Scholarly article on topic 'Capture of CO2 during low temperature biomass combustion in a fluidized bed using CaO. Process description, experimental results and economics'

Capture of CO2 during low temperature biomass combustion in a fluidized bed using CaO. Process description, experimental results and economics Academic research paper on "Chemical engineering"

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Energy Procedia
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{"CO2 capture" / "Calcium looping" / Carbonation / "Biomass combustion" / "Negative emissions"}

Abstract of research paper on Chemical engineering, author of scientific article — M. Alonso, N. Rodriguez, B. Gonzalez, B. Arias, J.C. Abanades

Abstract The process discussed in this work is aimed at the capture of CO2 with CaO in a circulating fluidized bed (CFB) combustor-carbonator reactor, where the combustion of biomass with air and the carbonation of CaO are taking place simultaneously. This process exploits the high reactivity displayed by most natural biomass materials during their combustion at low temperatures (around 700 °C) and the capability of CaO to absorb CO2 at these temperatures. This is a niche application for the calcium looping cycles being developed for other post-combustion and precombustion processes. The CaO necessary for the CO2 capture reaction by carbonation is being circulated from an oxyfired circulating fluidized bed calciner, operating at temperatures over 900 °C to permit the decomposition of CaCO3. In this publication, we present new experimental results proving the concept in a 30 kW test facility consisting of two interconnected CFB reactors (combustor-carbonator and combustor-calciner). An application case study is discussed, together with an economic analysis using published data of similar systems.

Academic research paper on topic "Capture of CO2 during low temperature biomass combustion in a fluidized bed using CaO. Process description, experimental results and economics"

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Energy Procedia 4 (20 11) 795-802

Energy Procedía


Capture of CO2 during low temperature biomass combustion in a fluidized bed using CaO. Process description, experimental results

and economics.

M. Alonsol*, N. Rodriguez, B. Gonzalez, B.Arias, J.C. Abanades

CSIC-INCAR, Francisco Pintado Fe, 26,33011Oviedo, Spain


The process discussed in this work is aimed at the capture of CO2 with CaO in a circulating fluidized bed (CFB) combustor-carbonator reactor, where the combustion of biomass with air and the carbonation of CaO are taking place simultaneously. This process exploits the high reactivity displayed by most natural biomass materials during their combustion at low temperatures (around 700°C) and the capability of CaO to absorb CO2 at these temperatures. This is a niche application for the calcium looping cycles being developed for other post-combustion and pre-combustion processes. The CaO necessary for the CO2 capture reaction by carbonation is being circulated from an oxyfired circulating fluidized bed calciner, operating at temperatures over 900°C to permit the decomposition of CaCO3. In this publication, we present new experimental results proving the concept in a 30 kW test facility consisting of two interconnected CFB reactors (combustor-carbonator and combustor-calciner). An application case study is discussed, together with an economic analysis using published data of similar systems.

© 2011 Publishe dby Elsevier L td.

Keywords: CO2 capture, calcium looping, carbonation, biomass combustion, negative emissions

1. Introduction

The integration of CCS with biomass energy system in the generation of electricity or other energy products leads to CO2 negative emissions. This was initially recognized by Ishitani and Johansson [1]. The positive implication of

* Corresponding author. Tel.: +3-498-511-8836; fax: +3-498-529-7662. E-mail address:


negative emissions energy technologies for long term climate change mitigation has been highlighted in some recent scenario exercises [2, 3].

All main capture routes are in principle valid for biomass firing [4]. In what refers to pre-combustion, one of the most innovative processes for hydrogen production from biomass is the Absorption Enhanced Reforming process[5-8], which is a technology for hydrogen rich gas generation from biomass gasification in the presence of solid particles of CaO. This novel gasification technology has already been demonstrated at the MW scale [5]. In this process, the CaO acts as a CO2 absorbent, capturing CO2 in the gasifier to form CaCO3, and shifting all the reforming and water gas shift equilibrium towards H2 production. The CaCO3 formed in the gasifier is then recirculated to a calciner where CaO is regenerated by burning residual char with air. This is a conceptually similar process to the Acceptor Gasification Process for coal [9], developed during the 60s and early 70s in the US. These gasification processes with high temperature CO2 sorbent (or "acceptor") were not designed for CCS applications, because the CO2 from the calcination of CaCO3 is released in diluted form in the flue gases of the air combustor-calciner. However, the interest for different versions of these technologies has grown in recent years by incorporating a calcination step to the process in conditions that allow for a pure stream of CO2 suitable for large scale storage [10-12]. At present the most realistic options for this type of calcination is the oxyfiring of more fuel in the calciner, so that the CO2 evolved in the calcination of CaCO3 is obtained together with the CO2 resulting from the oxycombustion of the fuel.

