Calcium Looping Cycle for Hydrogen Production ... - ACS Publications

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Calcium Looping Cycle for Hydrogen Production from Biomass Gasification Syngas: Experimental Investigation at a 20 kWth Dual Fluidized-Bed Facility Nina Armbrust,* Glykeria Duelli (Varela), Heiko Dieter, and Günter Scheffknecht Institute of Combustion and Power Plant Technology (IFK), University of Stuttgart, Pfaffenwaldring 23, 70569 Stuttgart, Germany ABSTRACT: The use of a calcium looping cycle for the production of hydrogen (Ca-LHP) from biomass gasification syngas is a promising alternative to the conventionally used catalytic water−gas shift process. The process takes place in two fluidized bed reactors connected by solid CaO flow. In the first reactor (the carbonator), the water−gas shift reaction is enhanced by CO2 capture via the carbonation reaction with CaO. In the second reactor (the regenerator), the captured CO2 is released. This paper presents results from a continuously operated 20 kWth dual fluidized-bed facility using a syngas-like mixture obtained from biomass gasification. The influence of the main process parameters, i.e., temperature in the carbonator (Tcarb), looping ratio (LR), and space time (τ) on the H2 concentration, and the CO conversion in the carbonator was investigated. A H2 concentration of 84.4 vol %dry and a CO conversion of 76.4% were achieved for a carbonator temperature of 639 °C, a looping ratio of 6 molCa/molCO+CO2, and a space time of 1.0 h. Additionally, experiments with different syngas compositions as obtained from various biomass fluidized bed gasification processes were performed and the mechanical stability as well as the chemical sorbent activity were studied. The potential of the calcium looping cycle for H2 production could be shown by the experimental results obtained and their comparison with literature.

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INTRODUCTION Renewable and sustainable production of hydrogen for industrial and chemical application or as a clean energy carrier is an important topic for the future. Biomass is considered as one of the most promising sources for carbon-neutral H2 production, because it is abundant, environmentally friendly, and renewable.1,2 Thermochemical gasification coupled with catalytic water−gas shift reaction (WGSR), followed by a CO2 separation and a purification step is the most widely practiced and economically viable process route for the conversion of biomass into hydrogen.3−5 An alternative approach to the conventional used catalytic WGS process is the use of a calcium looping cycle where the reversible carbonation reaction of CaO is used to capture CO2 (eq 1) and thereby enhance the WGSR (eq 2) at temperatures of 600−700 °C.

(iii) It simplifies the biomass-to-hydrogen process by integrating the WGSR and CO2 separation step in one reactor. (iv) The sulfur-sensitive shift catalyst is eliminated. This fact may become increasingly important in the future as organic residues with higher sulfur content might be used for biomass gasification. The Ca-looping cycle can either be integrated into the gasification process, as done, for example, in the CO2 Acceptor Gasification7 and Sorption Enhanced Reforming (SER)8−14 process, or the Ca-looping cycle can take place in a secondary reactor system, following the gasification step.15−18 Compared to SER gasification, the following benefits can be expected from a subsequent Ca-looping cycle: (i) A higher CO2 capture efficiency can be reached, because the syngas produced in the gasifier is passing the entire bed material of the carbonator. During SER gasification, the syngas is formed in the CaO bed or the gasification zone might be even located in the upper part of the gasifier bed, as a result of bed segregation.19 (ii) The temperature in the gasifier can be increased, since it is not restricted to 600−700 °C because of the carbonation reaction; thereby, the biomass conversion is increased and the tar yield is decreased. The first laboratory-scale experimental studies with a syngaslike gas mixture using a Ca-looping cycle as a subsequent step after gasification were performed by Han and Harrison15,20 in a

CaO(s) + CO2(g) ↔ CaCO3(s) ΔH25 ° C = ∓178.2 kJ/mol

(1)

CO(g) + H 2O(g) ↔ CO2(g) + H 2(g) ΔH25 ° C = ∓40.9 kJ/mol

(2)

This concept has a number of advantages over current used catalytic WGS process: (i) It uses limestone, a naturally occurring, widely available, relatively inexpensive, and nontoxic sorbent. (ii) The entire process occurs at elevated temperature. Therefore, the requirement of heat exchangers is reduced. Furthermore, high-quality heat can be extracted from exothermic carbonation reaction and from regenerator flue gas, offering the possibility to improve the overall efficiency of the process.6 © 2015 American Chemical Society

Received: Revised: Accepted: Published: 5624

January 8, 2015 May 2, 2015 May 5, 2015 May 5, 2015 DOI: 10.1021/acs.iecr.5b00070 Ind. Eng. Chem. Res. 2015, 54, 5624−5634

Article

Industrial & Engineering Chemistry Research

WGSR. Beside the H2 concentration, the CO conversion rate (eq 3) can be used as an indicator for the performance of the WGSR.

