating in the calciner, mol/s FPmole flow of solids purge in the calciner, mol/s FRmole flow of CaO coming from the calciner, mol/s hthickness of the product layer of a sorbent particle, nm ksintrinsic kinetic constant m 4 /(mol s) T0time of TGA multicyclic test, s MCaO molar mass of CaCO 3 , g/mol mCO2 volume flow rate per MW of a typical power plant of CO 2 in flue gas entering the carbonator, m 3 /s Ncalcination–carbonation cycles NCa mol of Ca in the carbonator, mol Ppressure (bar) rave;Daverage reaction rate in the diffusion regime, s 1 rave;Kaverage reaction rate in the kinetic regime, s 1 rN;Dparticle reaction rate in the diffusion regime, s 1 rN;Kparticle reaction rate in the kinetic regime, s 1 Save;Kaverage reaction available surface, m 1 SNreaction available surface in the N cycle, m 2 /m 3 Ttemperature (°C) ttime, s tKcharacteristic time at which kinetic phase end, s tmax Particle residence time in carbonator, s Vgas volume flow rate per MW of a typical power plant in flue gas entering the carbonator, m 3 /s VM;CaCO3 molar volume, m 3 /mol Wssolid inventory in the carbonator per MW of a typical power plant, kg Ws;csolid inventory in the calciner per MW of a typical power plant, kg Xcarbonation degree of a CaO particle Xaveaverage conversion of the sorbent Xave;Daverage conversion of the sorbent in the diffusion phase Xave;Kaverage conversion of the sorbent in the kinetic phase Xmax;avemaximum average carbonation degree of CaO in the solid population Xmax;ave;Dmaximum average carbonation degree of CaO in the solid population in the diffusion phase Xmax;ave;Kmaximum average carbonation degree of CaO in the solid population in the kinetic phase XNmaximum carbonation degree of CaO in the N cycle XND maximum carbonation degree of CaO in the N cycle in the diffusion phase XNK maximum carbonation degree of CaO in the N cycle in the kinetic phase Xrresidual conversion capacity of a sorbent particle j deactivation constant of a sorbent particle r N particle fraction in the N cycle q CaO CaO density, g/m 3 q g density of gas phase, g/m 3 s average residence time in the carbonator, s s a active space time of the carbonator, s C. Ortiz et al. / Fuel 160 (2015) 328–338 329 2. A brief review on the role of calcination conditions on the multicyclic CO 2 capture behavior of limestone Results of TGA tests show that the multicyclic CO 2 capture capacity of limestone derived CaO is strongly influenced by the conditions under which calcination is performed [22–24]. The most feasible method in practice to generate the heat required in the calciner while at the same time avoiding CO 2 dilution is to burn coal by oxy-combustion [4,11–14]. Thus, calcination is carried out in an environment containing between 70% and 90% vol. concentration of CO 2 at atmospheric pressure [11,23,25–27]. Calcination kinetics becomes extremely slow under high CO 2 partial pressures nearby equilibrium (T870–900 °C) [22]. This makes necessary to increase the temperature over 930 °C to rise the CO 2 partial equilibrium pressure well above the CO 2 partial pressure in the calciner and yield sufficiently fast reaction kinetics in short residence times [14,21,28]. 2.1. Sorbent deactivation The combination of high temperature and high concentration of CO 2 during calcination enhances the progressive deactivation of the sorbent as the number of cycles builds up due to enhanced sintering [29–32]. Enhanced sintering of CaO has a strongly negative effect on the active surface of the solids available for the gas–solid reaction in the kinetically driven fast phase. Grasa and Abanades [33] proposed to fit the observed evolution of CaO conversion X N (ratio of mass of CaO converted to CaCO 3 to mass of CaO initial) with the number of cycles by means of the Eq. (2): X N ¼1 1 1X r þjNþX r ð2Þ where j = 0.52 and X r = 0.075 are the deactivation constant and the residual conversion, respectively. This equation is however based on results from multicyclic calcination/carbonation TGA tests in which the sorbent is regenerated by calcination under low CO 2 partial pressure. Thus, conversion after 1st calcination (X 1 ) is close to unity. This high value of conversion is due to the relatively high porosity of the CaO particles derived by short-time calcination under low CO 2 partial pressure. Despite that Eq. (2) cannot reflect the effect of calcination under high CO 2 pressure as would be the case in the CaL technology, it has been routinely used in theoretical studies as representative of the sorbent behavior in order to model the carbonator reactor [15–18]. In more recent TGA studies [34,35], a modified equation was proposed using data obtained from multicyclic calcination/carbonation tests under more realistic calcination conditions involving high CO 2 partial pressure (70% v/v CO 2 , 30% air) at temperatures of 950 °C as well as fast transitions between the calcination and carbonation stages. The modified equation proposed is: X N X 1 ¼X r X 1 þ1 jðN1Þþð1 X r X 1 Þ 1 0 @1 Að3Þ where j = 0.776 and X r = 0.077 and X 1 = 0.48 is CaO conversion after 1st calcination. Carbonation/calcination conditions in these tests where 650 °C/5 min for carbonation under 15% CO 2 and 950 °C/5 min for calcination under 70% CO 2 . Eq. (3) looks similar to Eq. (2), but with the main difference that it takes into account the smaller value of conversion in the first cycle after calcination under high CO 2 partial pressure. Besides, the deactivation constant is higher as compared to that derived from tests carried out calcining under low CO 2 partial pressure. In the real process, a makeup flow of fresh limestone must be continuously fed into the calciner in order to compensate for sorbent deactivation. Limestone is therefore first calcined in the stationary state under high CO 2 partial pressure at the same conditions that the partially carbonated solids are calcined for CaO regeneration. Thus, the use of Eq. (3) for modelling the carbonator reactor might help simulating more realistic CaL conditions. 2.2. Fast and slow carbonation stages Two well differentiated stages can be observed during carbonation of CaO solid particles [30,36–38]. The first stage takes place on the free surface of the solid through the nucleation and growth of CaCO 3 and is governed by the kinetics of the reaction between CaO and CO 2 . The end of the fast stage takes place when a 40–60 nm thickness product layer is formed, which makes inaccessible a large fraction of CaO in the interior of the particles for fast kinetically controlled carbonation [39]. Once the CaO free surface directly available for fast carbonation has been covered the reaction continues in a second stage which is controlled by the solid-state diffusion of CO 3 2 and O 2 ions through the product layer, which is a much slower process [40,41]. TGA results show that the solid-state diffusion controlled carbonation is greatly enhanced when CaO is regenerated under high CO 2 partial pressure as compared to calcination under low CO 2 concentration [26,42,43].Fig. 1 shows the time evolution of Fig. 1. Time evolution of sorbent weight% during precalcination and subsequent carbonation/calcination cycles. Carbonation at 650 °C for 5 min (15% CO 2 /85% air vol/vol). Different calcination conditions are indicated. Reproduced from data reported in [35]. 330 C. Ortiz et al. / Fuel 160 (2015) 328–338 sorbent weight during the first calcination/carbonation cycles of two TGA tests in which calcination was carried out either under high CO 2 partial pressure or air. The results demonstrate that carbonation in the fast initial phase is extraordinarily hindered when the sorbent is regenerated under high CO 2 concentration. Even though the reaction rate is similar for the two tests, the maximum carbonation conversion in the fast phase is significantly hampered for the sorbent regenerated under high CO 2 partial pressure essentially due to severe sintering and drastic reduction of the surface area [22]. On the other hand, carbonation in the diffusion controlled phase is notably enhanced when calcination is carried out under high CO 2 concentration especially in the first cycles. A close examination of TGA test results reported in the literature for CaO derived from other precursors such as steel slag (Fig. 2a) and results obtained from pilot scale tests (Fig. 2b) reveals also that in those cases the relatively slower solid-state diffusion controlled phase contributes substantially to carbonation for residence times of a few minutes. Fig. 3 shows the multicyclic evolution of CaO conversion in the kinetically controlled (X NK ) and diffusion controlled (X ND ) carbonation stages derived from TGA tests reported in [35] in which calcination was carried out under high CO 2 partial pressure. Both X NK and X ND can be pretty well fitted by Eq. (3) using the parameters given in Table 1. As can be seen in Fig. 3, the contribution of the diffusion controlled phase to CaO conversion is quite relevant being of the same order than CaO conversion in the kinetically controlled phase for 5 min overall carbonation periods. Moreover, an increase of the carbonation phase above 5 min would lead to higher conversion in the diffusive phase since carbonation in this stage grows roughly linearly with time until carbonation is completed [42]. Thus, it seems clear that a carbonator model should consider carbonation in the diffusion stage in order to predict realistic values of efficiency and to draw useful conclusions on the optimum operating parameters in a commercial plant. This is the main goal of the present manuscript. 3. Carbonator model approach 3.1. Kinetic model As mentioned in the introduction, in order to predict the CO 2 capture efficiency of the CaL technology, many authors [15–18] propose to simplify the CaO conversion behavior by assuming that it reaches a maximum value, X N , at a constant rate (proportional to the free surface available of CaO and the thickness of the CaCO 3 layer formed) in a time t lim , after which the reaction rate drops to zero (solid lines in Fig. 4a). This simplifying assumption would be reasonable if calcination was performed under low CO 2 partial pressure. However, as pointed out above, the diffusive carbonation Fig. 2. (a) Time evolution conversion of steel slag during the first, 5th, 10th, 15th, and 20th cycles (reproduced from [48] with permission). (b) Average CO 2 carrying capacity (X ave ) vs. time and reaction constant (k sB ), derivation for sample removed from the CFB carbonator during operation in pilot-scale test (reproduced from [28] with permission). Fig. 3. CaO conversion in the kinetically controlled carbonation stage (X NK ) and in the diffusion controlled stage (X ND ) as a function of the cycle number. Precalcination and calcinations conditions: CO 2 vol% = 70%, t= 5 min, T= 950 °C. Carbonation conditions: CO 2 vol% = 15%, t= 5 min, T= 650 °C. The inset shows the time evolution of sorbent wt% in the first 3 cycles. Obtained from data originally reported in [42]. Table 1 Fitting parameters for sorbent deactivation curves of conversion in the kinetic and diffusion controlled stages (Fig. 3). Carbonation stage Kinetic Diffusion Time 0.3 min 5 min Deactivation constant j K = 0.676 j D = 0.871 Residual conversion X rK = 0.0296 X rD = 0.0408 1st cycle conversion X 1;K = 0.0218 X 1;D = 0.263 C. Ortiz et al. / Fuel 160 (2015) 328–338 331 phase plays a relevant role for calcination under high CO 2 partial pressure whereas the kinetically controlled carbonation stage is severely hampered. Using the multicyclic TGA results reported in [34,35,42], we develop below a modified carbonator model built upon the model proposed by Alonso et al. [18]. The main novelty of our model is that carbonation in the diffusion controlled phase will be considered to predict the carbonation efficiency. Fig. 4(a) shows experimental results from TGA tests in which calcinations were carried out under low CO 2 partial pressure. The solid lines are the kinetic curves assumed in the carbonator model described in [18]. As can be seen, the diffusion controlled stage is negligible in the first cycles although as the number of cycles is increased the diffusive conversion becomes more important (see results for cycle N=20inFig. 4(a)). As seen in Fig. 4(a) the time for the kinetic stage (t k ) decreases with the number of cycles as conversion in the diffusive phase is enhanced, which is however dismissed in the theoretical model (solid lines). Fig. 4(b) illustrates the kinetic curves proposed in our work as representative of the experimental curves reported in [42] in which calcinations were carried out under high CO 2 partial pressure. It is assumed that the time evolution of CaO conversion can be approximated by two straight lines of constant slope, one corresponding to the kinetically controlled carbonation stage and another one for the diffusive carbonation stage. In this way, the new kinetic model has into account the relevant role of the diffusion controlled phase. According to experimental measurements, conversion would be increased roughly linearly with time during the first minutes in the diffusion controlled stage against the widely accepted conception that once the fast kinetically controlled phase is completed, additional carbonation by solid-state diffusion is negligible. Thus, it may be anticipated that an important parameter influencing the carbonation efficiency would be the residence time of the solids in the carbonator. According to the kinetic model proposed in this work (solid lines in Fig. 4(b)), the reaction rate can be expressed as a function of residence time in the carbonator by the following expression: r N i ¼r N K ¼ X NK t K for t 6t K r N D ¼ X ND T 0 t k for t K <t6t max () ð4Þ where r N_i is the reaction rate in the i-phase (either kinetic or diffusion), t K is the time the kinetic phase takes, t max is the residence time of the particles in the carbonator, and X ND /(T 0 t K )is the rate of diffusive conversion that will be derived from experimental results. Using experimental data reported in [42], the reaction rates in the kinetic phase (r Nk ) and diffusive phase (r ND ) can be adjusted as a function of the number of cycles Nusing Eq. (3) as shown in Fig. 5. According to Bhatia and Perlmutter [38], the rate of CaO conversion in the kinetically controlled fast phase at atmospheric pressure, can be expressed by a first-order kinetic law: dX dt ¼r NK ¼k s S N ð1XÞ 2 3 ðC CO2 C CO2eq Þð5Þ where C CO2 and C CO2eq are the actual and equilibrium CO 2 concentrations, respectively, and k s is the kinetic constant. S N is the CaO specific area available for reaction after Ncycles, which is proportional to the particle conversion degree. As can be seen in Fig. 5, the reaction Fig. 4. CaO conversion as a function of time for different cycles showing in (a) the theoretical approximation proposed elsewhere for the kinetic model (solid lines) to describe the progress of the carbonation reaction with time (reproduced from [18] with permission). Data points are experimental results from TGA tests in which calcination was carried under low CO 2 concentration. In (b) the kinetic model curves proposed in the present article (solid lines) are shown superposed to experimental data from TGA tests reported in [49] in which calcination was carried out under high CO 2 concentration (dotted lines). Note the different scales in (a) an (b) for the CaO conversion degree. Fig. 5. Rate of kinetic and diffusive carbonation (r ND ,r NK respectively) as a function of number of cycles from experimental data reported in [42] obtained by means of TGA tests. The solid lines are best fit curves using Eq. (3) yielding residual values of the kinetic and diffusion rates of 0.096 and 0.0082, respectively and deactivation constants of the kinetic and diffusion rates of 0.315 and 0.917, respectively. 