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Fuel 113 (2013) 826–836
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Fuel
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Calcium looping for CO2 capture from a lignite fired power plant
Ilias Vorrias a,b, Konstantinos Atsonios a,b,⇑, Aristeidis Nikolopoulos a,b, Nikos Nikolopoulos b,Panagiotis Grammelis b, Emmanuel Kakaras a,b
a Laboratory of Steam Boilers and Thermal Plants, National Technical University of Athens, 9 Heroon Polytechniou Street, 15780 Zografou, Greeceb Center of Research & Technology Hellas/Chemical Process and Energy Resources Institute (CERTH/CPERI), Gr-15310, Greece
h i g h l i g h t s
" Process optimization of the CaLtechnology in a lignite-fired powerplant.
" Fuel pre-drying and solids heatexchanger enhance the systemefficiency.
" CaL process has greater efficiencythan MEA and oxyfuel.
0016-2361/$ - see front matter � 2013 Elsevier Ltd. Ahttp://dx.doi.org/10.1016/j.fuel.2012.12.087
⇑ Corresponding author at: ARKAT building, 357-3Athens, Greece. Tel.: +30 210 6501509; fax: +30 210
E-mail addresses: [email protected] (I. Vorrias), [email protected] (A. Nikolopoulos), [email protected] (P. Grammelis), [email protected]
g r a p h i c a l a b s t r a c t
a r t i c l e i n f o
Article history:Received 29 September 2012Received in revised form 24 December 2012Accepted 31 December 2012Available online 23 January 2013
Keywords:Calcium loopingCO2 captureLignite power plantSupercritical cyclePost-combustion CO2 capture
a b s t r a c t
Calcium looping (CaL) process is considered as a competitive technology for the capture of CO2 emittedfrom fossil fuel power plants. In this process two fluidized bed reactors are coupled. In the first reactor,i.e. carbonator, CO2 is absorbed by CaO and CaCO3 is produced, while in the second one, referred as cal-ciner, fuel is oxy – fired, so that calcium particles, already having absorbed CO2 in the carbonator in orderto produce CO2-rich flue gases, are regenerated. Numerous studies have focused so far on the implemen-tation of this process in coal fired plants. However, little work has been performed in terms of lignite firedplants. Thus here, the knowledge gaps are filled by means of a thorough investigation of the particulartechnology using a thermodynamic approach. It is worthwhile noticing, that the analysis is conductedon an existing lignite power plant, which is already operating and is located at the wider area of Meliti,Greece (330 MWe lignite fired power plant). For a more profound work, a series of parametric scenariosare investigated, for their optimum coupling setup, which subsequently lead to the conclusion that theenergy penalty for the sorbent regeneration in cases of lignite combustion is higher than in cases of coalcombustion. This is due to the higher moisture content and lower calorific value of lignite compared withcoal. Finally, in order to achieve a lower energy penalty, several configurations are presented, resultingthat the optimum is to burn pre-dried fuel, utilize solid heat exchanger and insert the fresh limestoneinside the calciner instead of the carbonator. All configurations are evaluated in terms of CO2 captureand net efficiency.
� 2013 Elsevier Ltd. All rights reserved.
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59 Mesogeion Ave., Halandri,[email protected] (K. Atsonios),[email protected] (N. Nikolopoulos),.gr (E. Kakaras).
1. Introduction
According to Kyoto protocol and Copenhagen Accord [1], allindustrialized countries are committed under binding obligations
Nomenclature
Ecarb carbonator capture efficiencyECO2 carbon capture rate of the total plantF molar flow rate (kmol/s)m mass flow rate (kmol/s)T temperature (�C)X carbonation conversion (kmol/kmol)ycomb ratio of the fuel that is burnt in the calciner to the total
fuel
subscripts0 fresh limestonecarb carbonatedcalc calcinedR circulating CaOref reference plant
AbbreviationsASU Air Separation UnitBFB Bubbling Fluidized Bed
DFB Dual Fluidized BedFGC Flue Gas CondenserFGD Flue Gas DesulphurisationFGR Flue gas RecirculationHEC Heat ExchangerHP High PressureIP Intermediate PressureMEA MonoethanolamineLHV Lower Heating ValueLP Low PressurePPC Public Power CorporationPFD Process Flow DiagramPH PreheaterPP Power plantRH ReheaterSH SuperheaterSPECCA Specific Energy Consumption for CO2 AvoidedST Steam Turbine
Table 1Simulation approaches for CaL process in literature.
Authors Carbonator model Calciner model
Abanades et al. [8] EQUIL EQUILWang et al. [9] EQUIL EQUILShimizu et al. [10] SIMPL SIMPLMartinez et al. [11] RIGOR RIGORBarelli et al. [12] EQUIL EQUILHughes et al. [13] EQUIL SIMPLMahishi et al. [14] EQUIL –Strohle et al. [15] SIMPL EQUILHawthorne et al. [16] RIGOR RIGORAlonso et al. [17] RIGOR –Romano [3] RIGOR –Wang et al. [18] SIMPL EQUILLi et al. [19] EQUIL EQUILKunze et al. [20] SIMPL SIMPL
I. Vorrias et al. / Fuel 113 (2013) 826–836 827
to reduce their CO2 and five other GHG’s emissions. In this light,the Intergovernmental Panel on Climate Change (IPCC) assessesthat the sufficient stabilization of atmospheric CO2 concentrationwould be achievable by the realization of policies and technicaladvancements. Among them, the energy demand reduction in theindustrial sector and the efficiency increase in old-fashioned tech-nologies, are included.