pressurized packed-bed reactor. This study was later extended by Müller et al.,17 using a synthetic and natural sorbent. Ramkumar and Fan21,22 performed extended thermodynamic analysis of a Ca-looping cycle for syngas obtained from different coal gasification processes, as well as batch experiments in a fixed-bed reactor using a mixture of N2, steam, and CO as feed gas. This work was then continued by Phalak et al.,18 who used a fluidized-bed reactor and Ca(OH)2 for semibatch experiments using H2O, CO, and N2, as well as a syngas-like mixture (H2O, CO, H2, CO2, and N2). Symonds et al. performed experiments in a thermogravimetric analyzer (TGA)23 and batch experiments in a dual fluidized bed under different calcination and carbonation conditions,24 also using a syngas-like mixture. The experiments performed showed that the Ca-looping cycle has the potential to achieve high CO conversion and produce a high purity of H2, making it possible to increase the efficiency of H2 production from coal-derived syngas.6 The investigations performed until now mainly focused on syngas composition derived from coal gasification and were performed either in small-scale reactors or noncontinuous mode. This work, for the first time, presents experiments conducted in a continuously operated 20 kWth dual fluidizedbed facility at the University of Stuttgart, investigating the Ca-looping cycle utilized for H2 production from syngas derived from various fluidized-bed biomass gasification processes. The main process parameters varied during this investigation were the carbonator temperature (Tcarb) and space time (τ), as well as the looping ratio (LR).

+ − − FCO FCO + FCO

XCO =

(3)

The efficiency of the carbonation reaction, which is the driving force of the WGSR, is described by the CO2 capture efficiency, which is defined as follows, including the CO2 fed into the carbonator and the CO2 formed in the carbonator via the WGSR: ECO2 =

+ + − FCO + XCOFCO − FCO 2 2 + + FCO + XCOFCO 2

(4)

Figure 1 shows that a molar flow of particles FCa with a carbonate content of Xcalc is entering the carbonator and, after capturing some of the CO2, is leaving the carbonator with a carbonate content of Xcarb. Making the assumption that no other reactions except the WGSR and the carbonation reaction are taking place in the carbonator and neglecting the makeup flow, the amount of CO2 captured in the carbonator can be described by the carbon molar balance as follows: + + ECO2(FCO + X COFCO ) = FCa(Xcarb − Xcalc) 2

(5)

The looping ratio (LR) for this process can be described as the ratio of the number of moles of calcium (FCa) coming to the carbonator and the number of moles of CO2 (F+CO2) and CO (F+CO) entering the carbonator per hour (see eq 6). The space time (τ) is defined as the ratio of moles CaO (nCa) in the carbonator and moles of CO2 and CO per hour in the syngas fed into the carbonator (eq 7).

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PROCESS DESCRIPTION AND THEORETICAL BACKGROUND The scheme of the Ca-looping cycle for hydrogen production (Ca-LHP) is shown in Figure 1; this scheme was first proposed

LR =

τ=

+ FCO

+ FCO

FCa + + FCO 2

(6)

nCa + + FCO 2

(7)

Another important parameter influencing the carbonation reactionand, thus, the H2 productionis the performance of the limestone. A measure that characterizes the limestone chemical activity is the average maximum carbonation conversion (Xmax,ave), which is defined as the maximum carbonation conversion that can be achieved by the average solid in the carbonator at the end of the fast reaction regime. Xmax,ave will be used within this investigation to describe the chemical property of the sorbent and is measured in a TGA. As reported in the literature by various authors,26−29 the average maximum carbonation conversion decreases with the amount of calcination/carbonation cycles that the limestone has experienced until it reaches its residual activity. In order to study the evolution of Xmax,ave during experimental investigation, the theoretical number of cycles (Nth) is used, which is defined by Charistos et al.30 as the amount of times the number of moles of CO2 captured could carbonate the total bed inventory (nCa,total) up to its average maximum carbonation conversion Xmax,ave (see eq 8).

Figure 1. Scheme of the Ca-looping cycle for hydrogen production (Ca-LHP).

by Shimizu et al.25 for post-combustion Ca-looping. It involves two fluidized-bed reactors, a carbonator, and a regenerator connected by solid transport flow. The first reactor, the carbonator, is fluidized with syngas derived from biomass gasification and is operating in a temperature range of 600−700 °C. In this reactor, in situ adsorption of carbon dioxide takes place via the CaO sorbent through the carbonation reaction (eq 1). Because of CO2 adsorption, the equilibrium of the WGSR (eq 2) is shifted to the hydrogen product. In the second reactor, the regenerator, calcination of the sorbent takes place. The goal of the process is to produce a large amount of H2. Therefore, as much as possible of the CO and steam fed into the carbonator should be converted to H2 and CO2 via the

Nth = 5625

∫0

t

+ + ) + XCOFCO ECO2(FCO 2

nCa,total X max,ave

dt (8)