332 C. Ortiz et al. / Fuel 160 (2015) 328–338 rate in the kinetic phase (r Nk ) decays with the number of cycles due to the effect of sintering that leads to a decrease of the available surface S N ). By integrating Eq. (5) and fitting it to the experimental data reported in [42], we obtain a value of the kinetic constant k s = 6.7 10 10 m 4 /(mol s), which is similar to the value reported by other authors [38,44] also for natural limestone but under different operating conditions as regards calcination. This suggests that the kinetics of carbonation in the fast phase is essentially determined by the CaO surface area and does not depend on the calcination conditions. The essential effect of calcination under CO 2 would be thus to enhance sintering, which greatly reduces the available surface area for carbonation and therefore the duration of the kinetic phase. The rate of conversion in the diffusion controlled phase can be expressed using an effective diffusion constant D eff [40]: dX dt ¼r ND ¼D eff ðC CO2 C CO2eq Þ 2 3 ð1XÞ  1 3  2 3 ðZþð1ZÞð1XÞÞ  1 3 ð6Þ where Zis the ratio of the molar volume of CaCO 3 and CaO. Eq. (6) can be adjusted to the experimental data (Fig. 5) to obtain a value for the diffusion constant. As seen from the experimental results (Fig. 5) the rate of carbonation in the diffusive stage remains after 2–3 cycles approximately constant with the number of cycles. In [42], some of the TGA tests presented were made by prolonging the carbonation stage up to 30 min showing that the rate of carbonation in the diffusion controlled phase did not change appreciably over this time interval. Thus, we may neglect as a first approximation the dependence of the diffusive reaction rate on conversion and use the simplified expression: r ND D  eff ðC CO2 C CO2eq Þð7Þ where D  eff is an effective diffusion constant. By using r ND 0:01 min 1 (Fig. 5) a value of D  eff = 6.5 10 5 m 3 /(mol s) is obtained. 3.2. Capture model In Fig. 6 we show the global scheme for the CaL system with the definition of main parameters used for defining and evaluating CO 2 capture. The average conversion of the CaO particles present in the reactor (X ave ) is given by the sum of conversion in the kinetic fast phase (X ave,k ) and conversion in the diffusive phase (X ave,D ): X ave ¼X ave;K þX ave;D ð8Þ For a given continuous supply of fresh limestone into the calciner and a given solids purge, the fraction of particles subjected to a given number of cycles Nis given by [30]: r N ¼F 0 F N1 R ðF 0 þF R Þ N ð9Þ The average maximum conversion (X max,ave ) can be calculated as the weighted sum of conversions after Ncycles: X max;ave ¼X N¼1 N¼1 r N X N ð10Þ X max;ave ¼X max;ave;K þX max;ave;D ¼X N¼1 N¼1 r N X NK þX N¼1 N¼1 r N X ND ð11Þ The average reaction rate depends on the phase the particles are reacting as: r ave i ¼r ave; K ¼ X max;ave;K t K si t 6t K r ave; D ¼ X max;ave;D T 0 t k si t K <t6 s 8 < :9 = ; ð12Þ where s is the average residence time of the particles in the carbonator: s ¼N Ca F R ¼W s 56F R ð13Þ The average conversion of the particles leaving the carbonator can be obtained from the sum of the average particle conversion reacting in the fast carbonation phase (X a v e;K ) and the average particle conversion reacting in the diffusive phase (X a v e;D ): X ave ¼X ave;K þX ave;D ¼f a Xj 6t K þð1f a ÞXj >t K ð14Þ Xj 6t K ¼R t K 0 r ave; K t 1 s e  t s dt 1e  tK s ð15Þ Xj >t K ¼X max;ave;K þR s t K r ave; D t 1 s e  t s dt 1e  tK s ð16Þ FCO2 (1-ECO2) carbonator T=650ºC Ws calciner T=950ºC Ws,c FR F0 FP FCO2 FRXave FCO2,cal Fig. 6. Scheme of the CaL post-combustion capture model. C. Ortiz et al. / Fuel 160 (2015) 328–338 333 where t D ¼ s t K and f a is the fraction of particles reacting in the kinetic phase, which is given by [18] f a ¼1e  tK s  ð17Þ The average surface area for reaction in the kinetic phase can be expressed by [44]: S ave K ¼V M CaCO3 X max;ave;K M CaO h q CaO ð18Þ Since conversion is typically low except for the first cycles, Eq. (5) can be simplified by dismissing the dependence of the reaction rate on conversion as done in [18]. Thus, it is possible to express the average reaction rate in each carbonation phase as: r ave K ¼k s S ave K ðC CO2 C CO2eq Þð19Þ r ave D ¼D  eff ðC CO2 C CO2eq Þð20Þ The capture efficiency in the carbonator E CO2 can be calculated as [18]: E CO2 ¼F R F CO2 X ave ð21Þ E CO2;K ¼F R F CO2 X ave;K ð22Þ E CO2;D ¼F R F CO2 X ave ; D ð23Þ where the capture efficiency in the fast phase (E CO2;K ) and the diffusive phase (E CO2;D ) have been considered as separate contributions in order to evaluate their relative importance. Assuming that the gas passes in plug flow through a bed of perfectly mixed solids, the carbon mass balance in the gas phase in a differential element of the carbonator reactor can be written as: F CO2 dE CO2 dz ¼A q CaO V M CaO ðf a r ave K þð1f a Þr ave D Þð24Þ Integrating Eq. (24) along the carbonation reactor, it results w¼ f 0 f 0 f e f 0 E CO2 þf 0 ðf 0 1Þ ðf 0 f e f 0 Þ 2 ln ðf 0 f e Þþðf 0 f e f 0 ÞE CO2 f 0 f e  "# ð25Þ where z¼F CO2 M CaO W s  q g ðS ave K f a k s þD eff sim ð1f a Þwð26Þ and f e and f 0 are the equilibrium and inlet molar fraction of CO 2 , respectively. 4. Model results A series of parametric calculations have been carried out to determine the CO 2 capture efficiency of the carbonator under different operating conditions. The calculations are made for a typical coal combustion power plant using a value of 0.15 for the volumetric fraction of CO 2 in the flue gas entering the carbonator and a flow rate F CO2 = 0.1 kg CO 2 /s as typical values [18]. In the first run we will use as key input parameters the solid inventory in the carbonator reactor (W s ), the flow rate of fresh limestone makeup (F 0 ) fed into the calciner and the flow rate of solids entering the carbonator coming from the calciner (F R ). As for the rest of parameters, the same representative values employed in previous works [18] will be used (Table 2). Figs. 7 and 8 show the calculated capture efficiency as function of the total solid inventory in the reactor W(normalized per MWth) for different values of F R and F 0 . As seen in Figs. 7 and 8, the model proposed in this paper yield results which are less sensitive than those obtained from the reference model [18] to changes in the solids recirculation flow rate (F R ) and the flow rate of fresh limestone introduced in the cycle (F 0 ). On the other hand, for fixed values of F 0 , increasing the solids inventory in the carbonator yields a significant increase of the capture efficiency as a result of having considered the enhanced conversion in the diffusive phase. According to the reference model [18], increasing the solids inventory over a certain value does not lead to an increase of the carbonation efficiency despite it has not reached the maximum attainable value (0.92) according to the inlet and equilibrium CO 2 concentration for the carbonation reaction at 650 °C. This would be the maximum capture efficiency Table 2 Values of the parameters used for the calculation of capture efficiency. Parameter Value References m CO2 /MW (kg/s) 0.1 [18] f 0 0.15 [18] P(bar) 1 [18] T(°C) 650 [18] h(nm) 50 [18,40] K s (m 4 /mol s) 4 10 10 [18] 0 0,1 0,2 0,3 0,4 0,5 0,6 0,7 0,8 0,9 1 0 50 100 150 200 250 300 350 400 ECO2 Ws (kg/MW) F0/FCO2=0,1 FR/FCO2=5 Calculated from model proposed in [18] FR/FCO2=5 Extracted from data shown in [18] FR/FCO2=5 Proposed model in this work FR/FCO2=20 Calculated from model proposed in [18] FR/FCO2=20 Extracted from data shown in [18] FR/FCO2=20 Proposed model in this work Fig. 7. Capture efficiency calculated as a function of the solids inventory in the carbonator for different F R /F CO2 ratios according to the model presented in [18] and the new model proposed in this work. 334 C. Ortiz et al. / Fuel 160 (2015) 328–338 achievable if the solids had fully reacted, which is not theoretically allowed if carbonation in the diffusion controlled phase is fully neglected as assumed in previous modelling works [18]. In the reference model [18] the active fraction of solids drops significantly after a short time once the kinetic controlled phase is ended. In contrast, the solids are considered to remain active for a long time in the proposed model in the present work as long as the subsequent carbonation in the diffusion controlled phase is taking place. Figs. 7 and 8 demonstrate that our model predicts a greater capture efficiency for longer residence times of the particles in the carbonator. For example, Fig. 8 shows that for a ratio F R /F CO2 = 5, and a solids inventory of about 200 kg/MW, the proposed model predicts a higher capture efficiency than the reference model [18].Fig. 9 shows the capture efficiency calculated as a function of the residence time in the carbonator using F R /F CO2 = 5 and F 0 /F CO2 = 0.01 and varying the value of solids inventory, which changes the residence time s (Eq. (13)). This figure shows also the ratio of the carbonation efficiency in the kinetic controlled phase to that in the diffusion controlled phase (right axis). 0 0,1 0,2 0,3 0,4 0,5 0,6 0,7 0,8 0,9 1 0 50 100 150 200 250 300 350 400 ECO2 Ws (kg/MW) F0/FCO2=0,01 FR/FCO2=5 Calculated from model proposed in [18] FR/FCO2=5 Extracted from data shown in [18] FR/FCO2=5 Proposed model in this work FR/FCO2=20 Calculated from model proposed in [18] FR/FCO2=20 Extracted from data shown in [18] FR/FCO2=20 Proposed model in this work Fig. 8. Carbonation efficiency as a function of the solids inventory in the carbonator at different F R /F CO2 ratios according to the model presented in [18] and the new model proposed in this work. Fig. 9. Left axis: carbonation efficiency as a function of the residence time in the reactor (modifying the solids inventory, Eq. (13)). Right axis: ratio of CO 2 capture efficiency in the kinetic controlled phase to carbonation efficiency in the diffusion controlled phase as a function of residence time. Calculations are made for fixed values of F R /F CO2 and F 0 /F CO2 as indicated. 2712172227 0,3 0,4 0,5 0,6 0,7 0,8 0,9 100 150 200 250 300 350 400 450 500 550 FR/FCO2 ECO2 τ(s) Ws=400 kg/MW F0/FCO2=0,01 ECO2 (calculated according [18]) ECO2 (proposed model) Fig. 10. Carbonation efficiency as a function of the residence time in the carbonator, which is varied by changing the F R /F CO2 ratio (upper horizontal axis). Calculations made for fixed values of W s and F 0 /F CO2 as indicated. C. Ortiz et al. / Fuel 160 (2015) 328–338 335 Fig. 9 shows that, according to the reference model [18], the capture efficiency is not improved further after a residence time of about 300 s since the fraction of active particles (considered as only those able to react in the fast phase) diminishes with time quickly. On the other hand, the capture efficiency is further improved by taking into account enhanced conversion in the diffusive phase as shown by the results obtained from the model proposed. As can be seen in Fig. 9 the capture efficiency in the diffusive phase (E CO2,D ) becomes the major contribution to the total carbonation efficiency just after tens of seconds. Besides, the capture efficiency continues to increase with the residence time and may reach a rather high value at residence times above 500 s. Fig. 10 shows the effect on the capture efficiency of increasing the residence time (horizontal bottom axis) by decreasing the flow rate of recirculated solids between the reactors (F R , horizontal top axis) while maintaining fixed a value for the solids inventory. As may be observed, by taking into account conversion in the diffusive phase, the capture efficiency is notably increased by increasing the value of the average residence time of the particles ( s ) while the solids recirculation flow rate is decreased (with a constant value of solids inventory in the carbonator). An additional benefit of reducing the recirculation flow rate can be that the solids activity would be extended for longer times since the frequency of calcinations is reduced, which would serve to minimize the makeup flow of fresh limestone needed. The results plotted in Fig. 11 correspond to simulations in which the solids recirculation flow rate F R is kept fixed while the flow rate of limestone makeup into the calciner (F 0 ) is varied. Interestingly, it is seen that high capture efficiency may be achieved by operating with low values of F 0 when W s is increased as a consequence of having extended the lifetime of the particles that remain active in the carbonator by virtue of diffusive carbonation. We must note, however, that neither sulphation nor ashes, which will reduce the activity of the sorbent, have been considered in the model. An additional makeup flow of fresh limestone will have to be employed in practice due to irreversible sulphation and deactivation by ashes as reported in many studies [7,16,25,45–47]. Let us now compare predicted results by the proposed model with experimental results obtained from pilot-scale tests. Capture efficiency obtained in pilot-scale tests is usually reported as a function of an ‘‘active space time’’ [13,14,28], which is expressed as the product of the residence time times the fraction of active particles reacting in the fast stage and the average conversion in this stage s a ¼ s f a X a v e;K . Taking into account also carbonation in the diffusive phase, s a would be defined in our model as [28] 0,4 0,5 0,6 0,7 0,8 0,9 1 0,04 0,08 0,12 0,16 0,2 ECO2 F0/FCO2 FR/FCO2 = 5 Ws=200 kg/MW calculated using [18] Ws=200 kg/MW Extracted from data [18] Ws=200 kg/MW proposed model Ws=400 kg/MW calculated using [18] Ws=400 kg/MW proposed model Fig. 11. Capture efficiency as a function of the ratio of makeup flow rate to CO 2 flow rate (F 0 /F CO2 ) at different solid inventories (W s ) and ratios of the recirculation flow rate to CO 2 flow rate (F R /F CO2 ). Results from the proposed model in this work are compared with results from the model presented in a previous work [18]. Table 3 Operating conditions used in the simulations whose results are plotted in Fig. 12. V gas = 1.15 (m 3 /s)/MW f 0 = 0.15 W s (kg/MW) = 100–800 Condition 1 Condition 2 Condition 3 Condition 4 F R /F CO2 510510 F 0 /F CO2 0.1 0.1 0.01 0.01 0,3 0,4 0,5 0,6 0,7 0,8 0,9 1 1 10 100 ECO2/ECO2eq τa (s) INCAR Condions 1 Condions 2 Condions 3 Condions 4 Condions 2. Calculated using [18] Model proposed in [16] Fig. 12. Normalized capture efficiency as a function of the active space time. Results obtained from the proposed model in this work using diverse operating conditions (displayed in Table 3) are compared with results obtained from previous models [16] [18] and with experimental data reported from pilot-scale INCAR tests [28]. 336 C. Ortiz et al. / Fuel 160 (2015) 328–338 parameters of this reference plant are employed. Combustion of 205 tonnes per hour of coal takes place in the steam boiler to generate 1335.7 MW th with an 88.6% efficiency. For the steam cycle, a 42% thermal to electric steam turbine efficiency is assumed based on results reported for similar plants [18,37]. Considering a 10% of global process parasitic electricity consumption, the reference CFPP net power is 505 MW e . Thus, a 33.5% overall plant efficiency is established [7]. The composition of the flue gas exiting the power plant, which is driven to the CaL cycle, is detailed in Table 2. 3.2. CaL integration For the CaL-CFPP integration analysis, mass and energy balances in the process have been calculated by using the commercial software ASPEN PLUS TM . A schematic view of the integration is shown in Fig. 3. In order to simplify the model a number of main assumptions are made: (i) steady state operation is assumed, (ii) pressure losses in the heat exchanger equipment are neglected, (iii) solid–solid heat exchange is simulated as a transfer of heat between both solids, (iv) for all gas–solid heat exchangers an approach temperature of 10 °C is established, and (v) gas passes in plug flow through a bed of perfectly mixed solids in the reactors. For the sake of clarity, the integration scheme (Fig. 4) will be analyzed by dividing it into 3 main areas such as the carbonator, calciner and secondary steam cycle. 