Concerning coal and lignite combustion, which is widely usedfor the production of energy and heat, one of the latest develop-ments for carbon footprint mitigation, includes CO2 capture andsequestration technologies [2]. One of these technologies is thecarbonation–calcination (CaL) process, which seems to be quitepromising as a post-combustion capture technology, as it takesadvantage of the reversible carbonation reaction of CaO, in orderto absorb, transport and release an almost pure CO2 stream forits subsequent treatment and storage. As this technology is consid-ered as highly competitive because of its aforementioned advan-tage, when compared with other technologies (e.g. aminescrubbing, oxyfuel combustion [3,4]), especially for retrofittingexisting plants, numerous modeling and experimental investiga-tion studies have been performed in order to get a further insightand optimization of the process [5,6]. The vast majority of thesestudies examine the application of CaL process into coal firedpower plants (PPs). However, lately Berstad et al. [7] investigatedthe coupling of a Natural Gas Combined Cycle case with the CaLlooping process.
The CaL process employs of two reactors, i.e. the carbonator andthe calciner, in which CO2 capture and sorbent regeneration takeplace, respectively. Therefore, a proper modeling of these reactorsis essential for process optimization. For the best achievement ofthis prerequisite, various approaches from simpler to more compli-cate can be followed. Table 1 depicts the modeling techniques usedby various authors for the simulation of both the carbonator andthe calciner. Primarily the simplest modeling approach (SIMPL) isbased on the assumption of a constant conversion of the carbon-ation and calcination reactions. The second one, which is consid-ered as the most common, is followed be the hypothesis ofthermodynamic equilibrium inside the reactors (EQUIL), an ap-proach closer to reality, especially in cases that the calcium reac-tion is conducted simultaneously with other reactions (i.e.gasification, combustion). However, simulation of CaL process isalso succeeded by applying the kinetic rates of the most significant
reactions taking place on the platform of more rigorous thermody-namic models, emerging from experimental data (RIGOR). Such anapproach can also take into consideration, except for the reactionkinetics, the governing granulate hydrodynamics and sorbentactivity.
In 1999, Shimizu et al. [10] investigated the integration of CaLprocess with a coal-fired power plant, depicting the high powerconsumption for (a) an almost 100% O2 stream utilized for theoxy-combustion of coal within the calciner production and (b)the pure CO2 stream compression exiting the calciner. Neverthe-less, they came to the conclusion that despite these drawbacksfrom an efficiency point of view, CaL process still has a highernet efficiency when compared to oxy-fuel combustion. Duringthe next years, following this first examination, several similarstudies [11,15,16,18] investigated the implementation of CaL pro-cess in existing coal-fired plants. This post-combustion captureprospect was proved to be marginally better than other options,e.g. O2/CO2 and amine scrubbing, in terms of plant overall net effi-ciency. Romeo et al. [21] performed the exergy analysis of an exist-ing coal-fired power plant integrated with CaL process, stressingout the importance of proper heat integration with the secondarysteam cycle. As a result, the following years several works focusedon the definition of the optimum process configuration and its cou-pling with the thermal cycle of the existing PP [8,22] paying atten-tion on the definition of the fresh limestone flow rate [23].
Table 2Raw lignite analysis.
C (w/w%) H (w/w%) S (w/w%) O (w/w%) N (w/w%) H2O (w/w%) Ash (w/w%) LHV (kJ/kg)
22.58 2.07 0.94 9.88 0.37 36.8 27.36 7831
Table 3Main characteristics of the reference case.
Coal consumption kg/s 98.16Gross power output MWel 330.60Gross electrical efficiency % 42.45Pumps electric consumption MWel 18.15Fan electric consumption MWel 8.37Net power output MWel 304.15Net electrical efficiency % 39.05Specific CO2 emissions kgCO2/MWel 959.20
828 I. Vorrias et al. / Fuel 113 (2013) 826–836
The vast majority of such literature examinations refer to coal-fired power plants, an exception of which is the work of Romeoet al. [24], who investigated a lignite-fired plant, assuming yet thatthe calciner is bituminous coal – fired. However, for power plantslocated near lignite mines, the use of another type of fossil fuelfor feeding the calciner is not respected as techno-economicallyfeasible. Thus, the utilization of lignite as a feeding fuel in the cal-ciner of CaL and as a mean for both power and heat productionmust be demonstrated. For this reason, this study for the first timedeals with the implementation of CaL technology for CO2 capturein an existing power plant exclusively fired with lignite.
2. Description and modeling approach of the energy system
The thermodynamic calculations are performed by the combi-nation of two commercial programs, i.e. ASPEN Plus™ and IPSEPro™. The detailed modeling of CaL capture process, including theoperation of the Dual Fluidized Bed (DFB) reactors, the essentiallignite drying process (to increase the overall efficiency of theplant), the CO2 stream purification and compression, as well asthe O2 production in the Air Separation Unit (ASU), are simulatedin ASPEN Plus™. The integration of this complex block, with thesecondary steam cycle of the PP is performed with IPSE Pro™. Inorder to extract safe conclusions for the performance of the CaLprocess with respect to other CO2 capture technologies, like MEAscrubbing and oxyfuel, it is essential to apply the same main pro-cess specifications for all cases examined. Therefore, the secondarysteam cycle and CO2 compression train characteristics and the spe-cific consumption for oxygen production in the ASU are consistentwith European Benchmark Task Force (EBTF) common framework[25]. The composition of the lignite used is presented in Table 2and it is in accordance with EBTF definitions. The ambient condi-tions, i.e. temperature, relative humidity and air pressure, are setto 15 �C, 60% and 1.013 bar respectively.