DOI: 10.1021/acs.iecr.5b00070 Ind. Eng. Chem. Res. 2015, 54, 5624−5634

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EXPERIMENTAL SECTION

Table 1. Chemical Composition of the Limestone

Experiments were conducted in the IFK DFB facility at the University of Stuttgart, which consists of a 12.4-m-high, 71-mm-diameter circulating fluidized bed (CFB) riser and a 150-mm-diameter bubbling fluidized bed (BFB). The facility is described in more detail in ref 31 and is shown in Figure 2. For

component

amount

CaO SiO2 MgO Al2O3 Fe2O3 Na2O TiO2 CO2

53.64 3.51 0.51 0.50 0.18 0.02 0.02 41.62

All experimental points presented within this paper are derived from steady-state operation, which is defined as a period of time lasting between 15 min and 1 h where temperatures, pressure drops, and inlet and outlet gas flows, as well as the circulation rate, remain constant. There was no continuous make-up flow of lime to the system, but fresh calcined lime had to be added to the facility from time to time to compensate for losses due to attrition, solid sampling, and possible cyclone inefficiency. Mass distribution in the system was calculated by pressure profiles provided from the pressure traducers. The outlet gas composition of the carbonator was measured by an ABB Advance Optima 2020 system. The circulation rate between the two reactors was measured by closing a butterfly valve in the return leg after the cone valve and measuring the time that the solids require to accumulate up to a specific height. After each steady state, solid samples were collected from both loop seals. The carbonate content (Xcarb/calc) of the samples was analyzed by means of a TGA 701 by LECO. The average maximum carbonation conversion Xmax,ave was measured for an inlet CO2 concentration of 13 vol % balanced by N2 and carbonation temperature of 640 °C using a TG analyzer developed by the University of Stuttgart. A moredetailed description of the Xmax,ave measurement performed can be found in ref 32. The particle size distribution of the samples was analyzed using a Malvern Mastersizer 3000 system. In the first part of the experimental investigation, the influence of three different operation parameters, namely, temperature (Tcarb), looping ratio (LR), and space time (τ) on the carbonator outlet syngas composition, the CO2 capture efficiency (ECO2) and the CO conversion (XCO) are studied. Experimental conditions are summarized in Table 2. During

Figure 2. Scheme of the 20 kWth IFK DFB facility at the University of Stuttgart.

experiments, the BFB was used as the carbonator and the CFB was used as the regenerator, because, at this facility, the experimental conditions are much easier to control in the BFB reactor. The mass flow between the two reactors is controlled by a cone valve. The entire facility is electrically heated and LabView software program is used for control of the facility and data acquisition. Syngas composition was adjusted by mass flow controllers (MFCs) and added to the carbonator via a gas mixing station. Steam was provided by a steam generator and fed into the distributor of the carbonator. Gas flows to the regenerator and loop seals were also controlled by MFCs. The chemical composition of the limestone used for the experiments is shown in Table 1. The limestone was completely calcined before the experiments. It was fed into the facility, which was heated to 600−700 °C in the carbonator and 850−900 °C in the regenerator. After reaching hydrodynamic stability, mixed syngas consisting of H2, CO, CO2, CH4, and N2 was fed into the carbonator, together with steam. The equilibrium gas concentrations were calculated using FactSage 6.2 software. Except for CH4, there were no further noncondensable hydrocarbons or tars present during experiments. The gas at the outlet of the carbonator was burned by a flare. Regenerator and loop seals were fluidized with N2.

Table 2. Basic Operation Conditions during Experiments value parameter temperature velocity solid circulation rate looping ratio, LR space time, τ

carbonator

regenerator

600−700 °C 800−920 °C 0.3−0.7 m/s 4−5 m/s 0.2−2.3 kg/(m2 s) 2−20 0.4−1.0 h

this investigation, the gas composition is kept constant; this is shown in Table 3 as the base case. Allothermal steam gasification with CaO as the bed material was chosen as the base case, because of its high H2 content, low CH4 content, and low tar content. In addition, the mechanical stability and the chemical activity of the limestone are studied. In the second part, the performance of the Ca-LHP for different fluidized-bed gasification processes is evaluated. Therefore, operation conditions are kept constant and syngas 5626


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Table 3. Base Case Gas Composition Used for Parameter Variation (Part 1) and Syngas Compositions Obtained from Different Ambient Fluidized Bed Biomass Gasification Processes A−D Used for Syngas Variation (Part 2)

gasification agent fuel bed material Tgasifier (°C) GR or S/B (kg/kgdaf)d H2 (vol %dry) CO (vol %dry) CO2 (vol %dry) N2 (vol %dry) CH4 + CxHy (vol %dry) H2O (vol %) H2O/CO (mol/mol) a

base case

A

B

lab-scale BFB facility at the University of Stuttgart steam wood pellets CaO 850 0.8