3.2.1. Carbonator model In the carbonator reactor, the CaO particles entering it from the calciner (at a flow rate F R ) and those already present in the bed (N Ca ) react with the flue gas flowing into it from the coal-fired power plant. CO 2 capture in this reactor is modeled according to an equilibrium reactor following the study reported by Ortiz et al. [32]. The operating carbonator temperature is set to 650 °C and it works at atmospheric pressure. An inventory of solids in the carbonator bed (W s Þof 200 tonnes of solids (400 kg/MW el )is considered as a typical value used in carbonator models [32,36,38]. The heat absorbed by the initial batch of solids to raise its operation temperature for the first time is not taken into account in the analysis since a steady operation regime is considered. Gas pressure loss across the carbonator reactor, which is operated in the continuous fluidized bed (CFB) regime, is calculated from the Kunii–Levenspiel (K–L) fluid dynamics model [39,40]. For the conditions of our simulation, the pressure loss across the carbonator reactor is about 100 mbar, making it necessary to use a compressor for the flue gas entering into it. In the carbonator model simulation carried out in the present work, sulfation of the solids is considered by setting a value of 99% for the SO 2 capture capacity [18,41,42]. According to the carbonator model reported in [32],CO 2 capture efficiency can be expressed as a function of the total solids inventory in the carbonator (W s ), the solids residence time in the carbonator ( s ) and the flow rate of fresh limestone makeup fed into Fig. 1. CaO conversion as a function of carbonation time for different cycles (N= 2, 10, and 20) showing the kinetic model curves proposed in [32] (solid lines) and experimental data from TGA tests reported in [34] (dotted lines) in which calcination was carried out under high CO 2 concentration. Conversion (X) and time (t) in the reaction controlled (X NK ,t K ) and diffusion controlled (X ND ,t D ) carbonation phases are indicated for cycles 20 and 2, respectively. Reproduced from [32] with permission. Fig. 2. Carbonation efficiency as a function of the solids residence time in the carbonator s , which is varied by changing the ratio of the flow rate of recirculated solids between the reactors to the CO 2 flow rate (F R /F CO2, upper horizontal axis). Calculations made for fixed values of the solids inventory W s and F 0 /F CO2 as indicated (F 0 is the flow rate of fresh limestone makeup fed into the calciner). Results from the carbonator model used in the present work (developed in [32]) and a carbonator model previously proposed by other authors [36] are shown for comparison (reproduced from [32]). C. Ortiz et al. / Applied Energy 169 (2016) 408–420 411 the calciner (F 0 ), among other variables such as the multicyclic CaO conversion used or the calcination conditions. According to carbonator model results, by using limestone derived CaO as sorbent, a CO 2 capture efficiency of about 70–90% can be attained. The almost free CO 2 gas stream that exits the carbonator is sent to a series of heat exchangers (see Fig. 4) in order to recover its sensible heat before the gas is vented into the atmosphere. As can be seen in Fig. 5, the post-combustion flue gas is preheated before entering into the carbonator reactor (FG-IN) through two heat exchangers (HE-FG and HE20) whereby heat is transferred from the CO 2 gas stream exiting the calciner previous to the storage stage (CO2-STO2) and from the calciner solids purge stream (PURGE3), respectively. The heat produced in the exothermic carbonation reaction, including the sensible heat recovered by the solids from the calciner (Q-CARB) is used in a secondary steam cycle for electricity generation. Given the large flow of solids recirculated between reactors in the CaL cycle, the energy production in the carbonator (and the energy consumption in the calciner) is rather high, which yields a power production in a secondary steam cycle similar to that produced in the reference power plant as obtained also from other previously published models [11,17]. This implies a significant penalty for the integration of the CaL cycle, albeit it could be reduced by inserting a heat exchanger between the solids leaving the calciner (F R in Fig. 6, with a temperature of about 950 °C) and the solids entering into it that must be heated up to the calcination temperature as detailed in Fig. 6. For this same purpose, Martinez et al. have proposed the insertion of alternative heat exchange systems between the calciner and the carbonator such as a cyclonic preheater [43] or a mixing seal valve [44]. An important consideration is the possibility of recarbonation of the partially carbonated CaO particles when they enter into contact with the gas/solid stream exiting the calciner. According to Valverde et al. [31], recarbonation of the solids would have a relevant effect on their multicyclic conversion behavior that should be considered. A possible solution if recarbonation is to be avoided could be the use of an indirect contact solid–solid heat exchanger, or a coupled pair of efficient gas–solid heat exchangers. In order to simplify the integration model in our work, a simple solid–solid heat exchanger is assumed with an approach temperature of 20 °C. Thus, the energy transferred to the secondary steam cycle Fig. 3. Schematic representation of the coal fired power plant (CFPP) used as reference for the integration of the CaL cycle. Table 1 Coal Pittsburgh No. 8 composition [3]. Proximate analysis %weight (as-received) %weight (dry) Ultimate analysis %weight (as-received) %weight (dry) Moisture 5.2 – Moisture 5.2 – Fixed carbon 48.1 50.7 Ash 8.6 9.1 Volatile matter 38.1 40.2 Carbon 70.2 74 Ash 8.6 9.1 Hydrogen 4.8 5.1 HHV (BTU/lb) 12.540 13.228 Nitrogen 1.5 1.6 Chlorine 0 0 Sulfur 2.2 2.3 Oxygen 7.5 7.9 Table 2 Flue gas composition of reference CFPP plant [3]. Component Flow (kmol/h) Flow (t/h) N 2 57752.53 1783.37 CO 2 10388.25 503.96 H 2 O 4955.28 98.4 O 2 2632.11 92.84 CO 473.68 14.63 NO 455.71 15.07 SO 2 126.39 8.93 H 2 41.81 0.093 NO 2 0.20 0.01 SO 3 0.07 0.06 S 0.01 0.04 Total 76826.03 2517.3 Ash 17.69 412 C. Ortiz et al. / Applied Energy 169 (2016) 408–420 is reduced, which lessens power generation from the CaL cycle and therefore allows for a smaller size of equipment and auxiliary systems. Both possible schemes (introducing or not the solids heat exchanger) are shown in Fig. 6. The sensitivity analysis that will be described in Section 4is devoted to a comparison of the energy penalty that results from both options. 3.2.2. Calciner model The partially carbonated solids at the carbonator outlet (CYC2-S in Fig. 7) are recirculated into another CFB reactor (the calciner), wherein CaO regeneration takes place through the endothermic calcination reaction at atmospheric pressure. Besides, a fresh makeup flow of limestone (MAKE-UP) that must be periodically introduced into the calciner to compensate for sorbent deactivation has to be also calcined (Fig. 7). As reported in previous works [26,45,46], for the calcination reaction to be fully achieved sufficiently fast in the calciner at atmospheric pressure under high CO 2 partial pressure, it must be operated at temperatures over 930 °C. At this temperature, which is 30–50 °C above the calcination equilibrium temperature, full calcination of the solids is achieved in short residence times, which yields a sufficiently high calciner efficiency [47]. Increasing the calcination atmosphere pressure promotes sintering of the sorbent and requires the use of higher temperatures [48], which could be detrimental for the plant performance. The imposition of a high calcination temperature makes it necessary to supply a large amount of heat to the calciner in order to achieve an acceptable calcination efficiency. This is accomplished by in-situ combustion of coal using pure O 2 (O2-IN). Oxycombustion allows retrieving a CO 2 rich stream at the reactor outlet (70–90% v/v) for subsequent compression and storage (COINSTO). As a previous step to oxy-combustion, coal decomposition (DRY-COAL) into its elemental components (IN-CALC) is introduced in the simulation model as may be seen in Fig. 7. For simulating the oxy-combustion of coal in the calciner reactor a model based on chemical and phase equilibrium through Gibbs’ free energy minimization methods is used. To obtain the necessary pure stream of O 2 for oxy-combustion in the calciner the most feasible option is to install an air separation unit (ASU). In this work, the energy penalty caused by this equipment is taken as 200 kWh per tonne of O 2 produced according to results from previous reports [42,49]. The necessary drying treatment of the gas stream exiting the calciner (before compression and storage) is not simulated in the Fig. 4. General CFPP-CaL integration scheme. Fig. 5. Schematic representation of the carbonator zone used in the simulations. C. Ortiz et al. / Applied Energy 169 (2016) 408–420 413 present work. Thus, only a one-stage compression model is simulated aimed at compressing the gas stream at the calciner outlet to a storage pressure of 100 bar. Before the compression stage, a cooling process is performed in order to improve the compression efficiency. The heat extracted from the CO 2 stream is used for preheating the flue gas entering into the carbonator (Fig. 4). The regenerated CaO particles are recovered at the exit of the calciner by a cyclone and circulated back into the carbonator for a new cycle. Energy consumption derived from solids transportation has been set at 20 MJ per tonne of solids [50]. Part of the gas stream is recirculated into the calciner (REC-CO22) in order to regulate the composition of gases in the reactor (Fig. 7). Thus, the resulting gas stream at the outlet of the calciner (CO-INSTO) has a high CO 2 content (about 70–90%v/v CO 2 ). The sensible heat of the calciner’s outlet streams (CO 2 , CaO and purge stream) is transferred to the secondary steam cycle (Fig. 6a) or to the inlet stream of CaCO 3 (Fig. 6b) through heat exchangers. Finally, a purge of solids is carried out in the calciner zone (PURGE3) to eliminate ashes and deactivated solids. 3.2.3. Secondary steam cycle The third main part in the simulation scheme is the secondary steam cycle zone (Fig. 9). This steam cycle is based on a reheat Rankine cycle with regeneration from five feedwater heaters (HEEXT1:4), one of which are of total mixer exchanger type (GASI). For this reason a series of steam extraction (EXT1:4) is realized. The steam operational parameters have been chosen from data of similar real power plants [51] and taking into account the CaL streams characteristics in order to maximize the integration efficiency without supplying more fuel for steam production. A makeup flow of water is introduced in the system (H2O-REP) to compensate for losses due to condensation in the turbine zone (COND). This water makeup flow is heated using the sensible heat of the solids purged (PURGOUT2). The steam boiler is modeled as a group of three independent zones: (i) the preheater, where the sensible heat flue gas stream (FG-OUT2) is used, (ii) the evaporator, where the heat Fig. 6. Schematic CaL cycle without (a) and with (b) solids heat exchanger between the calciner and carbonator. Fig. 7. Schematic representation of the calciner zone used in the simulations. 414 C. Ortiz et al. / Applied Energy 169 (2016) 408–420 produced by the exothermic carbonation reaction (Q-CARB) is employed for fluid evaporation through a vapor stream from the carbonator (STCARB2), and (iii) the super-heater, where a part of sensible heat of the CO 2 exiting the calciner (CO-INSTO) is used. As can be seen in Figs. 4 and 9, the stream of CO 2 exiting the calciner (at 950 °C) is passed through a series of heat exchangers, being the final stage previous to storage the steam re-heating between turbines facilities (HE-INT). Fig. 8 shows the main steam heating process to reach the live steam temperature. The condenser (HE5) is modeled as a heat exchanger using water as cold stream. Table 3 shows several input parameter of the steam secondary cycle zone. 4. Simulation results The model described in Section 3and using the solids heat exchanger between the calciner and the carbonator (Fig. 6a) has been simulated in order to determine the new global efficiency and penalty arising from the CaL cycle integration in the reference coal power plant. The values of the operation parameters used for this purpose are summarized in Table 3. In order to analyze the effect of integrating the CaL process in the reference plant it is necessary to quantify the cycle efficiency as affected by the new cycle penalty in order to assess whether retrofitting the plant with the CaL process is industrially feasible and competitive against other possible CO 2 capture methods. As specified above, the reference plant efficiency is assumed to be of 33.5%. The new efficiency taking into account the CaL integration is calculated by the following expression (1): g plant ¼ _ W CFPP þ _ W gen;CaL  _ W cons;CaL _ Q CFPP þ _ Q CaL ð1Þ where g plant is the new global efficiency value after integration, _ W CFPP is the reference plant generated power, _ W gen;CaL is the gener- Fig. 8. T–Qdiagram of the main steam heating process to reach the live steam temperature. Fig. 9. Schematic representation of the secondary steam cycle zone used in the simulations. C. Ortiz et al. / Applied Energy 169 (2016) 408–420 415 ated power in the CaL process (secondary steam cycle), _ W cons;CaL is the power consumption in the CaL cycle, _ Q CFPP is the reference plant heat consumption and _ Q CaL is the heat consumption in the calciner for oxy-combustion. The power consumption in the CaL cycle can be expressed according Eq. (2): _ W cons;CaL ¼ _ W FGCOMP þ _ W CO2COMP þ _ W COM3 þ _ W ASU þ _ W solid ð2Þ where _ W CO2COMP ,_ W COM3 and _ W FGCOMP are the power consumed in the CO 2 compressor for storage, CO 2 recycle stream compressor and flue gas compressor, respectively. _ W ASU is the power consumed in the ASU and _ W s olid is the power consumption needed for the transport of solids. Cycle integration and streams simulation results are summarized in Tables 4 and 5, respectively. 4.1. Sensitivity analysis Parametric tests have been carried out in order to evaluate the CaL integration potential in a real power plant. Percentage of penalty points was calculated over the plant efficiency. Firstly, simulation tests were carried out under certain particular conditions imposed on the secondary steam cycle (Table 3). In order to analyze how these conditions affect the efficiency penalty, an analysis changing the cycle conditions was performed by taking into account a 10% relative variation on the values of several parameters such as the live steam pressure (P VV ) and live steam temperature (T VV ) in the secondary steam cycle, the extraction flow, the rate of consumption of oxygen and coal for oxycombustion in the calciner ( _ W ASU and _ W solid , respectively), and the pumps and compression efficiencies. As can be seen in Fig. 10, a relative variation of these parameters within a 10% yields a maximum change on the order of 1% in the efficiency penalty. In regards to P VV and T VV , and the compressor efficiencies their increase would yield a remarkable decrease of the efficiency penalty. However, operation at increased P VV and T VV involves a higher equipment cost, making it necessary to carry out an additional sensitivity analysis from the economic standpoint. A variation of the extraction flow and pumps efficiencies does not have a critical effect on the energy penalty. On the other hand, the energy penalty is notably increased by the increase of the flow rate of coal and oxygen used for oxycombustion to raise the calciner temperature as might be expected. Next a test is performed to analyze the effect of increasing the solids residence time in the carbonator s by decreasing the recirculation flow rate of solids between the calciner and the carbonator (F R ). As may be seen in Fig. 11, the efficiency penalty is notably decreased as the solids residence time in the carbonator is increased. Moreover, the lower the flow rate of fresh limestone makeup the larger is the reduction of efficiency penalty. Thus, it can be concluded that a lower flow of solids recirculation between the reactors would lead to a remarkable improvement of the integrated cycle efficiency. This is a novel relevant result to be highlighted that arises from the new carbonator model in which the importance of carbonation of the CaO solids in the solid-state diffusion controlled phase is considered. An interesting parameter to assess the tradeoff between the achieved CO 2 capture efficiency and energy penalty is the specific energy consumption per kg of CO 2 captured (SPECCA). Thus, SPECCA represents the additional fuel consumption (in MJ) necessary to avoid the emission of 1 kg of CO 2 into the atmosphere [11,52,53]. It is calculated according to expression 3: SPECCA ¼3600 1 g plant  1 g ref E ref Eð3Þ where g ref ;g plant are the power plant efficiency without and with the CaL cycle integrated, and E ref ;Eare the emissions ratio (in kg CO 2 / kWh e ) before and after CaL cycle integration, respectively. Fig. 12 shows the results of the values obtained for SPECCA in our work to obtain a CO 2 capture efficiency of 80% and 90%, respectively, as depending on the solids recirculation flow rate. As can be seen, for a given value of the solids recirculation flow rate (F R /F 0 fixed), it is necessary to increase the amount of energy required in order to obtain a higher capture efficiency since the makeup flow rate of fresh limestone introduced into the system F 0 must be increased. Furthermore it is observed that a lower SPECCA is attained as the solids recirculation flow rate if decreased for a given capture efficiency (as was seen in Fig. 9). Thus, the additional energy consumption necessary to achieve a given capture efficiency is lowered as F R /F 0 is diminished. Fig. 13 shows the effect of changing the CO 2 compression pressure to storage on the steam cycle efficiency and efficiency penalty. Table 3 Main simulation model