2.1. The 330 MWe lignite power plant
As a reference PP, the 330 MWe lignite-fired unit of PPC locatedat the area of Florina (north Greece), has been selected. In order toenhance its efficiency, several modifications have been made to theexisting power plant, prior to its integration with any of the exam-ined CO2 capture technologies (CaL, amine scrubbing and oxyfuelcombustion). The major parameters that have been improved, arethe boiler air leakage, the boiler slagging and fouling, the con-denser vacuum, the air leakage at the air preheater (LUVO) andthe flue gas temperature at the LUVO outlet. The description ofeach section of the reference power plant is presented in the fol-lowing paragraphs.
The plant is equipped with a supercritical Benson type once-through boiler with single reheat. The boiler produces 265.36 kg/sof high pressure steam (235.4 bar, 540 �C) and 225.58 kg/s ofintermediate pressure steam (39.28 bar, 540 �C). The boiler’s effi-ciency is 91.6% and CO2 production is 81.1 kg/s.
The generated steam after being superheated enters the highpressure steam turbine, where it is expanded. It is subsequently re-heated and led to the IP steam turbine, the LP steam turbine andfinally to the condenser. The steam cycle includes an 8-stageregenerative water preheating (4-stage for condensate and 3-stagefor feedwater), using steam extractions from HP (High Pressure), IP(Intermediate Pressure) and LP (Low Pressure) steam turbinestages as follow:
� 56.97 bar, 42.45 bar 19.37 bar for preheating the HP feedwater.� 12.62 bar for using in the deaerator.� 5.69 bar, 2.22 bar, 1.04 bar and 0.225 bar for preheating the LP
feedwater.
The feedwater pump is electrically driven. The heat in the con-denser is dissipated by means of a cooling water cycle, which iscooled in a natural draught wet cooling tower. The main character-istics of the plant are presented in Table 3. The condenser pressureis set equal to 48 mbar and the cooling water inlet temperature at18.2 �C, with a terminal temperature difference between conden-sate and cooling water of 3 �C. Before the stack, there is a WetDesulphurization Unit that reduces the SOx emission below400 mg/Nm3. Therefore, it is assumed that no desulphurizationtakes place in the carbonator.
2.2. Calcium looping process
Fig. 1 depicts the CO2 capture process. The CO2 capture isachieved in the fluidized bed carbonator, where CO2 reacts withCaO particles, producing CaCO3, according to the followingreaction.
CaOþ CO2 ! CaCO3 DH� ¼ �178 kJ=mol ðR1aÞ
The equilibrium CO2 pressure of the reaction favors the CaCO3
formation at temperatures around to 600–650 �C. The regenerationof sorbents takes place in the second fluidized bed (calciner):
CaCO3 ! CaOþ CO2 DH� ¼ þ178 kJ=mol ðR1bÞ
Since the (R1b) is endothermic and is conducted at higher tem-perature (900–950 �C) than the (R1a), any additional fuel is neededto be maintained under autothermal conditions. In order to obtaina highly pure CO2 stream, the oxidizing agent is an O2-rich stream,produced in the ASU. Subsequently, the produced oxygen stream ismixed with a portion of the purified CO2 stream, originating fromthe exit of the calciner, in order to lower the O2 concentration inthe oxidizing mean, to avoid the introduction of quite hot spotsin the reactor, especially close to its inlet area (combustion under100% O2). This technological limitation is taken into considerationby some authors [7,23,24], whereas in other studies is assumedthat there is no need for CO2 recirculation [10,26]. The CO2 – richflue gas stream after the calciner exit, mainly composed of CO2
and H2O, is cleaned (fly ash removal), cooled and purified throughwater condensation, before being compressed and sent for storage.
Fig. 1. Sketch of calcium looping process.
Table 4Simulation assumptions.
Circulation rate FR/FCO2, (kmol/kmol) 7.0Fresh limestone F0/FCO2, (kmol/kmol) 0.1Carbonation temperature (�C) 650Calcination temperature (�C) 900Hot/cold temp. approach at the solid heat exchanger (�C) 10O2 content in pure oxygen stream (% v/v) 95.0O2 content in oxidizing mean (% v/v) 80.0
Fig. 2. Lignite drying system.
I. Vorrias et al. / Fuel 113 (2013) 826–836 829
The carbonation modeling is based on the chemical equilibriumbetween inlet CO2 (FCO2) and the maximum fraction of CaO avail-able to form CaCO3 (Xave�FR). The Xave is calculated by the empiricalequation proposed by Abanades et al. [8]:
Xave ¼fmð1� fwÞF0
F0 þ FRð1� fmÞþ fw; ð1Þ
where fm and fw are constants that depend on sorbent physical/chemical characteristics. The carbon capture rate within the carbo-nator can be calculated from the following equation:
Ecarb ¼FRðXcarb � XcalcÞ
FCO2; ð2Þ
where Xcarb is the sorbent carbonate content after the carbonatorand Xcalc is the sorbent carbonate content after the calciner.
Inside the calciner, apart from the lime regeneration, the lignitecombustion also takes place. As depicted in Table 1, the equilib-rium assumption is adopted by the majority of the researchers.This assumption, which reflects the requirement for Gibbs free en-ergy minimization, is as well adopted in the recent study. In phys-ical terms, this hypothesis does not deviate from the reality sincethe high operating temperature and the fluidization phenomenaenhance equilibrium conditions. The basic assumptions of the casestudy of the present simulation are summarized in Table 4.