40 kWth BFBa

100 kWth CFBb

steam/oxygen pine sand 810−830 1.1

steam/oxygen wood sand 837 1.7

51.4 20.9 23.4

27.0 43.5 20.0

21.9 34.0 33.2

4.3

9.5

40.0 3.2

40.1 3.0

C

D

8 MWth DFB gasifier in Güssingc steam wood chips olivine 850−900 n.a.

20 kWth DFB facility at University of Stuttgart steam wood pellets CaO 800 1.0 55.9 12.9 23.2

10.9

41.0 25.0 20.0 2.0 12.0

70.7 7.1

39.7 2.6

40.5 5.3

8.0

Data taken from ref 33. bData taken from ref 34. cData taken from ref 35. dThe subscript “daf” denotes dry ash free.

compositions obtained from different fluidized-bed biomass gasification processes, taken from the literature33−35 and our own previous experiments, are fed into the carbonator. Fluidized-bed gasification processes using steam (C+D) and a mixture of steam and oxygen (A+B) as gasification agent were chosen because they are most commonly used for biomass gasification and for their high potential for hydrogen production by producing syngas with low nitrogen content. Further gasification processes using different bed materials (sand (A+B), olivine (C), and CaO (D)) were selected to guarantee a wide variety of syngas compositions. Table 3 summarizes the gas composition as well as the basic operation parameters of the gasification processes.

shown by comparing H2,eq with the equilibrium H2 concentration (H2,eq 0) reached in a reactor at the same temperature without a CaO bed, which is calculated to 56.3 vol %dry and is only 2.7 vol % above the feed gas H2 concentration. Furthermore, this steady state proves that a H2 concentration of >80 vol %dry can be reached by a continuously operating dual fluidized-bed system for a long time. As described in the previous section, it is assumed for the carbon molar balance that no other reaction beside the carbonation reaction and WGSR are taking place in the carbonator. If this assumption proves correct, H2 is only produced and CO consumed via the WGSR and, therefore, the number of moles of CO converted in the carbonator (|ΔFCO|) should be equal to the number of moles of H2 produced (|ΔFH2|). Figure 4 shows these two values for all experimental

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RESULTS AND DISCUSSION In Figure 3, the dry inlet (dashed lines) and outlet (solid lines) gas concentrations of H2, CO, and CO2 for an example of a 1 h

Figure 3. Example of a 1 h carbonator steady state, Tcarb = 645 °C, LR = 14 molCa/molCO+CO2, τ = 0.8 h, and H2O/CO = 3.5 molH2O/molCO.

Figure 4. Flow rate of CO (|ΔFCO|, molCO/h) converted in the carbonator plotted against the flow rate of hydrogen (|ΔFH2|, molH2/h) formed in the carbonator for all experimental points conducted (feed gas composition: base case + (A−D)).

carbonator steady state are shown. CH4 concentration is not depicted and is balancing the other gas concentrations to 100 vol %dry. During this steady state, all process parameters are kept constant. It can be seen that, at the carbonator exit, the syngas contains 82.4 vol %dry of H2, which is close to equilibrium H2,eq concentration of 88.0 vol %dry, and less than 4 vol %dry CO2. The potential advantage of the Ca-LHP can be

steady states conducted during the investigation. Since the majority of all experimental points are very close to the 45° line, and small deviations must be accepted, because of deviation of the mass flow controller and measurement uncertainties, it can be stated from Figure 4 that no other reactions except the WGSR have a significant influence on the H2 production and CO consumption in the carbonator. 5627


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Figure 5. Carbon molar balance: CO2 adsorbed from solids (molCO2/h) versus CO2 removed from syngas (molCO2/h) for all experimental points conducted (feed gas composition: base case + (A−D)).

Figure 7. CO conversion (XCO) and CO2 capture efficiency (ECO2) measured and at equilibrium, versus the carbonator temperature (Tcarb) for τ = 0.7 h and a LR = 7 molCa/molCO+CO2.

Figure 5 is a proof of the quality of the obtained experimental data and is based on the carbon molar balance, described by eq 5. The number of moles of CO2 captured in the carbonator are calculated from the gas phase and plotted against the number of moles of CO2 derived from solid sample analysis. Both measurements are independent of each other and it can be seen that the majority of points are close to the 45° line, demonstrating a high accuracy of the analysis performed. Part 1: Parameter Variation. Effect of Temperature. In order to study the influence of the temperature in the carbonator, Tcarb was increased from 606 °C to 637 and 697 °C. The gas composition obtained for every steady state is plotted in Figure 6. The LR and τ values were kept constant for these