data. Equipment/zone Parameter Value F CO2 10388 kmol/h F 0 =F CO2 0.05 s 313 s F R =F CO2 10 W s 200 tonne Heat exchangers Approach temperature 10 °C Compressors, turbines Isentropic efficiency 0.85 Pumps Efficiency 0.9 Cyclones Efficiency 0.99 Type Barth 1 Correlation Muschelkrautz Reactor/calciner T950 °C P1 atm _ m coal 137 tonne/h P CO2 0.83 X calciner 1 Reactor/carbonator T650 °C P1 atm E CO2 0.77 ASU _ W ASU 200 kWh/tonne O 2 _ W solid 5.5 kWh/tonne solids Table 4 Main cycle results. Equipment/zone Parameter Value Reactor/calciner _ Q cal 1150.1 MW COM3/calciner Power 9.5 MW HEO2/calciner Heat duty 15.6 MW Reactor/carbonator _ Q carb 711 MW E CO2 0.77 FG-comp/carbonator Power 5.5 MW HE-FG/carbonator Heat duty 117.5 MW CO 2 -COM2/CO 2 storage Power 181 MW Pressure 100 bar Secondary steam cycle _ W tur;1 240.6 MW DP tur;1 128 bar _ W tur;2 284.6 MW DP tur;2 11.9 bar _ W B1 1.45 MW _ W B2 0.241 _ W B3 4.86 g(secondary steam cycle) 0.42 CaL cycle g(without CO 2 comp) 0.43 g(with CO 2 comp) 0.28 CFPP + CaL cycle (integration) g0.28 Integration penalty 5.60% points 416 C. Ortiz et al. / Applied Energy 169 (2016) 408–420 As may be observed, the final value of the CO 2 pressure has a significant role on the integration penalty. A 1% increase of the efficiency penalty is obtained by raising the storage CO 2 pressure from 50 bars to 140 bars. On the other hand, the steam cycle efficiency is enhanced by increasing the storage CO 2 pressure as seen in Fig. 13 since the sensible heat of the CO 2 stream to be compressed is used for steam re-heating between the high and low pressure turbines, which partially mitigates the increase of energy penalty due to higher compression work. Arguably, the efficiency of integration would be enhanced by using a more efficient compression process for the CO 2 produced, such as a multi-stage optimized process, or using the high temperature and pressure CO 2 stream in a secondary cycle, which would allow compression needs to be reduced. In this regard, Fig. 14 shows the CFPP-CaL efficiency penalty by taking into account an intercooled CO 2 multistage compression. In the present integration model, a high temperature of CO 2 compressed stream is needed since a part of sensible heat from the CO 2 stream is used for re-heating the steam flow. For this reason, the process is simulated by taking into account a constant inter-stage temperature of 170 °C and avoiding cooling in the last stage, which is followed by heating to the temperature required for the exchange with the steam cycle (around 780 °C). As seen from this analysis, by increasing the stages number it is possible to achieve a reduction of energy penalty of around 0.5%. Recent studies [54,55] have been devoted to a more detailed and optimized heat exchanger analysis aimed at achieving the best heat integration solution. Applying some of the methods proposed in these analyses, an even smaller energy penalty could be attained by operating under a higher solids residence time in the carbonator according to Figs. 11 and 12. A desirable option to reduce the integration efficiency penalty would be to decrease the calciner reactor temperature with a consequent reduction in fuel and O 2 requirements for oxycombustion. However, as mentioned earlier, fast calcination of the fresh Table 5 Main streams results. Streams Flow rate (kg/s) Temperature (°C) Pressure (bar) Main composition (mole fraction) Acronyms Definition CARB-IN Solids entering into the carbonator 1708.32 950 1.01 CaO (0.98) FG-IN Flue gas entering into the carbonator 634.35 384.5 1.10 N 2 (0.75) CO 2 (0.14) CYC2-S Solid stream exiting the carbonator 1807.29 650 1.01 CaO (0.90) CaCO 3 (0.08) CYC2-G Gas stream exiting the carbonator 534.21 650 1.01 N 2 (0.84) CO 2 (0.03) PURGE4 Purge stream flow exiting the system 170.78 88 1.01 CaO (0.98) CaSO 4 (0.02) a MAKE-UP Makeup flow of fresh limestone 170.78 25 1,01 CaCO 3 (1.00) O2-IN O 2 stream entering into the calciner 87.61 640 1.10 O 2 (1.00) CYC1-G Gas stream exiting the calciner 584.01 950 1.01 CO 2 (0.83) CYC1-S Solid stream exiting the calciner 1879.03 950 1.01 CaO (0.98) REC-CO22 CO 2 stream recirculated into the calciner 350.40 970 1.10 CO 2 (0.83) CO-INSTO CO 2 stream exiting the calciner zone 233.60 950 1.01 CaO (0.98) VAP3 Live steam 429.53 560 140.00 H 2 O (1.00) INTUR2 Reheated steam 340.19 472 17.00 H 2 O (1.00) VAP5 Water flow exiting the condenser 293.38 46 0.10 H 2 O (1.00) VAP-IN Water stream entering into steam boiler 429.53 251 140 H 2 O (1.00) H2O-1 Water stream entering into the condenser 305.50 25 1.00 H 2 O (1.00) H2O-REP Water makeup flow 15.08 25 1.00 H 2 O (1.00) a Does not take into account ash flow (33.06 kg/s). Fig. 10. Efficiency penalty as affected by the variation of several model parameters over the values shown in Table 3. Simulations made using the integration model described in Section 4. C. Ortiz et al. / Applied Energy 169 (2016) 408–420 417 makeup of limestone under high CO 2 partial pressure only takes place at temperatures over 930 °C. An alternative approach to mitigate the energy penalty would be thus to use a natural CaO precursor showing fast calcination at reduced temperatures as compared to limestone. This can be the case of natural dolomite as suggested by lab-scale TGA tests [3,6]. Let us finally assess the effect of introducing a solids heat exchanger between the calciner and carbonator reactors (Fig. 6). As can be seen in Table 6, the simulation including the heat exchanger yields a significant reduction of coal consumption and the thermal power produced by the CaL cycle (using an approach temperature of 20 °C), which implies a reduction of the efficiency penalty in the global system. Besides, the power production in the steam cycle is decreased and likewise the equipment size and investment cost. Data on the efficiency penalty and coal consumption in the calciner are shown in Fig. 15 as a function of the approach temperature in the solid heat exchanger. As might be expected, both penalty and coal consumption are decreased as the approach temperature is decreased. Fig. 11. Efficiency penalty for the integration of the CaL technology in a coal combustion power plant as a function of the solids residence time in the carbonator, which is varied by changing the ratio of solids recirculation to CO 2 flow rates (F R /F CO2 ). Calculations were made for several values of the ratio of limestone makeup flow to CO 2 flow rates (F 0 /F CO2 ). Fig. 12. SPECCA value as a function of the ratio of solids recirculation flow rate to CO 2 flow rate (F R /F CO2 ). Calculations made from simulations of the model described in Section 4by varying the ratio of fresh limestone makeup flow rate to CO 2 flow rate (F 0 /F CO2 ) to get a fixed CO 2 capture efficiency (E CO2 ) value of 80% (left) and 90% (right). Fig. 13. Steam cycle efficiency (left axis) and integration efficiency penalty (right axis) as a function of CO 2 storage pressure. Fig. 14. Efficiency penalty (left axis) and compression power consumption (right axis) as a function of stages number in the CO 2 to storage multi-stage compression. Table 6 Simulation results for the reference case with/without solid heat exchanger. Parameter Value Model inputs F CO2 10388 kmol/h F 0 =F CO2 0.05 s 313 s F R =F CO2 10 W s 200 tonne Heat solid exchangers (between calciner–carbonator) Approach temperature 20 °C Parameter Without heat exchanger (reference case) With heat exchanger Model outputs _ Q calciner 1150.1 MW 580.5 MW _ Q carb 711 MW 218 MW _ W steam cycle 525.2 MW 291.4 MW Coal consumption 137 tonne/h 69 tonne/h Integration penalty 5.6% points 5.35% points 418 C. Ortiz et al. / Applied Energy 169 (2016) 408–420 5. Conclusions In this work we have analyzed the energy penalty arising from integrating the CaL technology in a coal fired power plant for CO 2 capture. Simulation results show that the energy penalty ranges between 4% and 7% points over the reference plant efficiency (taken as 33.5%). A main novel feature of the integration model proposed is that it uses a carbonator model that considers the critical effect of solids residence time in the carbonator on the CO 2 capture efficiency, as inferred from lab-scale experimental results when the solids are regenerated by calcination under high CO 2 partial pressure in accordance with practical conditions. Thus, the integration energy penalty is significantly reduced as the solids residence time in the carbonator is increased within the practical operation range of a few minutes. The main reason for this result is the lower heat requirement in the calciner reactor as the recirculation flow rate of solids between the calciner and carbonator reactors is decreased. In addition, by increasing the solids residence time, a smaller size of process equipment would be necessary thus reducing investment costs as well as O&M cost. A further conclusion from the analysis is that optimizing the heat exchangers network and the power system parameters would significantly improve energy penalty results. Nevertheless, the large amount of heat required in the calciner leads to the production of a large amount of energy in a secondary power cycle, which results in a system of size similar to the actual reference power plant size (without the CaL cycle) as obtained in previous works, thus hindering investments prospective. Accordingly, it would be highly convenient to optimize the heat integration of the streams exiting the CaL with the reference power plant flows. Our work suggests that a notably lower energy penalty could be achieved together with a reduction of power production needs by adding s solids heat exchanger system between the calciner and carbonator reactors. To summarize, operation conditions that would essentially improve the industrial competitiveness of the CaL-CFPP integration would consist of prolonging the residence time in the carbonator reactor and optimizing the integration of heat produced. In a future work an economic feasibility study must be carried out to take into account the reduction of process equipment size possible when the CaL cycle is operated with a lower solids recirculation flow rate as suggested in the present work. Acknowledgements This work was supported by the Andalusian Regional Government Junta de Andalucia (contract FQM-5735), and by the Spanish Government Agency Ministerio de Economia y Competitividad (contracts CTQ2014-52763-C2-2-R and CTQ2014-52763-C2-1-R). References [1] Hanak DP, Biliyok C, Yeung H, Białecki R. Heat integration and exergy analysis for a supercritical high-ash coal-fired power plant integrated with a postcombustion carbon capture process. 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Design and analysis of heat exchanger networks for integrated Ca-looping systems. Appl Energy 2013;111:690–700. http://dx.doi.org/10.1016/j.apenergy.2013.05.044. [55] Lara Y, Lisbona P, Martínez A, Romeo LM. A systematic approach for high temperature looping cycles integration. Fuel 2014;127:4–12. http://dx.doi. org/10.1016/j.fuel.2013.09.062. 420 C. Ortiz et al. / Applied Energy 169 (2016) 408–420 F CaCO 3 ;clc  D t sun ¼F CaCO 3 ;crb 24 ð3Þ If, for instance, we consider a daytime D t sun of 8 h (clear sky), the ratio between the circulating streams in the calciner and the carbonator will be of 3 over the 24 h, while in case that the daytime is 12 h, the molar flow rates in the calciner side will be twice those in the carbonator side. A more sophisticated model (dynamic model) would be necessary in a framework of long-period control taking into account real data for solar energy input in order to obtain real energy production values along the year. Nevertheless, the scope of this paper is just to estimate the cycle efficiency for a fixed irradiation. A variation in the solar input power (as a function of solar irradiance) modifies mass and energy flows in the cycle, which requires improving the control strategy to maximize the cycle performance, including part-load strategies for the closed Brayton cycle. 2.1.2. Energy balance The first thermodynamics law applied to a reacting system is used in both the carbonator and calciner reactors for the energy balance: X i F i;out h i;out X i F i;in h i;in ¼U_ Wð4Þ F i;out F i;in ¼n m i ð5Þ where ndenotes the extent of reaction per unit time. Arranging and considering that output conditions are reactor conditions, it is: n D H R ðT react ÞþX i F i;in ðh i;react h i;in Þ¼U_ Wð6Þ with D H R ðT react Þ¼X i m i h i;T ¼ D H 0 R þX i m i Z T react ref c p;i dT ð7Þ being the reaction enthalpy change at the reaction temperature. Energy change in the control volume consists therefore of the heat of reaction at the reactor temperature n D H R ðT react Þand the heat required to bring the reactants from inlet to reactor’s conditions P i F i;in ðh i;react h i;in Þ. Eq. (4), when applied to the carbonator, serves to balance out the amount of CO 2 (F CO 2;S in Fig. 3) needed to remove all the heat that is not absorbed by reactants or dispersed through the walls assuming isothermal conditions at the reactor. On the other hand, Eq. (4), when applied to the calciner, balances out the CaCO 3 production in accordance with the net energy input of the system, which is the solar heat supply minus the heat losses occurring between the solar receiver and the reactor due to conduction, thermal radiation, absorption, etc. 3. CSP–CaL energy storage cycle analysis The proposed cycle configuration is the result of an optimization analysis based on the pinch-point method [60,61]. Thus, the heat exchanger network seeks for optimal internal heat-recovery performance in a broad range of operational conditions. In this section, the influence of the main cycle parameters are analyzed in order to determine the most efficient operating conditions. Table 2 shows several parameters used in the cycle simulations whose influence is discussed below. The proposed plant (Fig. 1) is equipped with a solid–solid heat exchanger HXA, two indirect gas–solid heat exchangers (HXB and HXF), a gas–gas regenerator (HXG) and two cyclone gas–solid direct heat exchangers (HXE and HXI). The results of the model for the proposed plant configuration will be reported as a function of the CaO conversion X, pressure ratio PR in the CO 2 turbine T1 (see Fig. 1), absolute carbonator pressure ðPÞ, T1 turbine outlet pressure ðTOPÞ, and carbonator temperature ðTÞ, which have been identified as the key operational parameters. 3.1. CaO conversion Conversion (X) of the CaO solids population recycled in the loop plays a relevant role in the plant performance. Changes in the value of this parameter cause remarkable modifications in the necessary flow rates of the circulating solids, CO 2 generation and CaCO 3 consumption in the calciner, heat balance in the calciner and useful carbonation heat and storage size, which have a significant effect on the plant efficiency. Fig. 3. Mass-balance schematics of the plant showing the flow rates of solids and CO 2 streams between the carbonator and calciner. Fig. 4. CO 2 storage conditions by considering a 12 h daytime (100 MW th CSP plant). 