2.3. Description of the pre-drying system
As the fuel utilized in this study is lignite, its moisture content ishigh and therefore a drying system before its oxy combustion atthe calciner is needed; thus contributing to overall energy effi-ciency increase of the plant. For the pre-drying process of lignitethe fluidized bed technology is proposed (Fig. 2). An option is touse the hot CO2 – rich stream as the fluidization mean. However,in order to avoid additional particle removal systems, this optionis not adopted, by the present study. Moreover, during the dryingprocess the CO2 stream may be contaminated by non-condensablecompounds as N2 and Ar. An alternative option is the exploitationof the low enthalpy steam, derived from the LP Steam Turbine ofthe secondary steam cycle. A third option is to use the H2O stream(derived from the evaporation of lignite water content), which isthe most beneficial among the aforementioned ones in efficiencyterms.
In this context, a portion of the evaporated moisture is used asthe fluidizing agent in the BFB (Bubbling Fluidized Bed) dryer,while another part of it is condensed inside an internal heat ex-changer, providing the required heat for drying. The condensateis used for the fuel preheating up to a temperature equal to around80 �C. The drying temperature is set to 110 �C. Two blowers areused in the system: one for the compression of the fluidizationsteam, and the other for the compression of the stream, which con-densates in the heat exchanger. The second compression is neededfor the increment of the condensation temperature by 10 �C great-er than the drying temperature, enabling in this way the efficientheat transfer. This technology is chosen among others as fluidizedbeds are characterized by high heat and mass transfer coefficientsthat beneficiate the reduction of drier size and the exergy effi-ciency. Although the addition of the blower increases the electricalconsumption, it offers the advantage of the drying process decou-pling from the rest of the system.
Table 5Dryer main characteristics.
Final moisture (% w/w) 12.0Drying temperature (�C) 110Specific energy consumption (kWe/kg dry fuel) 133.3Fraction of bed occupied by heat exchanger (%) 14.3%Recirculating steam (kgsteam/kgdry fuel) 0.502Part of steam that is utilized as fluidizing medium (%) 75.2
Fig. 4. Air separation unit flow diagram.
830 I. Vorrias et al. / Fuel 113 (2013) 826–836
The specifications of the pre-drying system are summarized inTable 5.
2.4. Solid recirculation heat exchanger
The common characteristic of the chemical absorption systemsis the different levels of temperature between the absorption andthe regeneration reactor, the temperature of the later to be higher.For reasons of energy penalty reduction, a heat exchanger is addedbetween the two main streams. In calcium looping case, the heatexchanging between the solid recirculation streams has beneficialimpact on total efficiency [23] because of the high temperature dif-ference between carbonator and calciner. However, due to the sol-idus phase of the streams, the configuration of the heat exchangershould be feasible. Thus, a configuration of concentric L-valves ispresented in this study. According to this, the L-valves of the twofluidized bed reactors are located concentrically and they are trea-ted as a co-current heat exchanger so as the outlet stream to havethe same temperature (see Fig. 3). The fact that the vertical parts ofthe L-valves are fixed bed enhances the heat transfer, reducing thedimensions of the configuration.
At the frame of this work is not easy to determine how this con-cept can be scaled-up, since there are several uncertainties regard-ing the coupled operation of the two reactors (dual fluidized bed-DFB) with the L-valve, in a large scale. Several of them are relevantto the size of the L-valve, limited by the existence of the cyclonesfor both the reactors. Nevertheless, the introduction of one or moreconcentric L-valves will contribute to the heat exchanging of atleast one portion of the circulating solids mass, as the presentstudy suggests, from a thermodynamic point of view.
2.5. Air separation unit
The best option for large scale pure oxygen production is thecryogenic method that typically takes place in an Air SeparationUnit (ASU). The separation of air into two streams, i.e. one purein nitrogen and one pure in oxygen, is based on the differentdew point of these compounds under high pressure conditions.The process flow diagram is depicted in Fig. 4. According to it,the air stream (1) enters the four staged inter-cooled compressorand is compressed (2). Subsequently, it is cooled down to 12 �C
Fig. 3. Concentric L-valves for heat exchanging between the circulating solids.
by passing through two stages of a Direct Contact Air Cooler(DCAC), so as to remove a part of its moisture. Downstream thissystem, it is further cooled down to 9 �C, at the evaporative coolersby exchanging heat with the pure nitrogen stream. Before the mainheat exchanger, the air separation is completed at the molecularsieves adsorbers where water and other impurities are completelyremoved (3).
Further, the air stream is cooled down near to dew point (4),and the air separation is performed, after passing through twodistillation columns with different operating pressures. Aroundto 34% of the separation is performed within the high pressurecolumn, while the outgoing streams (5, 6) undergo throttling andenter the low pressure column, where the separation of O2 isconcluded, without the need for external heating or cooling. Therequired cooling load for condensation at the HP column is ex-tracted from the evaporation heat that is released in the reboilerof the LP column. The operating pressures of the two distillationcolumns are determined by the temperature approach of the outletstreams of the reboiler-condenser. Finally, the outgoing productstreams from the low pressure column (9, 10) are heated up to15 �C in the main heat exchanger, by the air coming from themolecular sieves. The parameters of the ASU operation are setaccording to the EBTF [25], in terms of the specific powerconsumption for the production of 95% pure oxygen stream. Theprocess parameters and the main characteristics of the ASU arepresented in Table 6.
2.6. CO2 purification and compression unit
At this main component of the plant, the calciner off-gases un-dergo CO2 purification, resulting in a high purity CO2 stream (morethan 90% v/v). In addition, the EBTF limits for various components,i.e. O2, N2, H2O and other impurities, are as well satisfied. Hence, acomplete gas cleaning system for solids (ash and other particles)
Table 6Process specifications of the ASU modeling.
