Performance tests of ilmenite mineral as oxygen carrier in a laboratory fixed bed reactor - first step in developing a new

technical approach of implementing chemical-looping combustion

Georg L. Schwebel, Daniel Hein, Wolfgang Krumm

Department of Energy- and Environmental Process Engineering at the University of

Siegen

Abstract

At present, a worldwide increase in energy demand, which is predominantly accom-

plished by the use of fossil fuels on the one hand and the alarming climate change,

caused by anthropogenic greenhouse gas emissions on the other hand, can be

observed. The most important greenhouse gas in this context is carbon dioxide

(CO2). One method to combine the necessary use of fossil fuels and the also neces-

sary reduction of CO2-emissions is the adoption of carbon capture and storage tech-

nologies (CCS). One of the most promising methods in terms of avoiding losses in

power generation efficiency is chemical-looping combustion (CLC). In contrast to

conventional combustion techniques the required oxygen is transferred via a solid

oxygen carrier, i.e. a metal oxide, from air to fuel.

The aim of the work carried out at the University of Siegen is to develop a new tech-

nical approach of implementing the chemical-looping combustion for the use of fossil

and biogenic solid fuels.

Therefore, the successfully proven conception of Herhof-IPV-Verfahren® has to be

adopted. It is characterised by the parallel arrangement of a fixed bed acting as fuel

reactor and a fluidised bed operating as air reactor. The fluidised bed reactor is also

used for the solids circulation between both reactors. In the face of environmental

and economical concerns, the use of mineral oxygen carriers is preferred.

Currently, the performance of the mineral ilmenite, well known from former research

activities, is investigated in a laboratory fixed bed reactor. The results and the com-

parison with the existing literature regarding the new reactor concept are essential

parts of the present paper.

Keywords: chemical-looping, ilmenite, fixed bed

1 Introduction

The increased concentration of carbon dioxide (CO2) in the atmosphere has been

identified as one of the main drivers of the ongoing climate change (Solomon et al.

2007). About 75 percent of the total amount of anthropogenic CO2-emissions,

roundabout 30 Gt per year, are due to the use of fossil fuels (IEA 2007; Beising

2007). At the same time, we are facing a strong rise in world wide energy demand

which will extensively be covered by the use of fossil fuels. One method to balance

these requirements is carbon sequestration for reuse or storage. Chemical-looping

combustion is a new and innovative technology to implement carbon sequestration.

During their theoretical investigations, the prime intention of Richter and Knoche

(1983) was to reduce exergy losses in combustion to increase fuel exploitation by

splitting the combustion reaction. In contrast to conventional combustion techniques,

the needed oxygen is transferred from air to fuel via a solid oxygen carrier. Using this

method, contact between fuel and atmospheric nitrogen is avoided. This means that

there is no dilution of flue gases and almost only steam (H2O) and carbon dioxide are

released. Therefore, the oxygen carrier is circulated between an air or oxidation

reactor and a fuel or reduction reactor. The terms oxidation and reduction are asso-

ciated with the oxygen carriers conversion. Figure 1.1 shows the principle of a

chemical-looping combustion system.

Figure 1.1 Principle of chemical-looping combustion.

The denotation in Figure 1.1 is that the occurring global reactions in the system can

be postulated under usage of a common metal oxide (MeOy). The oxygen carrier is

reduced by fuel in the reduction reactor following eq. (1.1).

2

n4m + MeOy + CnHm ↔

2

n4m + MeOy-1 +

2

m H2O + n CO2 (1.1)

The evolving gas flow mainly consists of H2O and CO2. Nearly pure carbon dioxide

can be obtained by condensing out the water content. The reduced metal oxide

(MeOy-1) is conveyed to the oxidation reactor for regeneration of the oxygen carrier

(1.2).

2

n4m + MeOy-1 +

4

n4m + (O2 +

21

79 N2) ↔

2

n4m + MeOy +

84

)n4m(79 +N2 (1.2)

Besides the oxidised metal oxide, a nitrogen rich flow has to be discharged from the

air reactor which can be further utilised for power generation. Depending on the used

oxygen carrier, reduction can either be exothermic or endothermic, while oxidation is

always exothermic (Garcia-Labiano et al. 2005). Most research work has been con-

ducted to the use of gaseous fuels. i. e. natural gas or syngas, in chemical-looping

combustion. In the last few years an increased interest in conversion of solid fuels

can be observed (Cho et al. 2004; Leion et al. 2007). Due to availability and low

costs, utilisation of solid fuels is comparatively more attractive than the use of liquid

and gaseous fuels.

2 Technical implementation of chemical-looping combustion

As described before, chemical-looping is a new technology and therefore no com-

mercial plant yet operates based on this energy conversion principle. Until now, all

experimental investigation has been made in laboratory scale arrangements. Despite

this fact, the necessary requirements to be met by the process equipment have

already been postulated, e.g. in Brandvoll (2005).

• A good mass transfer between gas- and condensed phase to guarantee a

complete fuel conversion.

• Due to the mostly endothermic reactions in the reduction reactor an appropri-

ate heat transfer from oxygen carrier to the reaction zone has to be estab-

lished.

• Large streams of solids have to be circulated through the chemical-looping

system avoiding any gas leakage between the reactors.