594 R. Chacartegui et al. / Applied Energy 173 (2016) 589–605 The results of the model simulations carried out demonstrate that the power balance in the calciner side is favored by high values of Xthat allow decreasing the amounts of circulating solids, thus yielding lower heat-losses in the calciner side exchangers network and more heat available for calcination. CaCO 3 decomposition and CO 2 generation in the calciner are consequently enhanced and more heat is available in the discharge phase for carbonation. Carbonation heat is proportional to the amount of CaO and CO 2 produced in the calciner and is partially absorbed by reactants and partially removed by the CO 2 stream for power generation. For values of CaO conversion below a certain lower bound, the heat removed by inactive solids overcomes the heat of carbonation. In this situation, isothermal operation of the carbonator becomes compromised unless heat is provided by an external source. This critical value of conversion, X lim , is mainly dependent on the temperature of the inlet CaO, which is in turn imposed by the temperature at main turbine output through HXF. This means that, in practice, the ratio of carbonator pressure to the T1 turbine outlet pressure (PR) is the parameter that mainly determines the minimum value of CaO conversion needed for operation as will be seen ahead. As can be observed in Fig. 5, the higher CaO conversion the higher fraction of heat can be removed by the power fluid (CO 2 in the present case) given the lower flow rate of solids cycled within the system, which reduces the energy required in the form of sensible heat for the solids to reach the carbonation temperature. In the current cycle configuration, the presence of exchangers HXG and HXI ensure optimal thermal recovery and consequently the CO 2 mass flow rate in the cycle is perfectly proportional to the carbonation heat (as may be seen in Fig. 5). Fig. 6a shows the thermal capacity rate (defined as the product of the molar flow rate by the specific heat of the fluid of interest Fc p ) of the carbonator-side streams as a function of the CaO conversion. The cooling (CO 2 c) and preheating (CO 2 p) CO 2 streams’ split ratios (see Fig. 1) are determined by means of a pinch-point optimization method. When the amount of solids is reduced, a larger fraction of heat is regenerated in the heat exchanger HXG. This results in an increased size of the gas–gas regenerator HXG, unlike for HXF and HXI. High values of CO 2 molar flow rates for low pressure ratios PR can push up too much the size of the exchanger HXG. For this reason, too low PR values should be avoided. The efficiency curve is represented in Fig. 6b as a function of CaO conversion Xfor a fixed pressure ratio PR = 3.2 used as an optimum value. Efficiencies are shown for the cases of a carbonator operating under atmospheric pressure (expansion to sub atmospheric pressure) and a pressurized carbonator (atmospheric turbine output). As may be seen, efficiency increases with CaO conversion and reaches a maximum between 44% and 46% under the conditions considered. Importantly, the energy consumption related to solids conveying diminishes the efficiency especially in the case of low values of CaO conversion. Table 2 Values of the main parameters used in our work to simulate the proposed CSP–CaL integration model. Net absorbed solar flux in calciner 100 MWt Daylight hours (constant solar flux) 12 Thermal dispersions in carbonator 10% Calciner temperature 900 °C Carbonator temperature 875 °C Carbonator pressure 3.2 bar Turbine outlet pressure 1 bar Ambient temperature 20 °C Minimum temperature difference solid–solid HX 20 °C Minimum temperature difference solid–gas HX 15 °C Minimum temperature difference gas–gas HX 15 °C Minimum temperature difference CO 2 cooler 10 °C Intercoolings in CO 2 storage compression 5 Intercoolings in power cycle compression 2 Pressure losses in CO 2 circuit 0 bar CO 2 storage conditions 75 bar, Tambient Solid phase conveying energy consumption 10 MJ/tonne/100 m Equivalent length solids conveying carbonator side 100 m Equivalent length solids conveying calciner side 100 m Isentropic efficiencies (compression/expansion) 0.89 Fig. 5. Available heat of carbonation (left axis) and CO 2 mass flow rate through the carbonator (right axis) vs CaO conversion. Fig. 6. (a) Thermal capacity of streams (CaO, CaCO 3 , and CO 2 ) in the carbonator side heat exchangers (HXE, HXF, HXG, HXI) vs CaO conversion for a pressure ratio PR = 3.2. (b) Efficiency vs CaO conversion (using PR = 3.2) for selected absolute pressures of the carbonator and turbine outlet, respectively, as indicated. R. Chacartegui et al. / Applied Energy 173 (2016) 589–605 595 3.2. Pressure The carbonator pressure determines the upper pressure of the Brayton cycle and therefore plays a crucial role in the plant performance. The pressure ratio in the power turbine (T1) influences carbonator side’s parameters, mainly temperatures (first of all turbine outlet temperature), flow rates, exchangers’ nominal power (and size) and operational parameters. Some system variables show a direct dependence on the pressure ratio PR while others depend separately on the absolute values of the carbonator pressure and turbine outlet pressure. High pressure ratios yield a reduced temperature at the turbine outlet as seen in Fig. 7 where the effect on the temperature of the sorbent (CaO) at the carbonator inlet can be observed too. Note that the optimized heat exchangers configuration adopted in the model yields high and stable values of the CO 2 temperature at the carbonator inlet regardless of the pressure ratio. The trends of CaO and CO 2 temperatures at the carbonator inlet, both approaching the carbonator temperature in the limit PR ?1, yield an infinite CO 2 flow rate for removal of the carbonation heat. At higher pressure ratios, CO 2 inlet temperature is quite stable while CaO temperature decreases, absorbing more heat in the reactor and causing the CO 2 flow rate to drop as it is seen in Fig. 9b (green curve) 1 . For this reason, high pressure ratios are desirable in order to reduce costs related to the limited inventory storage capacity and to the plant machinery size. On the other hand, reduced values of PR allow for less strict operational limits regarding the critical value for CaO conversion below which no isothermal carbonator operation is possible. Fig. 8 shows the minimum CaO conversion required X lim in the proposed power cycle as a function of the pressure ratio PR. The trend is again justified by the lower CaO temperatures at the carbonator inlet for increasing values of PR. A detailed scheme of the CO 2 reaction turbine T2–CO 2 cycle compression C1 group is shown in Fig. 9a. The specific work of compression increases with the pressure ratio as seen in Fig. 9b. At a fixed pressure ratio, lower absolute pressures are associated with higher specific compression work, thus giving rise to an undesirable increase of compression power. This effect is shown in the blue curve of Fig. 10a where the power (calculated as the product of the CO 2 flow rate, following the trend shown in the green curve of Fig. 9b, by the specific work of compression, blue curve) is plotted vs pressure ratio for different values of turbine outlet pressure (shaded scale). As may be seen in Fig. 10a, lower carbonator absolute pressures (and TOP) imply a higher expansion ratio in the turbine T2 (from 75 bar to carbonator pressure) and a proportional increase of power generation (see the shaded scaling around the green curve in Fig. 10a). Thus, smaller absolute pressures determine higher compression power but also higher expansion power. The latter beneficial effect prevails over the former detrimental one, thus the net power consumption of the group is decreased in practice by low absolute pressure values. As can be seen in Fig. 10b, the gross power generation in the turbine depends only weakly on the absolute pressures. Hence, better performances are retrieved for lower absolute pressures at a fixed pressure ratio. This explains the global plant efficiency trend with pressures, which is shown in detail in Figs. 11 and 12 (note that the energy consumption in compressor C2 does not depend on PR). Fig. 11 demonstrates that, as usually occurs for regenerative Brayton power cycles, there is a maximum in the efficiency-PR curve, occurring in the present case at 3.2 bar. A more detailed representation of the global efficiency achievable, as depending on the carbonator pressure Pand turbine outlet pressure TOP, is given in the contour plot of Fig. 12 where a constant value of CaO conversion X= 0.5 has been assumed. Fig. 7. Variation of temperatures in the regions encircled in the diagram (according to the colors code) with pressure ratio. (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.) Fig. 8. Minimum value of CaO conversion for carbonator isothermal operation vs pressure ratio. 1 For interpretation of color in Fig. 9, the reader is referred to the web version of this article. 596 R. Chacartegui et al. / Applied Energy 173 (2016) 589–605 The effect of gas pressure drops in the CO 2 power cycle has been also investigated since viscous frictional forces can be enhanced by the use of fluidized bed reactors and open cyclone exchangers. Due to pressure drops, the power consumption of the plant increases and the efficiency decreases. At the same time, the maximum of efficiency is shifted to higher pressure ratios albeit the location of the maximum becomes less marked as the pressure drop is increased (see Fig. 13). The weighted dotted line joins all the maximums for different values of pressure drops in the power cycle (expressed as percentage of the absolute carbonation pressure). 3.3. Temperature In a Brayton cycle, the increase of turbine inlet temperature (which is, in the present case, the carbonator temperature T) naturally enhances the efficiency. It is therefore crucial to envisage which limits are imposed to the maximum carbonator temperature. A main limitation is linked to the equilibrium of the calcination/carbonation reaction. At a given temperature the reaction reaches equilibrium for a CO 2 pressure below which carbonation cannot take place. Equivalently, at a given CO 2 pressure, there is a maximum temperature above which carbonation is not possible. For example, under a pure CO 2 environmentatatmospheric pressure, equilibrium is reached at 895 °C. Carbonation slows down as practical operation conditions approach equilibrium [62]. Thus, the closer to equilibrium conditions the smaller will be the quantity of CO 2 reacting in the carbonator in practice (Fig. 14a). The equilibrium fraction of CO 2 in the carbonator (molar fraction of CO 2 at maximum capture efficiency) is defined as [63]: y eq ¼P eq P¼1 P4:137 10 7 exp 20474 T  ð8Þ where P eq (bar) is the CO 2 pressure at equilibrium. According to Eq. (8), at a fixed carbonator pressure P, an increase of the carbonator temperature yields an increase of the equilibrium pressure of CO 2 , and thus of y eq . At low temperatures, the CO 2 equilibrium pressure is small, thus the equilibrium molar fraction (y eq ) achievable is also very close to 0 and (almost) the entire stoichiometric amount of CO 2 can react even though the reaction kinetics will be hindered at very low temperatures [64]. On the other hand, if the temperature is increased, the equilibrium molar fraction of CO 2 achievable at the output approaches the CO 2 molar fraction at the input since the CO 2 pressure in the inlet gas stream becomes closer to the equilibrium pressure ðP CO 2 ;carb;in !P eq Þ. When the combination of temperature and CO 2 pressure yields an equilibrium molar fraction of CO 2 equal to the inlet molar fraction (y CO 2 ;carb;in ) carbonation is not possible anymore. In the present application, the value of y CO 2 ;carb;in has been maximized by working with pure CO 2 in the carbonator, with the consequent relevant improvements in reaction kinetics and reactor efficiency. The equilibrium CO 2 reaction efficiency E eq , defined as: E eq ¼F CO 2 ;in F CO 2 ;eq F CO 2 ;in ð9Þ represents the ratio of the CO 2 flow rate that reacts in the carbonator to the total CO 2 flow rate entering into it as limited by the reaction equilibrium. By increasing the carbonator temperature, the equilibrium CO 2 reaction efficiency is decreased (black curve of Fig. 14a). Thus, more CO 2 must be circulated in the loop in order to maintain the desired rate of thermal energy generation in the carbonator (Fig. 14b). It is therefore important not to work too close to the equilibrium limit in order to keep CO 2 flow rates at a feasible value. Moreover, the reaction kinetics would be slowed down as equilibrium is approached [64]. Note that the above means that higher carbonator pressures may yield an increase of efficiency through a carbonation temperature upgrade, as shown in the dotted curves of Fig. 14a. 3.4. Combined effect of temperature, pressure ratio and carbonator pressure The search for an optimum operation point that might result from a specific combination of temperature, carbonator pressure and pressure ratio is not straightforward. On one side, despite low absolute pressures are beneficial for increasing the efficiency at a fixed carbonator temperature, high pressures allow for further temperature upgrading, which yields higher efficiency. This is shown in Fig. 15 where the red curve, corresponding to low absolute pressures (turbine outlet pressure of 1 bar) shows a maximum efficiency of 42.6% at 875 °C for PR = 3.2. If absolute pressures are increased, operation at the optimum pressure ratio may be possible at an increased temperature of 950 °C, thus increasing the efficiency through a temperature upgrade despite the higher absolute pressure. Note that in Fig. 15 the CO 2 flow rate is calculated by taking into account the real equilibrium limitations (as shown in Fig. 14b). The dashed part of the curves in Fig. 15 represent a constrained operation range on the circulating rates set by an equilibrium molar fraction of CO 2 in the carbonator exceeding 0.3. Despite the maximum in the efficiency curve occurs at pressure ratios around 3, the molar flow rates (and the size of all exchangers, in particular the exchanger HXG) can be considerably reduced through an increase of the pressure ratio without efficiency penalizations. If the carbonator pressure is increased higher carbonation temperatures are in fact achievable and therefore the same efficiency can be reached through optimizing costs as shown in Fig. 16. Summarizing, the analysis carried out yields the following main conclusions: Fig. 9. (a) CO 2 reaction turbine–CO 2 cycle compression group. (b) Left axis: Specific work of compression vs pressure ratio PR for several turbine outlet pressures (TOP). TOP = 0.1 bar (dashed line) and 1.3 bar (solid line). Right axis: CO 2 mass flow rate vs PR. R. Chacartegui et al. / Applied Energy 173 (2016) 589–605 597 – Thermal optimization of the power cycle increase the global plant efficiency and invert the negative trend of the efficiency with increasing CaO average conversion in the carbonator (X) that was inferred from a previous work [17]. The optimum configuration with exchangers HXE, HXF, HXG and HXI yield high integration efficiencies. – Achieving high values of CaO conversion Xis beneficial for every aspect: higher efficiency, amount of passive solids reduction, lower size of silos and storage, lower size of exchangers and lower fresh sorbent makeup flow (in the case it was needed). Nevertheless, the nominal power of the regenerator HXG increases with Xand this can lead to a disproportionate increase of cost unless pressure ratio is increased. – The optimal pressure ratio value is found at low values (PR around 3.2) although several issues must be remarked: o Pressure ratios below the maximum must be avoided because they lead to a notable efficiency drop. o Pressure ratios above the maximum only generate a slow decrease in efficiency. o The presence of frictional pressure losses in the CO 2 circuit decreases the maximum efficiency and shifts the maximum to higher PR (around PR = 4.8 when pressure losses are 20% of the carbonator pressure P,PR = 7.1 when they are 50%P). Besides, it stabilizes the efficiency value for a wide range of pressure ratios (e.g. if 20%Ppressure losses occur, the optimum point is at PR = 5 but the efficiency variation is below 1% for PR in the range 2.8–10). Fig. 10. (a) Power balance in the T2/C1 group. Net power consumption (carbonator side) vs pressure ratio for different values of turbine outlet pressure (shaded scale, where arrows direction represents higher TOP values). Solid curves refer to atmospheric turbine output pressure. (b) Power balance in the carbonator side. Net power consumption, power generation in the turbine and net power production vs pressure ratio for different values of turbine outlet pressure are shown. A constant value of CaO conversion X= 0.5 is used. 598 R. Chacartegui et al. / Applied Energy 173 (2016) 589–605 o The CO 2 flow rate needed in the power cycle reaches very high values when pressure ratios are small (At X=0.5 it should be >300 kg/s for PR < 2; >200 kg/s for PR < 2.8; <100 kg/s for PR > 10). o Consequently, the nominal power (thus size and cost) of all heat exchangers can be reduced by increasing the pressure ratio. This concerns in particular the regenerator HXG (the most affected one by the CO 2 flow rate) whose size is reduced by 3–4 times when PR is increased from PR =3to PR =9. – Despite the (although small) efficiency drop, high pressure ratios are therefore desirable to reduce the circulating CO 2 flow rate and the size of exchanger HXG, in particular if the CaO solids show a high value of conversion X. The unavoidable Fig. 11. Plant efficiency as a function of carbonator to turbine outlet pressures ratio (for a fixed CaO conversion X= 0.5) and turbine outlet pressure TOP (shaded scaling). The solid line corresponds to TOP = 1 bar. Fig. 12. Contour plot of efficiency (assuming a constant value of CaO conversion X= 0.5) as depending on carbonator and turbine outlet pressure. Black lines indicate a constant value of pressure ratio. R. Chacartegui et al. / Applied Energy 173 (2016) 589–605 599 presence of gas pressure losses makes this choice less penalizing from the efficiency point of view. – The role of absolute carbonator and turbine outlet pressure on the plant efficiency is significant: o Lower absolute pressures yield higher performances at fixed carbonator temperature thanks to the higher CO 2 turbine T2 expansion ratio. o Nevertheless, high carbonator pressures ensure from equilibrium considerations higher temperature achievable in the carbonation reaction with the consequent efficiency upgrade. o At same time, at low pressures the specific volume of CO 2 may lead to an excessive increase of the size of the gas turbine and pipelines (especially for turbine outlet below atmospheric pressure). 