HPC and LPC pressure bar 5.50/1.92Oxygen pressure outlet bar 1.35Air compressor isentropic eff. % 86.5%O2 recovery efficiency % 99.1%N2 recovery efficiency % 99.7%Oxygen purity % v/v 95.0%Nitrogen purity % v/v 99.6%Specific O2 production kW h/tnO2 222.36
I. Vorrias et al. / Fuel 113 (2013) 826–836 831
and water removal should be implemented after the calciner. Anadditional desulfurization system is not required, on the groundsof fuel sulfur in situ capturing in the form of CaSO4. Additionally,the combustion conditions inside the calciner benefit for reducedNOx emissions compared to conventional air combustion [27,28].Thus, a de-NOx unit is probably not required. The removal of theH2O content is succeeded by its condensation at the ambient tem-perature. Subsequently, the pure-CO2 stream is compressed up to80 bar after passing through three inter-cooled stages. At eachinter-stage, the gas is cooled down to 28 �C in order to reducethe energy penalty. The supercritical outlet gas is also cooled downto 28 �C and pumped up to 110 bar, conforming with the require-ments for pressure to be equal to the appropriate delivery pressurefor storage.
2.7. Secondary steam cycle and heat integration
The CO2 capture process produces a significant amount of heatenergy, which is recovered, so as to increase the electrical powerproduction of the total plant. In this light, a secondary steam cyclewith single reheat is implemented in order to produce supercriticalsteam that is expanded at a secondary ST. The thermodynamicscharacteristics of both superheated and reheat steam are in accor-dance with EBTF and identical to the corresponding characteristicsof the produced steam from the main boiler. The water streamafter the feedwater pump is split in two portions; the first steamstream is preheated at a heat exchanger by the CO2 lean carbonatorgases, while the second stream is preheated by the CO2 rich calcin-er flue gases. The flue gases streams outlet temperature is equal to206 �C.
The preheated water streams are mixed and led to the Carbona-tor for evaporation, taking advantage both of the heat produced bythe exothermic carbonation reaction and the sensible heat of theincoming hot solid stream (CaO). The produced steam is thensuperheated by the CO2 lean-flue gases, thus exiting the carbonator
Fig. 5. Heat utilization fo
at a temperature equal to 650 �C. The final thermodynamic charac-teristics of the produced HP steam are 235.4 bar and 540 �C.
The steam reheating takes place at a heat exchanger, which uti-lizes the heat energy of the CO2 rich flue gas exiting the calciner ata temperature equal to around 900 �C. The thermodynamic charac-teristics of the reheated steam are 56.77 bar and 540 �C. The con-densate stream is preheated by passing through three heatexchangers without any steam extraction from ST; firstly by thehot CO2 stream compression train and the ASU cooler, secondlyby a portion of the CO2 rich stream that is extracted before the fluegas condenser and finally from the CO2 lean gases of the carbonator(before the stack). The condensates temperature entering thedeaerator is 147 �C.
Since the cooling water that exits the abovementioned heatexchangers has a significant mass flow rate and low temperature(1044 t/h, 88 �C), can be potentially used for district heating. In thiscase, the heating needs of about 58 MWth can be covered, with re-sult an extra increase in the mixed efficiency of the power plant.
In order to increase the overall efficiency from the heat utiliza-tion, the heat from the cooling of the purge stream (Ca particles) isused to preheat the oxygen and CO2 that fluidize the calciner. Theoutlet temperature of both gas and solids is 270 �C.
The secondary steam cycle is presented in Fig. 5.
3. Results and discussions
3.1. CaL process simulation results
In this section, the performance of the CO2 capture unit is inves-tigated. The numerical results are derived by a thermodynamicsimulation performed on the basis of the assumptions and thesub-processes presented in the previous sections, i.e. implementa-tion of fuel drying, solid heat exchanger with the process parame-ters being stated in Table 4 (case study). The simulation results areenriched with a sensitivity analysis of various parameters, so as to
r steam generation.
Table 7Process simulation main results.
Xcarb 0.13Xcalc 0.00Ecarb 89.65%ECO2 93.92%ycomb 0.367O2/CO2,capt (kg/kg) 0.336
Qheat (MJth/kgCO2 capt, LHV based) 3.65
Paux (MJe/kgCO2 capt, LHV based) 0.78
Table 8CO2 balance at the carbon capture system.
CO2 from main boiler flue gases 58.67%CO2 from lignite combustion at the calciner 36.35%CO2 from make-up limestone degradation 5.87%CO2 emitted 6.07%CO2 captured 93.93%
Fig. 6. Moisture content impact on CO2 capture process.
Fig. 7. Percentage distribution of the heat input of the secondary fuel (a) with solidheat exchanger and (b) without solid heat exchanger.
832 I. Vorrias et al. / Fuel 113 (2013) 826–836
fine tune the operation of the CaL process, in order to achieve themaximum net efficiency of the coupled CaL process with the mainsteam boiler circuit. The basic results of the case study are pre-sented in Table 7. The ycomb refers to the ratio of fuel combustedin the calciner to total fuel consumption:
ycomb ¼mf ;calciner
mf ;main boiler þmf ;calcinerð3Þ
The fuel consumption is increased by 36.8% compared to the nocapture case, where fuel is consumed only in the main boiler. Thecapture efficiency of the carbonator is 89.65%, whereas the overallcarbon capture efficiency is almost equal to 94%. The carbon bal-ance is demonstrated in Table 8. A considerable portion of the pro-duced CO2 originates from the additional oxy-fuel combustion inthe calciner. The specific heat requirements for CO2 capture aresimilar to MEA scrubbing. However, the relatively low operatingtemperature in the latter technology does not beneficiates thethermal exploitation of the outgoing streams, as CaL process does.The specific oxygen demand is about one third of the correspond-ing oxygen demand for the case of CO2 capture using the oxyfueltechnology (around 0.88 kg/kg [29]). The main characteristics ofthe most important streams (gas and solids) of CaL process are gi-ven in Table 9.