This paper discusses a new process that shares some of the principles outlined above but one which is focused on only combustion conditions. In the process discussed in this work, the CaO is able to capture CO2 "in situ" during the combustion of biomass in a circulating fluidized bed combustor-carbonator reactor [13, 14]. In a second reactor, the combustor-calciner, the reverse calcination reaction of CaCO3 takes place. Heat for the calcination reaction is supplied through the oxyfuel combustion of a fuel. Therefore, the overall reactions taking place in each reactor are:

C (from biomass) + CaO (s) + O2 (from air) ^ CaCO3 (s) + Heat (1)

CaCO3 (s) + Heat (from oxycombustion) ^ CaO (s) + CO2 (g) (2)

Heat is the only useful product of reaction in this CO2 capture process. It is therefore essential to allow for an effective heat recovery scheme in a steam cycle, using existing technologies for CFB combustion power plants. In this work, we present a case study to allow for an economic estimation of the key cost variables of the process (cost of electricity and avoided cost of CO2) under certain boundary conditions. These conditions are chosen using as a basis the new experimental data and design criteria derived from these experiments, together with a set of general cost parameters recently published for other applications of CCS to biomass firing [15].

2. Experimental results

The concept of biomass combustion with in situ CO2 capture was tested in experiments conducted in a 30 kW test facility, which has been described in detail elsewhere [16]. This is made up of two interconnected circulating fluidized bed reactors: a carbonator and an air-fired calciner (it is assumed that the sorbent performance is similar that when operating the calciner in oxy-fired mode). The carbonator is 6.5 m and the calciner 6.0 m in height. Both reactors have an internal diameter of 0.1 m. Temperatures, gas compositions at the exit of the reactors and pressure drops are continuously monitored. This last measurement allows for an accurate estimation of the bed inventory of solids at any given time in each reactor. Solid circulation rates and solid composition in the reactors or in the circulating solids are measured sporadically. The typical analysis on solid samples extracted from the experiments is the carbonate content (which allows, together with the solid circulation rates, a reasonable closure of mass balances as reported in [16]) and the maximum carbonate conversion (which is the maximum carrying capacity of the solids and allows the estimation of the fraction of active CaO in the circulating solids, Xactive=Xmax-Xcarb). Several tests were carried out at different carbonation reaction temperatures, solids inventories, superficial gas velocities, solids circulation flow rates and concentrations of CO2 generated by biomass combustion. Three different biomasses were used: sawdust, crushed olive pits and wood pellets. Although most of the experimental data were obtained at non steady-state conditions [16], we have also achieved many periods of trouble free operation in stationary state (of up to 14 hours duration). An example of recent experimental results is presented in Figure 1.

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Figure 1. Example of experimental results in the combustor-carbonator reactor in a typical experiment. (a) Combustor-carbonator exit gas concentrations of CO2 and O2 measured by the on-line gas analyzer and O2 zirconia probe (b) Experimental capture efficiency and maximum capture efficiency allowed by equilibrium.

In this example, the average CO2 concentration at the exit of the combustor-carbonator reactor was 3.1 vol%. The average oxygen concentration at the exit of the combustor-carbonator was 7.6 vol.% measured by the on-line gas analyzer and by the O2 zirconia probe installed at the top of the riser. From a combustion mass balance, the CO2 produced by biomass combustion was estimated to be around 13.6 vol.% Therefore, the average CO2 capture efficiency was 77%. The average temperature in the combustor-carbonator was 690 °C during the period, which according to the equilibrium of CO2 on CaO allows for a 2.4 vol% of CO2 (maximum efficiency allowed by the equilibrium of 81%, remarkably close to the experimental value). This means that at this particular set of conditions (superficial gas velocities around 2m/s, solid circulation rates around 1.6 kg/m2s and bed inventories around 200400 kg/m2) the active CaO in the reactor is an effective CO2 absorber of the CO2 generated during the combustion of the biomass. Indeed, solids samples extracted from the reactor during this period revealed a high activity of the material (maximum carrying capacity of Xmax = 0.5). Experiments are in progress to reproduce these results with solid materials with lower CO2 carrying capacities (Xmax between 0.1-0.2), more representative of the CaO derived from natural limestones to be used in continuous cycling systems.