These results are in good agreement with those found by other authors17,22 for their laboratory-scale fixed-bed reactor investigations. As described by Müller et al.,17 an increase in the temperature is accelerating the kinetics of the WGSR but, at the same time, decreases the equilibrium level. Furthermore, they found that the increase in kinetics is dominating over the decrease in the equilibrium concentration. These findings are confirmed by the results presented here, showing at Tcarb = 697 °C, a XCO and, therefore, H2 concentration slightly above those at Tcarb = 606 °C. For the temperature range investigated, ECO2 is close to or even above 90%. This is a relatively high CO2 capture efficiency, taking into account that some of the CO2 might be produced in the freeboard of the carbonator, where it cannot be captured by CaO. This fact can also be seen by the low wet CO2 outlet concentration, which is only ∼2 vol %, as the syngas is containing ∼40 vol % steam at the exit of the carbonator. XCO, on the other hand, remains, to some extent, away from equilibrium, indicating that the kinetics of the WGSR are relatively slow without an appropriate catalyst present in the temperature range investigated. This fact also has been described by other authors.17,36,37 CaO is catalyzing the WGSR to some extent,38,39 but seems to be less effective than other catalysts in a temperature range of ∼650 °C.17 Effect of Looping Ratio. In order to investigate the influence of the calcium looping ratio (LR), it was varied by adjusting the solid flow from the regenerator to the carbonator by opening and closing the cone valve. Tcarb was kept constant at 640 ± 4 °C and τ at 0.7 h during LR variation. As shown in Figure 8, the H2 concentration is increased from 78 vol % to 82 vol %dry and the CO2 capture efficiency is increased from 88% to 92% by increasing the LR from ∼2 molCa/molCO+CO2 to 7 molCa/molCO+CO2. A further increase of the LR above 7 molCa/molCO+CO2 does only result in slightly increasing ECO2 and H2 concentration. These results can be explained by the decreasing carbonated CaO fraction Xcarb in the carbonator with increasing LR. This leads to an increase of the CaO fraction available in the bed for carbonation, which is an important driving force for the carbonation reaction. Nevertheless, the effect of the LR seems to be limited by the CO2 partial pressure, which is reaching values close to equilibrium very quickly, indicating fast kinetics of the carbonation, even at low LR. These fast kinetics can be explained by two facts: (i) a high average maximum carbonation conversion Xmax,ave due to presence of water vapor,40,41 which will be discussed in more detail

Figure 6. Gas composition versus carbonator temperature (Tcarb) for τ = 0.7 h and a LR = 7 molCa/molCO+CO2.

steady states at a medium value of 7 molCa/molCO+CO2 and 0.7 h. It can be seen that H2 concentration is enhanced significantly in the carbonator, and the highest concentration of 82 vol %dry was measured at 637 °C while CO and CO2 concentration reach the lowest concentration measured. The CH4 concentration is not affected by the temperature in the carbonator. Figure 7 shows the CO2 capture efficiency (ECO2), the CO conversion (XCO) as well as the calculated equilibrium CO2 capture efficiency (ECO2,eq) and CO conversion (XCO,eq) for every steady state. Hence, the thermodynamic equilibrium of the combined reaction is predicting highest equilibrium CO2 capture efficiency (ECO2,eq) and CO conversion (XCO,eq) for the lowest temperature investigated (Tcarb = 606 °C), from Figure 7, it can be seen that highest ECO2 of 92.0% and XCO of 63.9% are reached during experimental investigation for a temperature of 637 °C. 5628


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Figure 10. Equilibrium-normalized CO2 capture efficiency (ECO2/ECO2,eq) versus looping ratio (LR) for Tcarb = 639 ± 4 °C and varying τ.

Figure 8. H2 concentration and CO2 capture efficiency (ECO2) versus looping ratio (LR) for Tcarb = 640 ± 4 °C and τ = 0.7 h.

later, and (ii) a high local CO2 partial pressure as the WGSR is catalyzed on the CaO surface.23 The impact of changing LR on the XCO is shown in Figure 9. Besides the variation of the LR at 640 °C, the effect was studied

Figure 11. Equilibrium-normalized CO conversion (XCO/XCO,eq) versus looping ratio (LR) for Tcarb = 639 ± 4 °C and varying τ.

reaching values close to equilibrium even at small LR. XCO/XCO,eq, on the other hand, is increasing with increasing τ for a constant LR, as shown in Figure 11. That increase in XCO can be explained by the gas residence time in the bed. By increasing τ, for a constant gas composition and constant molar inlet flow of CO2 and CO, the gas residence time in bed is increased, extending the time the combined WGS and carbonation reaction can occur. These outcomes are in good agreement with those of Müller et al.,17 who found, for variation of the amount of CaO sorbent in a fixed-bed reactor, that the carbonation reaction is reaching thermodynamic equilibrium even for a short contact time while the amount of H2 produced is increased with increasing mass of sorbent. Furthermore, it can be seen that an increase in τ, from 0.7 h to 1.0 h, has a more significant effect on XCO/XCO,eq than the increase from 0.4 h to 0.7 h, while the influence of LR is decreasing with increasing τ. This fact might be explained by an improved gas solid contact at τ = 0.4 h, as a result of higher fluidization velocity. The XCO value closest to equilibrium of 78.7% is reached for τ = 1.0 h and LR = 12 molCa/molCO+CO2, achieving a H2 concentration of 85.4 vol %dry. Although it would not be of commercial interest for process realization, a XCO even closer to XCO,eq might be achievable by increasing τ further. Mechanical Stability and Chemical Sorbent Performance. Attrition of the sorbent is studied by comparing the particle size distribution (PSD) of the solid samples taken and evaluation of the material losses during operation. Figure 12 shows the results of the PSD analysis of solid samples taken after