3.5. Limestone derived CaO conversion at CSP–CaL integration conditions The combined effects of pressure ratio, carbonator pressure and temperature must be carefully re-evaluated once the value of residual CaO conversion achievable under the specific CaL Fig. 13. Influence of pressure drops of CO 2 in the circuit on efficiency as a function of pressure ratio PR. Maximum efficiency is decreased and the optimum pressure ratio is increased. Turbine outlet pressure = 1 bar, CaO conversion X= 0.5. Fig. 14. (a) Left axis: Carbonator capture efficiency vs temperature in the case of fully achieved carbonation (blue curve) and equilibrium capture efficiency (black curve) for carbonator pressures of 3.2 (solid line) and 8 (dotted line) bar (fixed pressure ratio PR = 3.2). Right axis: equilibrium and inlet molar fraction of CO 2 vs temperature. (b) CO 2 molar flow rate through the carbonator vs carbonation temperature. Note that pure CO 2 is used in the carbonator as working fluid. Fig. 15. Efficiency as a function of PR for different values of carbonation temperature and turbine outlet pressure. Red curve: turbine outlet pressure = 1 bar, carbonation temperature = 875 °C; operation point (max efficiency): carbonator pressure = 3.2 bar (PR = 3.2). Purple curve: turbine outlet pressure = 4 bar, carbonation temperature = 950 °C; operation point: carbonator pressure = 14.8 bar (PR = 3.7). The dashed part of the curves corresponds to CO 2 equilibrium molar fraction > 0.3. It can be seen that working at the optimum point with higher absolute pressures enhances efficiency through a temperature upgrade. (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.) Fig. 16. Efficiency as a function of PR for different values of the carbonation temperature. Red curve: turbine outlet pressure = 1 bar, carbonation temperature = 875 °C; operation point: carbonator pressure = 3.2 bar (PR = 3.2). Blue curve: turbine outlet pressure = 1 bar, carbonation temperature = 950 °C; operation point: carbonator pressure = 9.2 bar (PR = 9.2). The dashed part of the curves corresponds to CO 2 equilibrium molar fraction > 0.3. The horizontal arrow shows that working at higher pressure ratios is possible without penalizing efficiency through a temperature upgrade. (For interpretation of the references to color in this figure legend, the reader is referred to the web version of this article.) 600 R. Chacartegui et al. / Applied Energy 173 (2016) 589–605 operating conditions for the CSP–CaL integration is determined from lab-scale experimental measurements. It is important to remind here that most of previous lab-scale experimental measurements on the multicycle conversion of limestone derived CaO are carried out at the standard conditions corresponding to the CaL process for CO 2 capture [28]. These are rather different from those to be found in the CSP–CaL integration at the optimum efficiency operating conditions inferred from the present work. The CaL process requires usually a periodic purge of deactivated sorbent, which must be compensated by a makeup flow of fresh limestone in order to keep the mass balance. This is indeed the case of the CaL process for CO 2 capture. The specific CaL conditions for this integration involve calcination under high temperature at high CO 2 partial pressure and carbonation under low CO 2 partial pressure, which give rise to a severe drop of the CaO conversion after only a few cycles. After a large number of cycles CaO conversion converges asymptotically towards a residual value of just about 0.07–0.08 [65,66]. However, the specific conditions for an efficient integration of the CaL process into CSP plants involve, as seen above, carbonation under high CO 2 partial pressure at high temperatures. Fig. 17 shows recently obtained experimental results from thermogravimetric tests on the conversion of limestone derived CaO as a function of the calcination/carbonation Fig. 17. Multicycle conversion of limestone derived CaO as obtained from thermogravimetric analysis (TGA) tests under different CaL conditions (reported in [44]). CSP storage conditions: carbonation at 850 °C under pure CO 2 and calcination at 700 °C under pure He. CO 2 capture conditions: Carbonation at 650 °C under a 15% CO 2 /85% air vol/vol atmosphere and calcination at 950 °C under a 70%CO 2 /30% air vol/vol atmosphere. Residence times at both stages were 5 min. The solid lines are best fits from Eq. (10) to data. Fig. 18. Efficiency variation derived from a sensitivity analysis on ambient temperature, heat exchangers and solar daylight hours using as reference values those shown in Table 3. R. Chacartegui et al. / Applied Energy 173 (2016) 589–605 601 cycle number N under CaL conditions specific for both CO 2 capture and CSP energy storage [44]. The experimental results shown in Fig. 17 highlight the radical different behavior of natural limestone derived CaO as depending on the CaL conditions. In order to find out a value of the residual CaO conversion X r , the semi-empirical Eq. (10) can be fitted to conversion data [66]. X N X 1 ¼X r X 1 þ1 jðN1Þþ 1 X r X 1  1 0 B @1 C Að10Þ Thus, the obtained residual CaO conversion is just about 0.08 for CO 2 capture conditions whereas it becomes as large as X r = 0.53 for conditions that correspond to the CSP–CaL integration. Expectedly, Table 3 Reference values and ranges of variation for the parameters considered in the sensitivity analysis. Parameter Sensitivity on Ref. value Range of variation Ambient temperature Efficiency, Storage volume 20 °C0°C40 °C Carbonator thermal dispersions (% of carbonator thermal power production) Efficiency 10% 2% 15% CO 2 storage pressure Efficiency 75 bar 50 bar 120 bar Minimum temperature difference solid–solid exchangers Efficiency 20 °C10°C30 °C Minimum temperature difference gas–solid exchangers Efficiency 15 °C5°C25 °C Minimum temperature difference gas–gas exchangers Efficiency 15 °C5°C25 °C Pressure drops in carbonator circuit (% of carbonator pressure) Efficiency 0% 0% 20% Conveying consumption factor Efficiency 10 MJ/ 100 m/tonne 3 MJ/ 100 m/tonne 18 MJ/ 100 m/tonne Total solid conveying equivalent distance Efficiency 200 m 50 m 350 m Solar daylight hours Daily electricity production 12 h 6 h 12 h Fig. 19. Efficiency variation derived from a sensitivity analysis on carbonator thermal dispersion, system transport and storage pressure using as reference values those shown in Table 3. 602 R. Chacartegui et al. / Applied Energy 173 (2016) 589–605 this value will be even higher when the CaO solids are carbonated under CO 2 at over atmospheric pressure, which will enhance the reaction kinetics and increase the equilibrium temperature, thus allowing carbonation at even higher temperatures [62]. Importantly, this high residual CaO conversion is obtained for residence times of just 5 min both in the carbonation and calcination stages [44]. This would allow in practice the use of circulating fluidized bed (CFB) reactors that ensure a highly efficient transfer of mass and heat for solid–gas reactions and whose suitability to the CaL process is already proven in CO 2 capture pilot plants. Note that fast calcination is achieved at a reduced temperature of just 700 °C under a gas which is easily separable from CO 2 (either He as in the TGA experiments or superheated steam as in Catalytic Flash Calcination technology [52]), which would allow the use of already mature and inexpensive metallic solar receivers thus reducing technological risks. 4. Sensitivity analysis Results from a sensitivity analysis of the intervening parameters on plant efficiency are shown in Figs. 18 and 19 using the data reported in Table 3.InFig. 18 a sensitivity analysis on the effects of ambient temperature, heat exchangers design and solar daylight hours is presented whereas in Fig. 19 the effects on global plant efficiency of carbonator thermal dispersion, transport losses and storage pressure are shown. Reference values are obtained for X=0.5, carbonator pressure of 3.2 bar, turbine outlet pressure of 1 bar and carbonator temperature of 875 °C. According to the results obtained from the sensitivity analysis carried out on diverse parameters (Figs. 18 and 19) the following main conclusions may be drawn: – The sensitivity of efficiency to ambient temperature is high. An ambient temperature variation of 20 °C leads to a 6% relative variation of efficiency. The same increase leads to roughly 250% bigger volumes when storage pressure is 75 bar. By increasing the storage pressure above 100 bar, the volume growth is more restricted (around 36%). Due to the high ambient temperatures typically present in CSP plant emplacements, the addition of gas-cooling devices upstream compression to avoid this effect should be considered. – A ±70% variation of exchangers’ approach temperatures (excluded the gas–gas regenerator) produces a modest efficiency variation, which is below 3% in relative terms. – Unlike for the other heat exchangers, the approach temperature of the regenerator HXG (gas–gas exchanger) has a critical effect on performance. Reducing the value of D T min by 10 °C leads to a relative increase of efficiency greater than 7%, passing from 42.6% to 45.6% in absolute value. – Thermal dispersions in the carbonator and pipework have a major influence on efficiency: a decrease of thermal dispersion from 10% down to 2% of the reaction heat can lead to over 10% relative higher efficiency. This issue highlights the importance of providing efficient insulation to the hot components. – A 150 m conveying distance variation produces a relative efficiency variation below 3%. The same variation is caused by an 8 MJ/100 m/ton change of the specific conveying consumption factor. – Solids conveying energy consumption, as noted in previous paragraphs, can be quite critical if high values of CaO conversion are not achievable due to the higher amounts of solids in the loop. – The impact of pressure drops in the gas circuit is very relevant, with a 12% relative efficiency decrease when going from 0% pressure losses to 20% losses (% referred to the carbonator pressure). – The number of solar daylight hours influences the amounts of dissociated products (CaO and CO 2 ) available at the beginning of the day for energy production in the discharge phase. 5. Conclusions This paper explores a thermochemical energy storage concept in Concentrated Solar Power plants (CSP) based on the Calcium Looping process (CaL), which allows a fully decoupled operation of charge and discharge phases for long periods. It uses an abundant, geographically widespread, cheap and non-toxic raw material such as limestone. Moreover, the multicycle conversion of limestone derived CaO at specific CaL conditions for integration in CSP plants converges towards a rather high residual value as demonstrated by lab-scale TGA results. Therefore, affordable costs can be expected. The proposed CSP–CaL integration is technically feasible with mature technologies already available in the market for input heat powers (net heat supply at the calciner) ranging from small to above 100 MW th values. It uses a pressurized fluidized reactor for carbonation, solids pneumatic conveying technology, gas–solid cyclone/closed heat exchangers, gas–gas regenerators, a pressurized storage vessel at 75 bar for CO 2 , two atmospheric silos for CaO and CaCO 3 storage, and conventional intercooling compressor and turbine systems. A main feature of the integration is that a closed CO 2 circuit is used for operation of both the CaL process and the power cycle. The analyses presented in the paper shows that for a ratio of carbonator to outlet turbine pressures of 3.2 bars, and a carbonator temperature of 875 °C, plant efficiencies above 45% may be reached. The optimized CSP–CaL integration provides several benefits over previous studies, including complete absence of CO 2 released from the system, improved kinetics in the carbonator at elevated temperatures and high CO 2 pressures, high carbonator temperatures attainable and high thermal efficiency values. First law efficiency of 40–46% and second law efficiency of 43–48% are achievable under realistic conditions and taking into account heat losses, pressure drops and conveying energy consumption. According to our study, the main parameters to improve the proposed cycle are similar to those used in other power cycles, namely, turbine inlet temperature, pressure ratio in the Brayton turbine, boiler/reactor/heat exchangers’ efficiency or pressure losses taking into account the particularities linked to chemical equilibrium and reaction kinetics as determined by CO 2 partial pressure and temperature. The competitiveness of the proposed integration is reinforced by the fact that the obtained efficiency is higher than that of Rankine cycles traditionally employed in CSP plants. A detailed economic assessment is out of the scope of the present manuscript. However, it is estimated elsewhere that an efficient CSP–CaL integration would reduce the cost of Levelized Cost of Electricity (LCOE) below 7c€/kWh (see SUNSHOT program [67]). This estimation takes into account further improvements in the solar field, receiver and power block and considers a thermochemical energy storage cost below $15/kWh th . Acknowledgements This work was supported by the Andalusian Regional Government Junta de Andalucia (contract FQM-5735) and Spanish Government Agency Ministerio de Economia y Competitividad and FEDER Funds (contracts CTQ2014-52763-C2-1-R, CTQ2014- 52763-C2-2-R and MAT2013-41233-R). The Functional Characterization service of the Innovation, Technology and Research Center of the University of Seville (CITIUS) is gratefully acknowledged. R. Chacartegui et al. / Applied Energy 173 (2016) 589–605 603 3.2. Case 2: addition of solid-gas heat exchangers Compared to the base case, case 2 (Fig. 4) incorporates the use of solid-gas heat exchangers on both calciner and carbonator sides. In this way, it is possible to make a more profitable use of the high temperature heat stored in the streams exiting the reactors, which leads to an improved thermal integration. Solids heating could be performed in a suspension preheater where gas and solids are sequentially contacted in risers and separated by cyclones, as usually performed in raw meal preheaters of cement plants [60]. All solid-gas heat exchangers have been modelled by assuming a coflow arrangement and a same exit temperature of the gas and solids streams (as shown in Fig. 5). GS-HE1 is used to exchange heat between the CO 2 stream exiting the calciner at 900  C and the solids stream entering it, which must be brought to the calcination temperature, while heat Fig. 3. Proposed configuration in this work for the CSP-CaL base case. Table 2 Energy balance resulting from the simulations of the base case. Parameter Base case (Fig. 3) “sun mode”“night mode”“sun mode w/o storage” Solar thermal power to the calciner (MW th ) 100 0 33.33 Heat exchangers Thermal Power(MWth) HRSG 21.73 e7.24 COOLER-1 0.24 e0.08 COND 15.85 e5.28 HP-COMP (intercooler) 4.66 ee COOLER-2 4.15 ee HEATER e2.14 e TURB1 (interheater) e0.77 e COOLER-3 13.24 13.24 13.24 HXG 75.91 75.91 75.91 Power production (MWe) CO 2 storage turbine (HPS-TURB) e1.05 e Main CO 2 turbine (M-TURB) 23.92 23.92 23.92 Steam Turbine (ST) 5.80 e1.93 Power consumptions(MWe) Steam cycle pump (P) 0.04 e0.01 Main CO 2 compressor (M-COMP) 12.94 12.22 12.94 CO 2 storage compressor (HPS-COMP) 5.33 ee Auxiliaries for heat rejection (calciner side) 0.20 e0.04 Auxiliaries for heat rejection (carbonator side) 0.11 0.11 0.11 Auxiliaries for solids transport calciner 2.17 e0.72 Auxiliaries for solids transport carbonator 0.72 0.72 0.72 Net power (MWe) _ W net;sun 8.27 - 11.31 _ W net;night - 11.93 0 Overall plant efficiency h 0.321 0.339 C. Ortiz et al. / Energy 155 (2018) 535e551540 Fig. 4. CSP-CaL modified integration scheme (case 2). Fig. 5. TQ diagrams for calciner/carbonator sides in both the base case and case 2 configurations. C. Ortiz et al. / Energy 155 (2018) 535e551 541 exchangers GS-HE2 and GS-HE3 are used to preheat the CO 2 stream entering the carbonator. As a consequence, solids are stored in the vessels at lower temperature and lock hoppers also operate at lower temperatures. Streams properties is reported in Appendix while the energy balance for this case is shown in Table 3. Because of the preheating of the gas in GS-HE2 and GS-HE3, this enter the carbonator at higher temperature compared to the base case (719  C instead 654  C) and a higher CO 2 flow rate can pass across this reactor to carry the heat released by the carbonation reaction. As a result, the power produced by the main CO 2 turbine increases by a 12%. Fig. 5 shows the temperature-heat (TQ) diagrams for both the base case and case 2. As shown in Fig. 5, a better heat integration is achieved in the calciner side for the case 2. An enhanced utilization of hightemperature stream temperature is achieved, with the consequent reduction in exergy loses. The lower storage temperature in the CaCO 3 /CaO storage vessel also causes a reduction of the solids temperature at the inlet of the calciner (stream s3). Therefore, in the “sun mode”, a higher portion of the heat provided to the calciner is taken as sensible heat of the solids stream and less power is available for the calcination reaction, which lowers the production of CaO and CO 2 . The reduced flow rate of CO 2 combined with the lower temperature at the HRSG inlet (stream g1-1) after the solids preheater GS-HE1 causes a reduction of the steam generated and of the steam turbine power (47% compared to the base case). The reduction of the solids flow rate and of the CO 2 flow rate generated in the calciner also lead to a reduction of consumption of the HP CO 2 storage compressor and the auxiliaries for calciner solids transport (26%). Regarding the energy balance, an overall plant efficiency of 34.7% has been obtained, i.e. 2.6% points higher than the base case efficiency. The calculated net efficiency without energy storage also increases to 36.0%. Due to the reduced consumption for CO 2 storage compression, the differences between these two calculated efficiencies and between the net power produced in “sun”and “night” modes are also reduced. 