The effect of the supplementary fuel pre-drying on fuel con-sumption and capture efficiency is depicted in Fig. 6. The require-ment for supplementary fuel can be reduced by 5% withoutconsiderably affecting the efficiency of capture rate. For the caseof raw-fuel entrance into the calciner, without that upstream beingpre-dried, the total carbon captured mass and its corresponding
Table 9CaL process main streams characteristics.
m T Mole composition (v/v)(kg/s) (�C) CO2 (%) H2O (%)
Gas streamsFlue gas from the main boiler 403.8 165.6 12.34 20.91Flue gas after the carbonator 332.3 650 1.44 23.56Air inlet to the ASU 187.7 15 0.03 1.01Oxygen stream 44.6 15 0.00 0.00Flue gas after the calciner 156.3 900.0 73.23 22.88Final CO2 stream 130.1 25 94.25 0.75CO2 recirculated 9.7 28.0 92.34 2.76Solid streamsSolids inlet to carbonator 700.5 772.4 – –Solids outlet to calciner 711.8 900.0 – –Solids inlet to calciner 770.3 762.4 – –Solids outlet to carbonator 770.3 650.0 – –Fresh limestone 17.7 25.0 – –Purge lime 11.3 772.4 – –Secondary fuel 58.3 25.0 – –
efficiency (ECO2) is greater compared to the case that an additionalstep of pre-drying is implemented, because of the increase of CO2
production in the calciner. On the other hand, the carbonator cap-ture efficiency (Ecarb) remains constant. The higher the moisturedecrease in the dryer is, the higher the energy penalty associatedwith this process is. Moreover, the available heat for steam produc-tion drops down, due to the reduction of calciner flue gases massflow rate.
Fig. 7 depicts the distribution percentage of the heat inputamong the various energy requirements, denoting the beneficialimpact of heat exchanging between the solid streams, on thereduction of secondary fuel consumption. Based on our calcula-tions, it is expected that if the heat exchanging concept is adopted,
Mass composition (w/w)N2 (%) O2 (%) Ar (%) CaO (%) CaCO3 (%) Ash (%)
62.83 2.98 0.75 – – –70.80 3.36 0.85 – – –77.30 20.74 0.92 – – –
1.51 95.00 3.49 – – –0.76 1.88 1.24 – – –0.75 2.42 1.60 – – –0.96 2.38 1.57 – – –
– – – 100.0 0.00 0.00– – – 97.78 0.00 2.22– – – 79.23 20.77 0.00– – – 79.23 20.77 0.00– – – 0.00 100.0 0.00– – – 97.78 0.00 2.22– – – – – –
Fig. 8a. Effect of circulating solids on capture efficiency and on fuel consumption.
Fig. 8b. Effect of fresh limestone on capture efficiency and on fuel consumption.
I. Vorrias et al. / Fuel 113 (2013) 826–836 833
17% of fuel heat is consumed in order to increase the sensible heatof solids from the inlet temperature up to the calcination temper-ature. Moreover, almost half of the heat input is used for the sor-bent regeneration. If a solids heat exchanging stage is notapplied, the required heat input percentage consumed for the cir-culating solids temperature increase increases significantly. Thedecrease in fuel consumption (i.e. ycomb) is preferable with respectto economic cost and efficiency. However, high fuel consumptionbeneficiates the heat of outgoing gases and steam production.
As concerns the solid circulation rate, Fig. 8a reveals that a fur-ther increase of it does not affect straightforward increase the cap-ture efficiency above a certain ratio, equal to around 7. On theother hand, fuel consumption increases, due to the higher solid in-let flux to the calciner that must be heated up to the calcinationtemperature. To conclude, the fresh limestone injection rate hasa small beneficial effect on the capture efficiency, though it in-creases the ycomb ratio (Fig. 8b).
Table 10Summary of the results of all examined cases.
Overall gross power output MWel
ASU electric consumption MWel
CO2 compressor electric consumption MWel
Overall net power output MWel
HP steam production of secondary cycle kg/sMass recirculating CO2 kg/sRaw supplementary fuel consumption % of total fuelSpecific raw supplementary fuel consumption kg/MWhe net
Dry supplementary fuel consumption % of total fuelSpecific supplementary dry fuel consumption kg/MWhe net
Specific CO2 capture rate kg/MWhe net
3.2. Process integration results
3.2.1. Sensitivity analysisSimulations have been performed for different cases, which
are differentiated by the following three parameters, (a) thecombustion of raw or pre-dried lignite in the calciner, (b) the appli-cation of a heat exchanger between the solid streams recirculatedbetween the reactors and (c) the place where fresh limestoneis introduced in the CaL process. A very short characteristicdescription of the cases examined is given below:
� (CASE A-Case study): In this case, the supplementary fuel that isused for oxy-combustion in the calciner is pre-dried, coupledwith a heat exchanging step of the circulating solids. This canbe achieved during the CaL process by the implementation ofan L-valve system. Moreover, the fresh make-up limestone isconsidered to entering the calciner.� (CASE B): In this case, the supplementary fuel undergoes com-
bustion without this first upstream being pre-dried.� (CASE C): In this case no heat exchange between the circulating
solids occurs.� (CASE D): In this case, the fresh make-up limestone enters the
carbonator instead of the calciner.