Figure 9. CO conversion (XCO) versus looping ratio (LR) for varying Tcarb and τ = 0.7 h.

at carbonator temperatures of 604 ± 4 °C and 700 ± 5 °C. The LR has the same influence on XCO for all three temperatures investigated; it is increasing with increasing LR. Furthermore, the diagram confirms the trend shown in the previous section that, for a constant LR, the highest XCO can be reached at Tcarb = 640 ± 4 °C. Effect of Space Time. The third parameter studied in this paper is the space time (τ). For this study, steady states for three different τ values, namely, 0.4, 0.7, and 1.0 h, are conducted for different LR values and a carbonator temperature of 639 ± 4 °C. τ is altered by varying the amount of gas fed into the carbonator and thereby the carbonator velocity from 0.34 m/s to 0.68 m/s. Even though higher gas velocity would be more realistic, high space time values of 0.7 and 1.0 h can only be reached at these relatively low gas velocities as the amount of solids in the carbonator is fixed, because of the overflow of the reactor. Figure 10 shows the equilibriumnormalized CO2 capture efficiency (ECO2/ECO2,eq), and Figure 11 shows the equilibrium-normalized CO conversion (XCO/XCO,eq) for the variation of τ. The equilibrium-normalized values are shown, making it possible to quantify the performance of the WGS and carbonation reaction, in comparison to the maximum allowed by equilibrium. As shown in Figure 10, ECO2/ECO2,eq is not affected by τ and is, as shown in the previous section, 5629


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experiments, as well as the theoretical decay of Xmax,ave for a “average” limestone defined by Grasa and Abanades.26 In addition, the Xmax,ave of the calcined limestone is shown. After the precalcination step, the Xmax,ave with 0.12 molCaCO3/molCa is relatively low, which can be explained by thermal stress during the precalcination step where the limestone is exposed to high temperature (∼900 °C) for several hours.30,45 After a few cycles, Xmax,ave is increasing to values between 0.16 and 0.29 molCaCO3/molCaO and remained almost constant for up to 20 theoretical cycles. A rapid decay, as per the theoretical curve from Grasa and Abanades,26 is not recorded. Symonds et al.23 has reported that the presence of H2 and CO has no impact on the Xmax,ave. One possible explanation can be the presence of water vapor, which seems to cause an improvement of the limestone activity by changing the microstructure of the particles, as reported by many authors.24,32,40,41 Duelli et al.32 found during the experimentation performed, using the same limestone, that the residual Xmax,ave almost doubles from 0.10−0.15 molCaCO3/molCaO up to 0.20 molCaCO3/molCaO when water vapor is present during carbonation or calcination. Part 2: Syngas Variation. During variation of the syngas composition, Tcarb was set to 640 °C, τ to 1.0 h, and LR to 9 molCa/molCO+CO2. These values were chosen due to promising results, which have been achieved during parameter variation. Figures 14−17 show the dry feed gas composition at the carbonator inlet (dashed lines), the dry product gas composition measured at the carbonator outlet and the CO conversion (solid lines), as well as the equilibrium hydrogen concentration H2,eq and equilibrium CO conversion (XCO,eq)

Figure 12. Cumulative particle size distribution of solid samples after calcination, after 8 h of operation, and after 15 h of operation.

precalcination, after 8 h of operation, and after 15 h of operation. This figure shows that the mean particle size (dp50) is reduced from 402 μm to 356 μm after 8 h of operation and further reduced to 300 μm after 15 h of operation. Material losses were determined by collecting the material lost to cyclones and filters. It was found to be 1.7 wt % of the total system inventory per hour of operation, which is not a significant problem during experiments. However, it must be taken into account that precalcined limestone was used for the experiments. Scala et al.42 classified attrition phenomena in three categories: primary fragmentation, attrition by abrasion, and secondary fragmentation. Primary fragmentation occurring during the first calcination is reported in the literature to have a significant influence,43 and the particles entrained from the BFB during the precalcination step are not included in the attrition calculation. During continuous operation, attrition contributes to abrasion and secondary fragmentation as primary fragmentation becomes less significant. The attrition rate measured during the experiments are in good agreement with those measured by Charitos et al.,31 using the same facility (same configuration BFB carbonator/CFB regenerator) but a different limestone for post-combustion CO2 capture experiments. Duelli et al.44 found attrition with 0.8 wt %/h to be significantly lower, using the same limestone but different facility configuration (CFB carbonator/BFB regenerator). This lower attrition rate can be explained by lower mechanical stresses due to lower velocities during BFB regeneration and low calcination rates, as well as mechanical strengthening of the particles, because of the high CO2 concentration present in the regenerator. Chemical sorbent activity during the experimental investigation is evaluated by measuring the average maximum carbonation conversion (Xmax,ave). Figure 13 shows the results of these measurements for all solid samples taken during

Figure 14. Inlet and outlet gas concentration of feed gas composition A obtained from steam−oxygen BFB gasification. Conditions: H2O/CO = 3.0 molH2O/molCO, Tcarb = 642 °C, LR = 8.6 molCa/molCO+CO2, τ = 1.1 h.