Since the purpose of this work is to illustrate the technical and economic viability of the proposed process under certain conditions, it is necessary to extrapolate from the available experimental information obtained at a small scale to the expected conditions in large scale systems. In previous work, a simple model for biomass combustion with in situ CO2 capture was proposed [16] and the fitting of this model to the new series of experimental data such as that of Figure 1 allows for the tuning of the main model parameter. An overall effectiveness factor of 0.74 in the combustor-carbonator reactor is obtained from this exercise, indicating that the carbonation rate of reaction is indeed the controlling step of the overall process of combustion and CO2 capture (the effectiveness factor would be 1 if the reactor was behaving as an ideal plug flow reactor for the gas, perfectly mixed for the solids and totally controlled by the intrinsic carbonation reaction rate). This also indicates that the rate of carbonation is sufficiently fast to achieve high capture rates of CO2 in beds formed by CaO particles with a certain active fraction (Xactive) reacting in the fast carbonation regime. This is illustrated in Figure 3, where the CO2 capture efficiency predicted by the model in a combustor-carbonator operating at conditions within the range of what is acceptable in existing large scale biomass boilers, assuming also an excess oxygen concentration in the combustor-carbonator of 3% (CO2 mass flow rate from combustion before capture is 8.0 mol/m2s at a combustion efficiency of 95%).

Solids inventory, kg/m2

Figure 2. Example of model predictions of CO2 capture efficiencies for different bed inventories in the biomass combustor-carbonator. T=700°C, superficial gas velocity=4m/s, solid circulation rate of CaO between risers= 5 kg/m2s.

The maximum carbonation efficiency allowed by the equilibrium is 84%, which it is represented by the horizontal dashed line. As it can be seen, the availability of highly reactive solids (as those present in the experiment of Figure 2, with Xmax around 0.5) would make it possible to achieve maximum efficiencies with very low bed inventories. This is however unrealistic for CaO derived solids from natural limestones, because the known tendency of the activity to decay as the number of times that the particles experience a carbonation-calcination cycles increases. Although there is intense research activity trying to develop better sorbents (see recent review of Blamey et al [17]), Figure 2 also shows that with modest values of sorbent activity (around 0.15) it is possible to achieve high CO2 capture efficiencies when the system operates with solid inventories well below 1000 kg/m2 (which are still within the range of what is possible in existing large scale CFB combustors). However, as indicated above, this sorbent activity is twice the 7-8% residual activity reported for CaO cycled many times and derived from natural limestones [18], and therefore, a substantial make up flow of fresh limestone (that displays a typical activity close to 70%) is still required in the system to operate with Xmax=0.15 (about 0.007 mol of fresh CaCO3/mol of CaO circulating between reactors). The method to solve the population balances that yield the relationship between average activity and make up flows is described elsewhere [19].

The previous analysis is used in the following paragraphs to carry out a case study for a large scale application of the combustor-carbonator reactor. It is implicitly assumed in all cases that the technology for continuous sorbent regeneration and CO2 production (through oxy-fired CFB combustion and calcination) is readily available from independent developments in other projects.

3. Case study

Figure 3 contains the main results from mass and energy balances applied to a new power plant implementing this concept. The temperature in the combustor-carbonator is 700 °C and the excess air ratio in the combustor-carbonator is 10%. A net power output from the full CFB system of 250 Mwe is assumed at a typical net power efficiency of 42% (before penalties due to the air separation unit for the calciner and CO2 purification and compression). In these conditions, the resulting thermal input to the system is 595.2 MWt. This heat input has to be split between the biomass heat input to the combustor-carbonator and the coal heat input to the combustor-calciner. To estimate this split in the heat balance, it is necessary to define key variables in the circulating system [13]. This is because heat requirements for calcination of the CaCO3 arriving to the calciner (fresh limestone make up flow and the calcium carbonate that comes from the combustor-carbonator) depend on the CO2 capture efficiency. Heat is also used in the calciner to heat up the incoming streams up to the calcination temperature. Therefore, in order to close the mass and energy balances, we assume that the CO2 capture efficiency in the combustor-carbonator is 80%, only slightly below the equilibrium maximum (84 %) following the discussion in previous sections. The molar carbonation conversion of CaO to CaCO3 is assumed to be 10% (Xcarb) while the maximum average carbonation conversion of the solids in the system is 15% (Xmax). The make up flow of fresh limestone account for 6.9 kg/s of fresh limestone (note that this is a substantial fraction of the mass flow of biomass to the combustor-carbonator (0.25), and therefore, the purge of solids should find a use a clinker precursor or as desulfurization feedstock for a larger coal power plant) .