In order to compare the results of the aforementioned cases, interms of efficiency, the SPECCA coefficient is introduced. SPECCA(Specific Energy Consumption for CO2 Avoided) expresses the addi-tional fuel (in MJ) that is required to be consumed in order to avoidthe emission of 1 kg of CO2:
SPECCA ¼ 3600 �1g� 1
gref
Eref � EðkJLHV=kgCO2Þ; ð4Þ
where E is the CO2 emission rate in kg CO2/kW hel, and g the netelectrical efficiency of the plants. The index ref refers to the refer-ence plant without the implementation of any CO2 capture system.
The results of the aforementioned cases simulation are pre-sented in Table 10, and Figs. 9a–9c. It is derived that the CO2 cap-ture with CaL technology increases significantly the electricproduction, due to the higher heat production from the captureplant, compared to a plant which does not have a CaL system in-stalled. This additional heat can be used for the production of en-ergy in a secondary steam cycle.
Fuel pre-drying significantly decreases the fuel demands in thecalciner (comparison between Case A and B). As a result, reductionof the CO2 produced in the calciner and consequently of the steamat the secondary cycle (about 11.1%) takes place, thus leading inturn to a reduction in electricity production. Although fuel pre-dry-ing requires an additional electrical consumption, this additionalstage results in an increased net electric efficiency, due to the low-er fuel consumption, the lower extra fuel demand for the CO2 cap-
CASE A CASE B CASE C CASE D
546.7 571 634.8 532.636.4 44.3 49.6 34.340.2 43.4 45.6 39.4428.2 445.8 490.5 416.8146 162.3 210 135.511.1 14.1 15.7 10.937.86 42.98 45.85 36.89502.8 565.3 609.9 495.842.90 – 59.70 26.56360.7 – 438.2 355.81267.0 1323.6 1265.3 1286.1
Fig. 9a. Gross electric efficiency.
Fig. 9b. Net electric efficiency.
Fig. 9c. SPECCA index.
834 I. Vorrias et al. / Fuel 113 (2013) 826–836
ture (SPECCA index) and the lower electric consumption for theASU and the CO2 compressors.
The installation of an intermediate solid heat exchanger reducesthe temperature difference between the solids flowing betweenthe calciner and the carbonator (comparison between Case A andC). This decrease results in the reduction of the supplementary fuelconsumption and the heat output of the carbonator. As a result,these reductions cause a significant decrease in the steam produc-tion (about 30.5% compared) and consequently a decrease the elec-tric production of the installation. On the other hand, lower fuelconsumption reduces the efficiency penalty of the capture, sinceless pure oxygen is needed. Additionally, the lowered CO2 capturedmass reduces the electric consumption for CO2 compression. Thesebenefits result in a considerable overall net electrical efficiency in-crease and a significantly smaller in size calciner.
The introduction of the fresh make-up limestone in the carbo-nator is also examined and compared with the respective case thatthis is introduced into the calciner (examined case D). For this to beachieved additional energy should be consumed for heating up thiscold stream and therefore the steam production in the carbonatorwill be reduced. However, the fuel consumption and the amount ofCO2 released within the calciner are as well reducing. All in all, the
overall steam production is 7.2% less compared to case A, while thenet electric production and efficiency are decreasing.
Fig. 10 presents the energy flow diagram (SANKEY diagram) forcase study (Case A). The heat power entering the carbonator com-prise of 86 MWth contained within the flue gases of the main PPand 430 MWth within the incoming solid stream (CaO), whereas303 MWth are yielded by the (exothermic) carbonation reaction.The outgoing heat from the carbonator is distributed among theflue gases stream (poor CO2), the solid stream (rich CaCO3) andthe evaporated steam, in order to increase their temperature.
As far as the heat input into the calciner is concerned, this com-prises of 516 MWth provided by the combustion of the pre-driedfuel, 620 MWth by the solid stream (rich CaCO3) exiting the carbo-nator, 20.1 MWth carried by the preheated oxygen and the CO2
recirculated stream. However, only 356 MWth out of the total heatinput is utilized for the sorbents regeneration. 162 MWth from thehot CaO exiting the calciner is recovered from the solid stream thatenters the calciner, in the concentric L-valve heat exchanger.
As a result, the total heat input produced in the CaL processbeing able to be recovered from the secondary steam cycle is equalto 322 MWth from the carbonator (contained within the flue gases,and the heat release along the riser which is utilized for steamevaporation) and 180 MWth from the flue gases exiting the calcin-er. In addition, 18 MWth are recovered from the intercooling of CO2
stream and air in the ASU and are utilized for the LP feedwater pre-heating. Finally, 216 MWe gross is produced by the overall 520MWth, which are recovered from the secondary steam cycle. Theremaining heat of 295 MWth from the overall 520 MWth is dis-carded via the condenser and the cooling tower out to the atmo-sphere and 9 MWth are mechanical and electrical loses at thepumps the STs and the electrical generator.
3.2.2. Effect of CO2 recirculation for the dilution of the oxygen in theoxidizing agent
Beyond the aforementioned cases, the effect of CO2 recircula-tion for the dilution of the oxygen in the oxidizing agent on theplant efficiency is examined. The cold CO2 recirculation requiresfor an extra heat demand in the calciner, increasing the supple-mentary fuel consumption and consequently reducing the net elec-tric efficiency of the PP by ca. 0.05%. However, the use of an inertgas stream to dilute the O2 concentration in the oxidizing meanis essential for an oxy-combustion environment, in order to keepthe combustion temperature into acceptable levels for the calcinermaterials and to create a capable gas flow to transfer the fuel intothe combustor.