3.3. Case 3: introduction of solid-gas heat exchangers and intercooled compression Compared with case 2, case 3 contemplates the use of an intercooled main CO 2 compressor, which allows reducing the consumption for CO 2 compression and achieving a higher cycle efficiency thanks to the presence of the regenerator. The lowpressure intercooled compressor is used to compress around 140 kg/s of CO 2 from atmospheric pressure to the carbonator pressure. In this case, two intercoolers are assumed with an intercooled temperature of 40  C, leading to compressor outlet temperatures of 73  C from each stage. This leads to a reduction of the CO 2 stream temperature at compressor outlet (g8) compared to case 2 (73  C instead 143  C), which allows for a higher heat exchange in HXG and therefore a reduced need for cooling in cooler-3. Results for this configuration can be seen in Table 3. By using a 2 intercooled-stages main compressor (M-comp), electric Table 3 Energy balance for the cases 2,3 and 4. Parameter Case 2 Case 3 Case 4 Including: solids heat exchangers Including: solids heat exchangers Intercooled compression Including: solids heat exchangers Intercooled compression P¼Patm carbonation “sun mode” “night mode” “sun mode w/o storage” “sun mode” “night mode” “sun mode w/o storage” “sun mode” “night mode” “sun mode w/o storage” Solar thermal power (MW th ) 100 0 33.33 100 0 33.33 100 0 33.33 Heat exchangers thermal Power (MWth) HRSG 11.63 e3.88 11.37 e3.80 11.51 e3.84 GS-HE1 4.39 e1.46 4.45 e1.48 4.42 e1.47 COOLER-1 0.25 e0.08 0.32 e0.11 0.28 e0.09 COND 8.49 e2.830 8.23 e2.78 8.40 e2.80 HP-COMP (intercooler) 3.47 ee 3.45 ee 3.46 ee COOLER-2 3.08 ee 3.06 ee 3.07 ee HEATER e1.58 ee1.57 ee1.58 e TURB1 (interheater) e0.57 ee0.56 ee0.74 e COOLER-3 15.54 15.54 15.54 6.02 6.02 6.02 6.04 6.04 6.04 HXG 84.60 84.60 84.60 93.27 93.27 93.27 93.68 93.68 93.68 GS-HE2 2.91 2.91 2.91 3.17 3.17 3.17 3.04 3.04 3.04 GS-HE3 7.19 7.19 7.19 7.15 7.15 7.15 7.17 7.17 7.17 Power inlet (MWe) CO 2 storage turbine (HPSTURB) e0.77 ee0.76 ee1.01 e Main CO 2 turbine (M- TURB) 26.87 26.87 26.87 26.64 26.64 26.64 26.77 26.77 26.77 Steam Turbine (ST) 3.10 e1.03 3.03 e1.01 3.07 e1.02 Power outlet (MWe) Steam cycle pump (P) 0.02 e0.01 0.02 e0.01 0.02 e0.01 Main CO 2 compressor (M- COMP) 14.67 14.11 14.67 13.26 12.75 13.26 13.28 13.20 13.28 CO 2 storage compressor (HPS-COMP) 3.98 e0.00 3.95 e0.00 3.96 e0.00 Auxiliaries heat calciner 0.12 e0.02 0.12 e0.02 0.12 e0.02 Auxiliaries heat carbonator 0.12 0.12 0.12 0.12 0.12 0.12 0.12 0.12 0.12 Auxiliaries solids transport calciner 1.60 e0.53 1.59 e0.53 1.60 e0.53 Auxiliaries solids transport carbonator 0.53 0.53 0.53 0.53 0.53 0.53 0.53 0.53 0.53 Wnet _ W net;sun 8.94 12.01 10.09 13.18 10.21 13.30 _ W net;night 12.88 14.01 13.92 - Overall plant efficiency h 0.347 0.360 0.381 0.395 0.381 0.399 C. Ortiz et al. / Energy 155 (2018) 535e551542 consumption is reduced by 6.8% and 8.8% in the “sun”and “night” modes, respectively, which implies an overall plant efficiency increase of 3.4%. 3.4. Case 4: carbonator at ambient pressure Case 4 allows operating the carbonator at ambient pressure. As a benefit, high temperature lock hoppers for solids pressurization would not be necessary. On the other hand, hermetic machines and heat exchangers to avoid air in-leakages as well as larger turbomachines (to handle the higher volume flow rate), larger carbonator and larger heat exchangers (to compensate the decrease of heat transfer coefficient with gas density) must be employed. Moreover, it is expected to achieve a lower reaction rate in the carbonator, which is favored by high CO 2 pressure [61]. Fig. 6 shows the case 4 scheme. M-comp compresses the CO 2 stream coming from the carbonator (g13) up to atmospheric pressure (plus about 0.1 bar to overcome pressure drops). Unlike in previous layouts in which g3 and g13 compressions were made completely in the M-compressor, in case 4 the CO 2 coming from the calciner (g3) is simply blown by a fan and is directly sent to the carbonator side, without passing through the M-comp. As shown in Table 3, there are no big differences in terms of efficiency between cases 3 and 4. Therefore, the most advantageous configuration between pressurized and atmospheric pressure carbonator must be chosen based on techno-economic considerations depending on the technical challenge and cost of a high temperature lock hopper system or the sealed components for a specific facility. The techno-economic analysis is being carried out by considering a quasi-stationary simulation to better assess the equipment sizes and costs and will be reported in a future publication. Note that all configurations have been modelled with PR ¼3. A higher temperature is achieved at the carbonator inlet in all cases where solid-gas heat exchangers are integrated (as shown in Appendix), which leads to a higher temperature of heat introduced in the thermodynamic cycle and therefore to a higher cycle efficiency. As a consequence, a larger mass flow rate of CO 2 can be heated up to 850  C for power production. Thus, 133.5 kg/s of CO 2 enter in the main turbine for the base case whereas in case 4 this value is increased up to 151 kg/s. Compared to the base case (where the global efficiency is 32.1%), the overall plant efficiency increases for all the modified configurations. As mentioned in the base case, also in these schemes an efficiency increase is achieved when no storage and SM ¼1are considered as the high-pressure compressor is not used. Thus, it is important to point out that the base operation mode selected for the simulations (“sun mode”,SM¼3) represents the worst scenario since throughout the day the solar power hitting the solar receiver will vary and therefore most of the time less than 100 MW th will be produced in the solar field. In practical operation conditions, the net global efficiency achieved in each case will be comprised between the weighed efficiency (1/3e2/3) and the efficiency in “sun mode w/o storage”depending on the characteristics of solar radiation for the selected site and on the solar field off-design performance. 4. Sensitivity analysis and discussion This section is devoted to compare the performance of the four configurations proposed for different gas cycle pressure ratio and Fig. 6. CSP-CaL modified integration scheme (case 4). C. Ortiz et al. / Energy 155 (2018) 535e551 543 number of main compressor intercoolers from the overall efficiency perspective. Moreover, to further explore the CSP-CaL integration, a sensitivity analysis has been carried out on case 3 about the effect of the turbine inlet temperature, CaO conversion, steam cycle conditions, turbomachinery efficiency, minimum D T in the main CO 2 heat exchanger, and CaO-CO 2 storage conditions. 4.1. Effect of pressure ratio In this section, all previous schemes are compared to analyze the most advantageous configuration. To this end, the pressure ratio ðPRÞ;defined as the ratio of the carbonator pressure to the turbine outlet pressure, has been chosen as dependent variable. Note that in the previous section all configurations are analyzed at PR ¼3. Fig. 7 shows the resulting overall plant efficiency as a function of the pressure ratio for the different cases and assuming a different number of intercooling stages in the main CO 2 compressor (for cases 3 and 4 only). As can be seen, cases 3 and 4 clearly show a higher efficiency than cases 1 and 2, particularly for higher pressure ratios. This illustrates the benefits of using the multi-stage intercooled compression combined with the gas cycle regenerator HXG. The simplest configuration (base case) reaches the maximum efficiency with PR of about 2.8, whereas in case 2 (where gas-solids heat exchangers are considered) efficiency is enhanced to near 35% at PR ¼3.2. Gas-solids preheating allows increasing the temperature of solids stream entering the carbonator and a higher CO 2 flow rate can be expanded in the main turbine (M-TURB). Thanks to the inter-cooled compression, the overall efficiency is further increased in case 3, reaching a value near 40% at PR between 4 and 5. Compared to previously reported CSP-CaL schemes [8,21], this relatively high efficiency is achieved in case 3 with a configuration of lower technical complexity and based on equipment already used in the cement industry, which would imply also a lower system cost. 4.2. Effect of turbine inlet temperature (TIT), CaO average conversion (X) and pressure ratio As it was shown in Fig. 7, there is a carbonator pressure for which the integration efficiency reaches a maximum, which justifies the relevance of an analysis on the pressure and temperature conditions in the carbonator to determine the optimal CaL- power cycle integration. Since the carbonator is directly connected to the turbine inlet, increasing the carbonator temperature leads to a higher power cycle efficiency. However, the maximum temperature in the carbonator is limited by the thermodynamic equilibrium of the calcination-carbonation reaction (Eq. (1)). According to thermochemical data, the equilibrium temperature T eq is related to the CO 2 partial pressure in the carbonator environment Pby means of Eq. (6) [28]: Teq ¼ 20474 lnP 4:137$10 7 (6) where Pis measured in bar and T eq in K. Thus, for a given value of P there is a maximum carbonator temperature (T max ¼T eq ) above which the CO 2 partial pressure is not sufficiently high for the reaction to be shifted toward carbonation (Fig. 8). Since in all considered cases carbonation is carried out under a pure CO 2 atmosphere, CO 2 partial pressure coincides with the carbonator absolute pressure. In order to get a grip on the role of turbine inlet temperature and pressure ratio, case 3 was simulated by considering the cycle without storage, thus all the CO 2 and CaO exiting the calciner side is sent to power production (SM ¼1). The analysis also contemplates several values of the average CaO conversion (X). This analysis is important because of the high uncertainty on the CO 2 carrying capacity of CaO under the carbonation and calcination conditions assumed in this work. Moreover, as discussed in section 2, CaO conversion is highly dependent on the carbonator-calciner conditions as well as on the CaO precursor. Table 4 shows the energy balance obtained by varying the carbonator temperature (and therefore the TIT) and the CaO average conversion. Note that for a better understanding, these balances are referred to the conditions without storage and therefore consumption of CO 2 storage compressor and power generated by the CO 2 storage turbine are zero. As expected, an increase in either CaO conversion or TIT leads to a higher overall plant efficiency. If TIT is raised from 850  Cto 950  C the overall efficiency is enhanced by about 1.2% and 2.7% points for X ¼0.1 and X ¼0.4, respectively. As shown in Table 4, Fig. 7. Sensitivity analysis on pressure ratio (PR). Fig. 8. Maximum carbonator temperature as a function of CO 2 partial pressure in the carbonator atmosphere according to thermodynamic equilibrium (Eq. (6)). C. Ortiz et al. / Energy 155 (2018) 535e551544 power consumptions due to solids conveying are the same in the calciner and carbonator sides since storage is not considered in this analysis. As can be seen, the steam cycle production is enhanced with the carbonator temperature. As the carbonator temperature is raised, the temperature of the solids entering the calciner (s2) increases as well, reducing the sensible heat required to heat them up to 900  C or, equivalently, increasing the CaCO 3 flow rate to achieve full calcination. Thus, a higher amount of CO 2 can be produced in the calciner (see g1 stream in Table 5). Furthermore, the net power produced by the CO 2 closed cycle is also enhanced with the carbonator temperature. An increase in CaO conversion enhances the steam cycle production albeit it does not lead to a higher net power output in the CO 2 closed cycle. This is mainly due to the reduction of CO 2 temperature at the carbonator inlet (Table 5) because of the lower solids to gas ratio in the CO 2 preheaters, which leads to a lower temperature of heat introduction into the gas cycle and therefore to a decrease of the cycle efficiency. A reduction of the CO 2 temperature at the carbonator inlet causes that a higher amount of carbonation heat is used to bring the reactants to the carbonation temperature and therefore, a lower amount of CO 2 can be recirculated (g10).Table 4 shows the carbonator heat used to preheat the reactants to the carbonator temperature. By considering TIT ¼850  C and X ¼0.1, a high amount of solids (compared with the X ¼0.4 case) is circulated from the calciner (at 900  C) to the carbonator, whose temperature is reduced. Therefore, this solids stream carries additional heat from the calciner to be used in the power cycle (as can be seen by the positive sign of the reactants pre-heating thermal power in Table 4). Fig. 9 shows the global CSP-CaL efficiency as a function of X and TIT. The different trends followed by the overall efficiency with X in the cases of TIT ¼850  C and 950  C may be explained by an analysis of the CaO temperature entering the carbonator from the calciner at 900  C. In the case of TIT ¼950  C, a part of the energy released in the carbonator must be used to bring the solids up to the carbonator temperature, being this effect more important when X is reduced (see reactants preheating power in Table 4). In the case of TIT ¼850  C, it is not necessary to heat up the solids since these come from the calciner at higher temperature, and therefore the loss of efficiency by reducing the CaO conversion is mitigated. Fig. 9 shows also efficiency results obtained by including or not auxiliaries' consumption. In regard to increasing the CaO conversion, the energy consumption linked to solids conveying is decreased, which Table 4 Energy balance for the case 3 as a function of TIT and X (SM ¼1, PR ¼3.5). Parameter X ¼0.1 X ¼0.4 TIT ¼850  C TIT ¼950  C TIT ¼850  C TIT ¼950  C Solar thermal power (MW th ) 33.33 33.33 33.33 33.33 Heat exchangers (MWth) HRSG 3.368 4.796 4.138 4.859 GS-HE1 1.279 0.868 2.072 1.693 COOLER-1 0.097 0.111 0.128 0.137 COND 2.457 3.499 3.019 3.545 COOLER-3 5.943 5.157 5.719 5.104 HXG 80.923 83.050 78.760 82.764 GS-HE2 2.891 4.299 3.261 4.077 GS-HE3 9.108 10.804 4.975 5.601 Power inlet (MWe) Main CO 2 turbine (M-TURB) 27.226 26.044 26.479 25.943 Steam Turbine (ST) 0.899 1.280 1.105 1.297 Power outlet (MWe) Steam cycle pump (P) 0.006 0.009 0.008 0.009 Main CO 2 compressor (M-COMP) 13.523 12.030 13.335 12.083 Auxiliaries heat calciner 0.020 0.029 0.025 0.029 Auxiliaries heat carbonator 0.120 0.105 0.117 0.105 Auxiliaries solids transport calciner 0.677 0.824 0.275 0.291 