Figure 3. Schematic overview of the mass and energy balances for a combustor-carbonator of 250 MWe. Biomass composition (41.39 w% C, 5.38 w% H, 39.90 w% O, 0.04 w% S, 0.71 w% N, 7.20 w% H2O, 5.38 %w ash). Coal composition (70.40 w% C, 4.90 w% H, 13.00 w% O, 1.30 w% N, 9.00 w% H2O, 1.40 %w ash)

Incorporating these assumptions into the mass and heat balances, the heat requirement in the calciner is calculated to be 54.4 % of the total heat input to the system. A sensitivity analysis on the effect of key operating variables on this figure reveals that it is relatively difficult to expect a drastic reduction in the heat requirements in the calciner (below 40%). However, since the calciner is operating at temperatures in excess of 900°C (950°C in the case study) it is important to stress that this heat input is effectively recovered for the steam cycle, extracting it from

the high temperature mass streams that leave the calciner and from the carbonation reaction heat (128.6 MWt) released in the combustor-carbonator at 700°C when the CO2 reacts with CaO. It is beyond the scope of this work to discuss the approach for the most effective integration of the steam cycle in this system. But it is clear that the necessary technology can be adapted from the existing CFB combustors, that operate with similar circulating solids and thermal conditions.

The similarities with existing and emerging CFB power systems also allows for an economic analysis of the system proposed, using the same methodology as in previous works for other Ca-looping processes [20, 21]. The target is to estimate the cost of electricity (COE) and the CO2 avoided costs with respect to a well defined reference case. In this work, the cost is estimated for a net power of 250 MWe to allow comparison with other technologies for CCS for biomass recently reviewed by the IEAGHG [15]. Table 1 shows the cost estimations for a coal reference plant (Coal Ref.) assumed to be supercritical CFB boiler [15], a biomass plant without CO2 capture (Bio.CFBC) of identical power output (based on the circulating fluidized bed technology and assuming a penalty respect to the reference case of 0.03 as justified in reference [15]), a coal oxy combustion plant based on circulating fluidized bed technology (Coal Oxy), a biomass oxy combustion plant based on circulating fluidized bed technology (Bio. Oxy) and the system described in this work (Bio. Ca-loop). Coal and biomass cost data, fixed fraction cost, fixed charge factor, and variable costs are common for all systems and taken from reference [15]. Also from this report are the specific capital cost of the first two reference systems (1400 €/kWe). We assume a capital cost for the oxyfired CFB power plant of 2500€/kWe and a reduced capital cost for the Bio. Ca-loop system of 1994 €/kWe. This is calculated assuming that the full system is made up of two components (one at an specific capital cost of 1400 corresponding to the 45.6% of the heat input coming from the combustor-carbonator and a second part at a cost of 2500 corresponding to the 54.4% of heat input of the oxy-fired calciner) [20]. Efficiency penalties for the new system are also calculated as proportional to the fraction of biomass used in the system, and the fraction of heat required in the calciner (that determines the size of the air separation unit to supply oxygen to the calciner).