Figure 15. Inlet and outlet gas concentration of feed gas composition B obtained from steam−oxygen CFB gasification. Conditions: H2O/CO = 7.1 molH2O/molCO, Tcarb = 643 °C, LR= 10.6 molCa/molCO+CO2, τ = 1.1 h.

Figure 13. Average maximum carbonation conversion (Xmax,ave). 5630


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Furthermore, the results show that the Ca-looping cycle can achieve XCO of over 70% for all syngas composition studied and is able to generate high H2 concentration, even for low H2 feed gas concentrations. Comparing the H2 concentration reached during experimental investigation with those obtained during SER gasification,13,46,47 it can be seen that conventional steam gasification (feed gas composition C, Figure 16) coupled with a Ca-looping cycle can reach H2 concentration in the same range as during SER gasification. Advantageous of the SER gasification is the simpler overall reactor design, as gasification and CO2 adsorption are taking place in one single reactor. Therefore, the heat released by the exothermic carbonation reaction is directly used for endothermic gasification reaction. However, the S/C ratio typically used in SER gasification is up to twice that used in conventional steam gasification. When a CaO bed is used in the gasifier, higher H2 concentration can be reached at the gasifier exit. The CaO bed is catalyzing the WGSR in the gasifier, leading to higher CO conversion and, therefore, H2 concentration during gasification.17,39 When a gasifier using a CaO bed (feed gas composition D, Figure 17) is coupled with a Ca-looping cycle, higher H2 concentrations can be reached, compared to those obtained in SER gasification, as a result of the improved CO2 capture efficiency. Furthermore, the biomass conversion will be increased and the tar yield decreased, as a result of the higher gasification temperature. This fact will become even more important when organic residues, such as, for example, sewage sludge or manure, are used for gasification.48−50 The CaO in the gasifier will also have a positive effect on the ash melting point,51 which might become a problem when organic residues are gasified under high temperature. Using CaO as the bed material in the gasifier further offers the possibility that regeneration of the sorbent could take place in the combustion part of the gasification process, simplifying the process concept52 and decreasing the overall energy requirement. Conventionally used catalytic WGS reactors are usually divided into two or more stages, including high-temperature shift (HTS) and low-temperature shift (LTS) reactor. In these reactor systems, a CO conversion of >90% can be achieved.53 Corella et al.54 and Aznar et al.55 published results from hightemperature and low-temperature CO-shift reactor downstream from a steam−oxygen-blown fluidized-bed gasifier and a steam reforming catalytic bed. During their experiments, they could reach up to 90% CO conversion and a H2 concentration of 71−73 vol %dry. Results shown in Figure 14 demonstrate that, with a comparable gas composition (see Table 3, feed gas composition A), the Ca-looping cycle could reach up to 80% CO conversion and a H2 concentration of 72 vol %dry. Fail et al.56 published experimental results of 100 h of operation of a WGS unit using a side stream from the CHP Plant in Oberwart, Austria. During this operation, 72% of the CO was converted in the WGS unit, reaching a H2 concentration of 50 vol %dry, which was further increased to 76.1 vol %dry by a membrane separation unit. Additional steam had to be added to prevent coking at the surface of the WGS catalyst. With feed gas composition C, obtained from CHP in Güssing, producing similar gas composition, the Ca-LHP can reach a value of XCO = 74.8% and a H2 concentration of 74.5 vol %dry (Figure 16). By comparing the results obtained during the experimental investigation shown in this paper with those from literature, it becomes clear that the Ca-LHP can produce comparable H2 purity and CO conversion by combining the two staged WGS reactors (HTS and LTS) and CO2 removal in one single

Figure 16. Inlet and outlet gas concentration of feed gas composition C obtained from steam gasification in Güssing. Conditions: H2O/CO = 2.6 molH2O/molCO, Tcarb = 642 °C, LR = 8.7 molCa/molCO+CO2, τ = 1.1 h.

Figure 17. Inlet and outlet gas concentration of feed gas composition D obtained from steam gasification with CaO bed. Conditions: H2O/CO = 5.3 molH2O/molCO, Tcarb = 643 °C, LR = 8.6 molCa/molCO+CO2, τ = 1.0 h.