Auxiliaries solids transport carbonator 0.677 0.824 0.275 0.291 Carbonator heat (MWth) carbonation heat 19.253 23.442 25.717 27.137 reactants preheating (sensible heat) 2.217 4.698 0.320 2.573 losses 0.194 0.234 0.252 0.274 carbonator heat to power cycle 21.276 18.510 25.145 24.290 Power output CO 2 cycle (MWe) 13.704 14.014 13.144 13.860 Power output steam cycle (MWe) 0.893 1.271 1.097 1.288 Overall plant efficiency ð h Þ0.393 0.405 0.406 0.433 Table 5 Trends in selected streams by varying X and TIT. X TIT ¼850  C TIT ¼950  C g10 (kg/s) g1 (kg/s) s2 (TºC) s3 (TºC) g10 (kg/s) g1 (kg/s) s2 (TºC) s3 (TºC) 0.1 127.43 4.93 673.17 692.25 111.75 6.00 769.82 780.31 0.15 126.18 5.49 647.71 676.88 111.73 6.28 736.96 755.71 0.2 125.39 5.86 627.67 666.43 111.59 6.48 712.35 738.90 0.25 124.87 6.11 611.42 659.11 111.51 6.64 692.97 726.95 0.3 124.48 6.31 597.72 653.70 111.42 6.76 676.86 717.88 0.35 124.19 6.46 586.25 649.83 111.37 6.86 663.22 710.84 0.4 123.95 6.58 576.19 646.81 111.32 6.94 651.52 705.28 C. Ortiz et al. / Energy 155 (2018) 535e551 545 serves to enhance the overall net efficiency. These trends are also appreciated when overall average daily performance is considered. As showed in section 2, the energy density of the storage system is highly influenced by the CaO conversion, as shown in Fig. 10.CO 2 vessel volume is not affected by X, but is only dependent on pressure and temperature conditions. Note that the energy density is calculated by using Equation (2), and therefore considers not only the CaCO3 as solid material, but also the CO 2 and CaO vessels, the bulk porosity and packing density of the granular solids. Thus, energy density of the storage system infrastructure varies between 0.27 and 0.77 GJ/m 3 . In Table 6 three more cases, in which temperature and pressure in the carbonator are varied taking into account thermodynamic equilibrium restrictions (Fig. 8), are analyzed. As can be seen in Table 6, an increase of the carbonator temperature (and therefore of TIT) from 950  C up to 1050  C serves to enhance the overall plant efficiency by 0.8% points due to the increase of net power output in both the CO 2 and steam power cycles. As expected, an increase of the carbonator pressure (keeping fixed temperature) does not improve the cycle efficiency (as seen in Fig. 7 for other conditions). According to these results, the increase of the carbonator temperature improves the overall efficiency, albeit it would also add a higher degree of complexity for the turbine and the solid filtration system. On the other hand, varying the PR from 3 to 9 has an important effect on the heat exchanged in HXG (Table 6). This effect is caused by the reduction of the CO 2 temperature at the carbonator inlet, which decreases from 820  C (at PR ¼3) to 749  C (at PR ¼9) as a consequence of the isentropic CO 2 expansion process. A higher CO 2 temperature inlet allows a higher CO 2 mass flow rate to enter the power cycle since less carbonator heat is used to preheat the reactants to the carbonator temperature. Thus, the CO 2 mass flow rate decreases from 395 kg/s (at PR ¼3) to 231.3 kg/s (at PR ¼9). As a summary, case 3 simulation results indicate that increasing the temperature in the carbonator, and therefore the TIT, systematically leads to a higher overall plant efficiency. However, the maximum temperature in the carbonator is limited by the thermodynamic reaction equilibrium, which depends on the carbonator pressure. The process efficiency is also enhanced as the PR is increased up to a certain value of around 3e3.5 from which a further increase of the PR does not improve efficiency. Moreover, increasing the CaO conversion enhances the overall efficiency (reaching values about 43% at X ¼0.4) mainly because of the significant reduction of auxiliaries power consumption for solids transport. 4.3. Effect of steam cycle conditions, auxiliaries consumption and turbomachinery-heat exchangers efficiencies Following with the sensitivity analysis, this section assesses how the overall efficiency is affected by important parameters whose values were fixed in previous calculations (Table 1). Reference case conditions as well as the range of variation of these parameters are shown in Table 7. The values of those parameters not specified in Table 7 are left unchanged. Results of the sensitivity analysis on overall CSP-CaL efficiency are shown in Fig. 11. Compared to CSP-CaL schemes reported in previous papers [8,21,37] a main novelty of the present work is the introduction of a steam power cycle to take advantage of the high temperature CO 2 Fig. 9. Overall plant efficiency as a function of X for two different TIT (no storage is considered: SM ¼1). Fig. 10. Storage vessels and energy density as a function of CaO conversion. CO 2 is stored at 25 C; 75 bar. C. Ortiz et al. / Energy 155 (2018) 535e551546 stream exiting the calciner for power production. For this purpose, a simple Rankine cycle with moderate live steam conditions (400  C/40 bar) has been considered from the small power output of this cycle (around 3 MW e at plant design point). Considering these conditions, the net electric efficiency remains at 26.5% whereas for 540  C/100 bar live steam conditions the efficiency increases by 5% points. By computing the overall efficiency after this live steam conditions change, the new efficiency value is 0.6% Table 6 Energy balance for case 3 as a function of TIT and PR (without energy storage). Parameter X ¼0.15 T carb ¼950  CP carb ¼3 bar T carb ¼950  CP carb ¼9 bar T carb ¼1050  CP carb ¼9 bar Solar thermal power (MW th ) 33.33 33.33 33.33 Heat exchangers (MWth) HRSG 4.951 4.032 5.482 GS-HE1 1.082 1.401 0.885 COOLER-1 0.128 0.110 0.128 COND 3.613 2.941 3.999 COOLER-3 5.283 4.914 4.232 HXG 96.221 40.601 41.844 GS-HE2 4.388 3.794 5.241 GS-HE3 8.329 11.420 12.917 Power inlet (MWe) Main CO 2 turbine (M-TURB) 25.702 27.245 25.974 Steam Turbine (ST) 1.322 1.076 1.463 Power outlet (MWel) Steam cycle pump (P) 0.009 0.008 0.010 Main CO 2 compressor (M-COMP) 11.860 13.503 12.167 Auxiliaries heat calciner 0.030 0.024 0.033 Auxiliaries heat carbonator 0.105 0.112 0.100 Auxiliaries solids transport calciner 0.607 0.546 0.640 Auxiliaries solids transport carbonator 0.607 0.546 0.640 Carbonator heat (MWth) carbonation heat 25.010 22.496 26.361 reactants preheating (sensible heat) 3.714 3.893 10.062 losses 0.250 0.226 0.258 carbonator heat to power cycle 21.046 18.376 16.041 Power output CO 2 cycle (MWe) 13.842 13.741 13.807 Power output steam cycle (MWe) 1.313 1.069 1.453 Overall plant efficiency ð h Þ0.414 0.407 0.415 Table 7 Parameters variation throughout the sensitivity analysis on the overall CSP-CaL efficiency. Selected scheme case 3 (section 3.3) Main fixed parameters TIT 950  C PR 3.5 X 0.15 SM 3 Sensitivity analysis Parameter Ref. value Variation ID-Fig. 11 lower limit upper limit Rankine cycle live steam 400  C;40 bar 360  C; 40 bar 540  C; 40 bar Rankine 1 360  C;100 bar 540  C; 100 bar Rankine 2 Turbomachinery isentropic efficiency M-TURB 0.9 5% þ5% Isentropic 1 HPS-TURB 0.8 5% þ5% Isentropic 2 M-COMP 0.87 5% þ5% Isentropic 3 HPS-COMP 0.8 5% þ5% Isentropic 4 Turbomachinery mechanical efficiency M-TURB 0.98 2% þ2% Mechanical 1 HPS-TURB 0.96 2% þ2% Mechanical 2 M-COMP 0.98 2% þ2% Mechanical 3 HPS-COMP 0.96 2% þ2% Mechanical 4 Turbomachinery intercooled stages HPS-COMP 5 3 7 HP-ic-stage M-COMP 2 0 4 LP-ic-stage Heat exchangers minimum DT HXG (regenerator) 15  C10  C20  CHXG Cooling temperature Intercoolers Cooler-2 Cooler-3 40  C30  C50  CT-cooling Heat exchangers pressure drops Coolers HXG solid-gas HX 1% 5% 3% 10% þ10% P-drops Auxiliaries Heat rejection 0.8% of heat released 0.4% 1.2% Heat-rejec Solids conveying consumption 10 MJ/ton 5 MJ/ton 15 MJ/ton Solids C. Ortiz et al. / Energy 155 (2018) 535e551 547 higher than for the reference case. As shown in Fig. 11, the most critical parameters on the overall efficiency are those related to turbomachinery efficiency. Thus, increasing by 5% the reference isentropic efficiency value for the main CO 2 compressor (M-COMP) the overall efficiency rises up to 44%. Regarding the number of intercooled stages, variations in lowpressure compressor have a stronger effect on the overall efficiency compared to the high-pressure compressor since the penalty over the cycle caused by the former is higher. Fig. 11 shows also that variations in energy consumption linked to solids transport plays a significant role in the plant efficiency, as was also seen in Fig. 9. Regarding heat exchangers, a reduction by 5  C of the minimum temperature difference in the regenerator (HXG) leads to an enhancement of efficiency by 1%. 5. Conclusions The present work is focused on an energy integration analysis of a CSP-CaL plant. Four novel integration schemes have been developed by progressively increasing the level of complexity. Plant operation considering solar multiples (SM) of 3 (design-point for a constant power production throughout the day) and SM ¼1 (avoiding the energy storage equipment consumptions since lower solar radiation is assumed) were tested. As no high-pressure compressor is needed in the second case, the overall efficiency is increased by 1.8% points and therefore showing the penalty associated to the proposed high-pressure CO 2 storage system. The base case presents an overall efficiency of 32.1%. By adding gas-solid heat exchangers on both the calciner and carbonator sides (case 2) the overall plant efficiency reaches a 34.7%. Furthermore, if a 2-intercooled stage compression is used in the low-pressure compressor (which is the equipment with greater energy consumption in the plant) efficiency is enhanced up to 38.1%. Cases 3 and 4 present similar efficiencies. The optimum configuration for a specific plant will depend on techno-economic considerations. Results from a sensitivity analysis on the pressure ratio (PR) in the Brayton cycle shows that the highest efficiency is achieved for PR around 3e4.5 in all cases. A sensitivity analysis on the effect of CaO conversion and carbonator pressure/temperature shows that increasing the TIT under the limits posed by thermodynamic equilibrium enhances the overall plant efficiency. The overall efficiency is also enhanced as CaO conversion is increased (reaching values about 43% at X ¼0.4), mainly because of the significant reduction of auxiliaries power consumption. A new expression for the energy density is proposed in this paper, which takes into account the size of the infrastructure (including all vessels and the packing density of solids). For the proposed cases, the energy storage density, mainly dependent on CO 2 pressure, CO 2 temperature and CaO conversion, varies between 0.2 and 0.9 GJ/m 3 . Our study gives support to the potential benefits of using the Calcium-Looping process for thermochemical energy storage in CSP plants. Major technological challenges are yet to be faced for the plant assessed in this work, especially related to the design of the solar receiver/calciner and to the high temperature solids pressurization and depressurization system. In future works, a design study of the solar receiver, hour-by-hour simulations of the plant considering variable heat input in the solar calciner and technoeconomic analysis need to be undertaken to further assess the feasibility and competiveness of the CSPeCaL integration. Acknowledgments This work has been supported by the Spanish Government Agency Ministerio de Economia y Competitividad (MINECO-FEDER funds), contracts CTQ2014-52763-C2-2-R, CTQ2017-83602-C2-2-R and FPI contract granted to Carlos Ortiz Dominguez (BES-2015- 0703149) within the project: Hybrid thermochemical storage of concentrated solar power (CTQ2014-52763-C2-2-R). Fig. 11. Sensitivity analysis on overall CSP-CaL efficiency results (see ID definition in Table 7). For an easier understanding, bars are presented in the same order than legend and a colour code is used. The reader is referred to the web version of this article. (For interpretation of the references to colour in this figure legend, the reader is referred to the Web version of this article.) C. Ortiz et al. / Energy 155 (2018) 535e551548 Appendix Table 8 shows the main stream data for all simulated cases. Table 8 Main stream data for the CSP-CaL integration ID Base case Case 2 Case 3 Case 4 PðbarÞTðºCÞ_ mðkg=sÞPðbarÞTðºCÞ_ mðkg=sÞPðbarÞTðºCÞ_ m ðkg=sÞ PðbarÞTðºCÞ_ mðkg=sÞ s1 1.10 850 72.19 1.20 850 53.47 1.20 850 53.06 1.00 850 53.27 s1-1 eee 1.17 718.92 53.47 1.17 718.81 53.06 0.97 718.90 53.27 s1-2 eee 1.14 665.03 53.47 1.14 659.62 53.06 0.94 662.40 53.27 s2 1.10 850 216.56 1.14 665.03 160.40 1.14 659.62 159.17 0.94 662.40 159.80 s3 eee 1.11 692.17 160.40 1.11 687.39 159.17 0.91 689.84 159.80 c1 1.00 900 193.75 1.00 900 143.51 1.00 900 142.41 1.00 900 142.97 c2 1.00 900 64.58 1.00 900 47.84 1.00 900 47.47 1.00 900 47.66 v1 0.074 40.13 7.14 0.074 40.13 3.82 0.074 40.32 3.74 0.074 40.32 3.78 v2 45 40.53 7.14 45.00 40.53 3.82 45.00 40.72 3.74 45.00 40.72 3.78 v3 40 400 7.14 40.00 400 3.82 40.00 400 3.74 40.00 400 3.78 v4 0.075 40.32 7.14 0.075 40.32 3.82 0.075 40.32 3.74 0.075 40.32 3.78 g1 1.00 900 22.81 1.00 900 16.89 1.00 900 16.76 1.00 900 16.83 g1-1 eee 0.97 692.17 16.89 0.97 687.39 16.76 0.97 689.84 16.83 g2 0.97 52.03 22.81 0.94 56.92 16.89 0.94 62.03 16.76 0.94 58.95 16.83 g3 0.96 40 22.81 0.93 40 16.89 0.93 40 16.76 0.93 40 16.83 g4 3.21 142.69 15.21 3.31 148.29 11.26 3.31 74.36 11.18 1.14 40 11.22 g5 75.75 123.66 15.21 75.75 124.27 11.26 75.75 124.27 11.18 75.75 124.27 11.22 g6 74.25 130 7.60 75.00 130 5.64 75.00 130 5.59 75.00 130 5.61 g7a 3.21 142.69 126.25 3.31 148.29 141.8 3.31 74.36 140.58 1.14 45.28 141.30 g7b 3.21 38.57 7.60 3.31 39.08 5.64 3.31 39.08 5.59 1.14 21.92 5.61 g8 3.21 138.83 133.72 3.31 143.44 142.38 3.31 72.78 141.19 1.14 72.81 142.14 g8-1 eee 3.31 143.44 5.05 3.31 72.78 4.98 1.14 72.81 4.77 g8-2 eee 3.21 665.03 5.05 3.21 659.52 4.98 1.10 662.39 4.77 g9 3.05 654.1 133.72 3.14 679.55 142.38 3.14 679.54 141.19 1.08 679.56 142.14 g9-1 eee 3.14 679.05 147.43 3.14 678.87 146.17 1.08 679.00 146.91 g9-2 eee 3.05 718.93 147.43 3.05 718.81 146.17 1.05 718.94 146.91 g10 3.00 850 126.25 3.00 850 141.8 3.00 850 140.58 1.00 850 141.30 g11 1.00 694.53 126.25 1.00 694.54 141.8 1.00 694.54 140.58 0.33 694.56 141.30 g12 0.95 153.83 126.25 0.95 158.65 141.8 0.95 87.98 140.58 0.32 88.01 141.30 g13 0.94 40 126.25 0.94 40 141.8 0.94 40 140.58 0.31 40 141.30 Notation Acarbonator cross-section, m 2 T i;vessel storage temperature for of the i-component,  C c p;i specific heat, kJ/(kmol$K) TIT turbine temperature inlet, ºC E den energy density, GJ/m 3 n i specific volume for the i-component, m3/kmol ε i porosity of the i-component V i storage vessel volume for the i-component, m 3 F i molar flow rate of component i, kmol/s _ V CO2 CO 2 volume flow rate, m 3 /s F CaCO 3 CaCO 3 molar flow rate _ W net average electrical power, MWe F CaCO3;crb CaCO 3 molar flow rate (calciner side) _ W net;night net electrical power for the night mode, MWe F CaO;unr molar flow rate of unreacted CaO (calciner side) _ W net;sun net electrical power for the sun mode, MWe F CO 2 ;clc CO 2 molar flow rate at calciner outlet _ W MTURB power produced by the main CO 2 turbine, MWe F CO 2 ;crb;in CO 2 molar flow rate at carbonator inlet _ W MCOMP power consumed by the main CO 2 compressor, MWe F CO 2 ;crb;out CO 2 molar flow rate at carbonator outlet _ W HPSTURB power produced by the high-pressure CO 2 turbine, MWe F R;crb recirculating molar flow rate (carbonator side) _ W HPSCOMP power consumption of high pressure intercooled CO 2 compressor for the storage system, MWe F R;clc recirculating molar flow rate (calciner side) _ W ST power produced in the steam turbine cycle, MWe h i Enthalpy, kJ/kmol _ W P power consumed in the steam turbine cycle, MWe _ mmass flow rate, kg/s _ W PSOLCAL Power consumptions for solids transport in the calciner side, MWe Ncycle number _ W PSOLCAR Power consumptions for solids transport in the carbonator side, MWe Ppressure, bar _ W AUXPOWCA auxiliary power consumptions in the calciner side, MWe P carb absolute carbonator pressure, bar _ W AUXPOWCR auxiliary power consumptions in the carbonator side, MWe P eq CO 2 partial pressure at equilibrium, bar Xaverage CaO conversion PR pressure ratio X r residual CaO conversion _ Q input solar power input D Ppressure drop at carbonator, bar SM solar multiple D t sun average daytime period (h) Ttemperature,  C D H R ðT react Þreaction enthalpy at the reactor temperature, kJ/mol T calciner calciner temperature,  C D H 0 R standard enthalpy of reaction, kJ/mol T carb carbonator temperature,  C h global net efficiency T eq equilibrium temperature, ºCfpacking density C. Ortiz et al. / Energy 155 (2018) 535e551 549 [Document text truncated for crawler view.]