In order to highlight the importance of regulation to promote this and other biomass-CCS systems we introduce as in reference [15] two usual values of carbon prices (14 €/tCO2 under the ETS market in Europe) and a green certificate (it is an additional benefit on the cost of electricity added to the proportion of electricity produced by the biomass),which is characteristic in several locations in Europe. Three cost scenarios can now be compared, as indicated at the bottom part of Table 1. As expected, the first scenario (no revenues) indicates that the cost of electricity from biomass based systems, with or without capture, is substantially higher than the cost of a coal based CCS system due to the much higher cost of the fuel. The system proposed in this work seems substantially more economic than the stand-alone biomass system, but this is mainly a distortion from the large contribution of cheaper coal as a fuel in the calciner. In order to change the relative situation between CCS technologies it is necessary to incorporate a certain carbon tax through the ETS value to the cost of electricity. When the ETS value is around 29 €/tCO2, the preferable technology would be the coal oxy combustion. If the ETS value increases at around 40€/tCO2 the Bio-Ca-loop system becomes economically attractive in order to exploit the negative emission component from the biomass firing in the combustor-carbonator. This tendency would also be followed by other biomass-based CCS systems as it is the case for biomass oxy combustion. In these scenarios of high ETS prices it would be increasingly more attractive to use biomass as a fuel in the calciner if this option is technically viable, even if this translated into slightly lower capacity factors and higher capital cost. Finally, the last line in the table shows the impact of both an ETS market and a certain green certificate for the MWh generated from biomass fuel. When de green certificate value is less than 30€/MWh, the preferable technology is the system proposed in this work as the cost of electricity and the avoided cost of CO2 are the lowest. For the green certificate chosen in reference [15], 50€/MWh, the avoided cost of CO2 as well as the cost of electricity are the lowest with a stand-alone biomass plant without capture. Under these high incentives conditions for biomass electricity, the proposed system in this work would be competitive with higher cost of carbon under the ETS and/or a larger share of biomass use by operating the calciner with biomass and not coal.

Table 1: Cost estimations for the system and comparison with other technologies.

1 Coal Ref. Bio. CFBC Coal Oxy. Bio. Oxy Bio.Ca-loop

Biomass cost, FC €/kWh 0.03 0.03 0.03

Coal cost, FC €/kWh 0.01 0.01 0.01

Capital cost, TCR €/kWe 1400 1400 2500 2500 1994

Fixed fraction cost, FOM % (CC) 3.7 3.7 3.7 3.7 3.7

Capacity factor, CF % 90 88 90 88 88

Fixed charge factor, FCF 0.1 0.1 0.1 0.1 0.1

Variable cost, VOM €/kWhe 0.01 0.01 0.01 0.01 0.01

Carbonator capture efficiency, Ecarb 0.8

Capture efficiency, Eoxy 0.95 0.95 0.95

Overall capture efficiency, E 0.95 0.88

Power fraction non oxycomb 1 1 0 0 0.46

Penalty ASU 0.07 0.07 0.04

Penalty Compresion 0.05 0.05 0.05

Penalty Biomass 0.03 0 0.03 0.01

Penalty total 0.03 0.12 0.15 0.10

Net efficiency, h kWhe/kWh 0.45 0.42 0.33 0.30 0.35

emission factor kg CO2/MWhe 880 0 60 -1463 -457

ETS price €/t CO2 14 14 14 14 14

Green certificate €/MWh 50 50 50 50

Cost of electricity without revenues, COE €/kWhe 0.051 0.100 0.073 0.144 0.092

Avoided cost without revenues, AC €/t CO2 avo. 56 27 40 31

COE with ETS only €/kWhe 0.063 0.100 0.074 0.123 0.086

Avoided cost with ETS only €/t CO2 avo. 42 13 26 17

Cost of electricity with revenues, COE €/kWhe 0.063 0.050 0.074 0.073 0.063

Avoided cost with revenues, AC €/t CO2 avo. -14 13 4 0

4. Conclusions

The concept of in situ CO2 capture during low temperature biomass combustion seems experimentally feasible when operating at around 700°C to maximize both combustion and CO2 capture efficiencies in circulating fluidized beds fed with a continuous supply of CaO from a calciner. New experimental work in continuous mode of operation in a 30 kW interconnected fluidized bed test facility confirms the need to guarantee a high sorbent solids inventory of active CaO in the combustor-carbonator reactor, as well as an intense circulation of solids between this reactor and the calciner. This leads to CO2 capture efficiencies over 80%, remarkably close to those allowed by the equilibrium and the combustion mass balance. The similarity of the process conditions with respect to existing commercial CFB combustors for biomass, coal and emerging oxy-combustion CFB systems allows for an estimation of the main cost figures in the process. This reveals that the concept can only be competitive respect to other CCS systems if there is an incentive (green certificate) for electricity generated from biomass and/or a sufficient price for carbon emissions.

5. Acknowledgments

This work has been carried out thanks to the financial support from Gas Natural Fenosa, within the CENITCO2 and MENOSCO2 projects. M. Alonso, N. Rodriguez also acknowledge their fellowships received through the JAE-Doc Programme from CSIC and FICYT.

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