3.2.3. Comparison with other CO2 capture technologies for a retrofitcase
In order to evaluate the CaL technology as a retrofit option forCO2 capture, the results of case A are compared against other com-petitive CO2 capture technologies, such as amines scrubbing andoxyfuel combustion [30–32]. Table 11 depicts that the couplingof a PP with CaL process produces more electric power than thetwo others, owed to the benefit of the additional steam cycle inte-gration, as afore-explained. Moreover, the net electric efficiency ishigher for CaL, especially compared to the MEA case. Concerningthe oxyfuel combustion, its net efficiency as a CO2 capture technol-ogy is smaller compared to the respective CaL, without referring tothe fact that the implementation of this technology in a retrofitcase requires multiple important technical modifications at theboiler island (extended modifications in heat transfer surfaces,burners, flue gas recirculation and utilization of special materialsfor construction of O2 transfer tubes).
Fig. 10. SANKEY diagram for the secondary steam cycle.
Table 11Comparison with other CO2 capture technologies.
Ca-Lprocess
MEA Oxyfuel
Overall gross power output (MWel) 546.7 299.4 357.44Overall gross electrical efficiency (%) 43.53 38.44 45.89ASU electric consumption (MWel) 36.4 – 53.4CO2 compressor consumption (MWel) 40.2 24.1 24.1Overall net power output (MWel) 428.2 243.43 258.22Overall net electrical efficiency (%) 34.09 31.25 33.20Energy penalty (%) 4.96 7.8 5.85Specific fuel consumption (kg/MWhe net) 1328.0 1450.8 1378.8Specific mass of CO2 captured (kg/MWhe net) 1267.0 1346.4 1270.8SPECCA (kJLHV/kgCO2) 5.897 10.238 7.153
Table 12Comparison between retrofit and new build plants using the carbonation–calcinationtechnology.
Retrofit(Case A)
Newbuild
Overall gross power output MWel 546.7 552.3Overall gross electrical efficiency % 43.53 44.33ASU electric consumption MWel 36.4 35.6CO2 compressor electric consumption MWel 40.2 34.9Overall net power output MWel 428.2 435.3Overall net electrical efficiency % 34.09 34.94Specific raw fuel consumption kg/MWhe net 1328.0 483.8Specific mass of CO2 that captured kg/MWhe net 1267.0 1237.2SPECCA kJLHV/kgCO2 5.897 4.072
I. Vorrias et al. / Fuel 113 (2013) 826–836 835
3.2.4. Comparison of a new-built with a retrofit PP, applying CaLtechnology
The case of a new-built power plant equipped with the CaLtechnology is as well examined. For the carbonation–calcinationcapture technology the configuration again of case A is applied.In order to be able to operate the plant with and without CO2 cap-ture, two independent steam cycles with two STs have beenconsidered.
For this to achieve and to maximize the overall electric effi-ciency, heat integration of the two STs with the main boiler andthe CaL unit should be performed, thus requiring significant inter-ventions for this case, compared to a retrofit one. Additionally, thefirst LP water preheater of the secondary steam cycle takes advan-tage of the heat that the cooling water carries after exiting the FGC.The cooling water for intermediate cooling between the CO2 com-pression stages and the ASU is at first used to preheat the combus-tion air of the main boiler prior to LUVO and then for heating up
the feedwater flowing inside the second LP preheater of the mainpower plant. Moreover, the first LP PH of the main PP can usethe heat of the condensate exiting the first LP PH of the secondarysteam cycle, thus allowing for the disregard of steam extractionfrom ST.
To conclude, Table 12 presents the comparative characteristicsof the retrofit case A and the new built PP, integrated with CaL pro-cess, where the net electric efficiency of the new built PP is in-creased approximately by 2.5%, while the SPECCA is reducedapproximately by 30% compared to the retrofit option.
4. Conclusions
In this study, the post-combustion capture in a lignite-firedpower plant employing the calcium looping technology was inves-tigated. All the components of a retrofit and/or a new built PP, were
836 I. Vorrias et al. / Fuel 113 (2013) 826–836
presented, described and modeled in detail. This was performed byutilizing two accredited commercial software (ASPEN Plus™ andIPSE Pro™). From the detailed engineering of all cases examined,one can derive the main interventions that need to be performedin order to optimize the efficiency of the plant by simultaneouslyreducing the supplementary fuel consumption. The most impor-tant among these, include the integration of the PP cycle with a lig-nite pre-drying step and an intermediate solid heat exchanger atthe CaL unit. On the contrary, CO2 recirculation slightly reducesthe electric efficiency, but its implementation is considered as nec-essary so that (a) to keep the combustion temperatures at the samelevel with the air-fired combustion and (b) the flue gas has the nec-essary mass flow rate and momentum to transfer the lignite parti-cles towards the calciner.
As any CO2 capture technology, CaL process reduces consider-ably the net electric efficiency when integrated in a PP. However,this technology when compared to other more standard CO2 cap-ture technologies, as MEA and oxy-combustion, allows for a higherelectric production and efficiency both for a retrofit and a new builtPP.
Acknowledgements
This study has been carried out in the framework of the Euro-pean Commission – Research Fund for Coal and Steel ContractNo. RFCR-CT-2010-00013 (CAL-MOD).
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