(dotted lines) for the different feed gas compositions A−D (see Table 3). The CH4 inlet and outlet concentrations are balancing the other concentrations to 100 vol %dry. It can be seen that the highest H2 concentration of 83 vol %dry is reached for gas composition D obtained from steam gasification with CaO bed. For this product gas composition also, the lowest CO2 concentration of 2.8 vol %dry is achieved. The highest XCO of 84.4% is achieved for feed gas composition B obtained from CFB steam-oxygen gasification and is shown in Figure 15. For this steady state, the LR was 11 molCa/molCO+CO2, which is slightly above the set point. However, recall that, as stated in a previous section, an increase of the LR above 7 molCa/molCO+CO2 only slight increases the CO conversion. For this gas composition also, the highest equilibrium CO conversion was calculated. Feed gas composition B has the highest H2O/CO ratio, which is an important factor, because the presence of excess steam drives the WGS reaction in the forward direction. Feed gas composition D obtained from steam gasification also has a high H2O/CO ratio but results show the lowest CO conversion (Figure 17). The reason for that might be the high H2 and low CO feed gas concentration constraining the WGSR. The opposite of this finding is shown by the high XCO obtained from feed gas composition A (Figure 14), containing 19 vol %dry CO and only 11.8 vol %dry H2. Feed gas composition C has the highest CH4 concentration (see Figure 16). The CH4 concentration is also restricting the achievable H2 concentration, because it cannot be converted in the carbonator. Generally, it can be seen from Figures 14−17 that the CO conversions (XCO) achieved during experiments follow the trend predicted by equilibrium calculations. 5631


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Industrial & Engineering Chemistry Research reactor. Nevertheless, one must remember that, during the experimental investigation performed within this paper, a mixed syngas that contains no tars or other impurities such as sulfur or halides is used. While the CaO in the carbonator is capable of reducing the concentration of tars,57,58 as well as sulfur and halides, in the syngas,21 these impurities might have a negative effect on the CO2 adsorption and, therefore, on the performance of the carbonator.

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CONCLUSION Calcium looping cycle for the production of hydrogen (Ca-LHP) from biomass gasification-derived syngas has been realized in a 20 kWth continuous dual fluidized-bed facility. Experiments were performed under realistic operation conditions using syngas-like gas mixtures. Through continuous operation of the facility, it could be shown that high CO conversion (>70%) and H2 concentration (>80 vol %dry) can be achieved over a long period of time. The temperature in the carbonator was varied between 606 °C and 697 °C. The highest H2 concentration and XCO was reached at a carbonator temperature of 637 °C. LR was varied between 2 and 20 molCa/molCO+CO2 for different τ values (0.4, 0.7, and 1.0 h). The effect of the LR was found to be limited, since ECO2 is close to equilibrium, even for small LR values. The value of τ was determined to have a significant influence on the XCO value in the carbonator. A H2 concentration of 84.4 vol %dry and XCO of 76.4% were achieved for LR = 6 molCa/molCO+CO2 and τ = 1.0 h. The attrition rate during experimental investigation was determined to be 1.7 wt %/h, with respect to the total solid inventory in the DFB facility. Chemical sorbent activity was evaluated by the average maximum carbonation conversion and was found to be ∼0.2 molCaCO3/molCaO for all experimental points obtained. The potential of the Ca-LHP from syngas obtained from different fluidized-bed gasification processes was evaluated. For all syngas compositions used, H2 concentrations of >70 vol %dry and XCO values of >70% could be achieved. By comparing the results obtained during experimental investigation with those found in the literature for SER gasification and catalytic WGS units, the high potential of the Ca-looping cycle became clear, because comparable or even higher H2 concentrations can be achieved.

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F−y = molar flow of component y exiting the carbonator (mol/h) GR = gasifying ratio, steam and oxygen-to-biomass ratio (kg/kgdaf) LR = looping ratio (molCa/molCO+CO2) nCa = amount of Ca in carbonator (mol) Nth = theoretical carbonation calcination cycle number S/B = steam to biomass ratio (kg/kgdaf) Xcalc = sorbent carbonate content entering the carbonator (molCaCO3/molCa) Xcarb = sorbent carbonate content exiting the carbonator (molCaCO3/molCa) XCO = CO conversion in the carbonator (%) XCO,eq = equilibrium CO conversion in the carbonator (%) X max,ave = average maximum carbonation conversion (molCaCO3/molCa) y+ = inlet gas concentration of component y (vol %) y− = outlet gas concentration of component y (vol %) τ = carbonator space time (h)

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AUTHOR INFORMATION

Corresponding Author

*Tel.: +49 (0)711 685 63393. Fax: +49 (0) 711 685 63491. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors gratefully acknowledge the financial support from the Helmholtz Association under the Project No. HA-E-0001 and thank Dipl.-Ing. (FH) Vladimir Stack for his active participation during experimental campaigns and M.Sc. Craig Hawthorne for his helpful suggestions to improve this manuscript.

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NOMENCLATURE ECO2 = CO2 capture efficiency (%) ECO2,eq = equilibrium CO2 capture efficiency (%) F+y = molar flow of component y entering the carbonator (mol/h) 5632


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