The most popular arrangement for purposes concerning the implementation of

chemical-looping technologies is the interconnected fluidised bed. In this conception,

both reduction and oxidation reactor consist of a fluidised bed. Currently, laboratory

facilities are run at various research institutions. Among the first are the ones oper-

ated at Chalmers University, Gotenburg, Sweden. A 300 Wth arrangement was used

for CLC of different gaseous fuels, described for example in Johansson et al. (2006).

Based on a different reactor geometry and using a carbon stripper Berguerand and

Lyngfelt (2008 a and b) conducted some CLC experiments with different coals and

petroleum coke in a 10 kWth plant. Another example is a 10 kWth arrangement which

is run in Spain, in which de Diego et al. (2007) investigated the use of copper based

oxygen carriers. In Vienna, Austria, tests with ilmenite and nickel based oxygen

carriers were performed in a 120 kWth interconnected fluidised bed, described in

Kolbisch et al. (2008) as well as in Bolhar-Nordenkampf et al. (2008). A new 1 MWth

facility is being erected at the University of Darmstadt to test the technical feasibility

of coal as fuel in chemical-looping as well as the application of carbonate looping

Epple and Ströhle (2008).

In contrast to concepts based on interconnected fluidised bed reactors, a new ap-

proach of a two-stage configuration for the chemical-looping process based on a

parallel-arranged fluidised bed and a fixed bed reactor has been suggested (Figure

2.1). This new reactor configuration in the Herhof-IPV-Verfahren® is nowadays a

successfully proven gasification process for waste-derived fuels and solid biomass.

One of the main advantages of this gasification process is to avoid the product gas

diluting with nitrogen from air, because the required heat is provided by the circulat-

ing hot bed material out of the fluidised bed combustor instead of partial combustion

of the fuel. Design, construction, operation and results of the Herhof-IPV-Verfahren®

are presented in Hamel et al. (2007), Weil et al. (2004) and Hamel et al. (2004). A

detailed experimental analysis of the deactivation and regeneration of catalytic min-

eral bed material whilst circulating between the oxidating and reducing athmospheres

is given in Hein and Krumm (2008).

Figure 2.1 Scheme of the postulated arrangement based on the Herhof-IPV-

Verfahren®.

To adapt this system to a chemical looping process the fixed bed acts as fuel reactor

and the fluidised bed operating as air reactor, while ensuring solids circulation.

Especially in connection with the utilisation of solid fuels, integrating a fixed bed

reactor for the reduction step seems to be very promising. The most significant ad-

vantages are avoiding fuel segregation connected with less char at the reactor exit,

adjustable residence time and less power demand for fluidisation.

Following Leion et al. (2007), remaining char conveyed from fuel to air reactor is a

problem in connection with utilisation of solid fuels in a CLC - system. One reason for

remaining char from solid fuels conversion is the varying residence time of fuel and

oxygen carrier in the reactor. The fuel residence time in the reactor is shorter than

neccassary for the full conversion. In a fixed bed, the residence time of fuel and

oxygen carrier is relatively constant and depends on reactor geometry and mass

flows. To solve this problem in a fluidised bed, the integration of a carbon stripper

downstream of the reactor which is also designed as a fluidised bed is recom-

mended, aiming at the separation of the low density char from high density oxygen

carrier by segregation effects due to fluidisation. However, it has to be noted, that

these segratation effects probably lead to varying residence times in the fuel reactor.

In a fixed bed reactor the segregation due to fluidisation can be neglected. Adjusting

the residence time a necessary carbon stripper can be scaled down to a minimum.

Furthermore, the fixed bed reactor is characterised by reaction zones for combustion

and reburning, see Figure 2.1. This means that the devolatilisation gases will pass a

layer of fresh oxygen carrier towards the reactor exit and therefore a complete fuel

conversion is expected. The oxidation reactor is designed as an expanded fluidised

bed reactor. Besides realising the oxidation of the reduced carrier, the fluidised bed

works as a driver for the solids circulation in the process. When using solid fuels, the

oxygen carrier stream is also loaded with ash, which has to be separated in order to

keep up an efficient oxygen content. Assuming a lower density compared to the

oxygen carrier, ash and fine-grained oxygen carrier can be separated at the top of

the fluidised bed. The energy consumption of the mentioned reactor arrangement is

cut down to almost half due to the fact that only one fluidised bed reactor is required.

Considering the optional utilisation of biomass, e.g. in combination of carbon capture

and reuse concepts, the required plant capacities are around 20 MWth.

3 Ilmenite as oxygen carrier in chemical-looping combustion

Efficient chemical-looping systems depend on the the quality of the applied oxygen

carrier, (e.g. Ishida und Jin 1996; Adanez et al. 2005; Abad et al. 2007 a und b; de

Diego et al. 2007; Erri und Varma 2007; Hossain et al. 2007 and Johannson 2007).

Jerndal et al. (2006) used thermal analysis to identify applicable oxide/metal sys-

tems. Synthetic and natural oxygen carriers based on Cu, Mn, Fe and Ni were found

to be promising due to their ability of fuel conversion and available oxygen.

Synthetic oxygen carriers are well-known, but not yet available in a technical scale

due to its expensive production method. The common used methods in literature are

freeze granulation, as described in Cho et al. (2004) or wet impregnation, see Cor-

bella et al. (2005).

To be suitable for CLC, reactivity of both types of oxygen carriers has to be sufficient

for reduction and oxidation purposes, (Cho et al. 2004). The abundance of oxygen

for the reduction step is crucial. Among others, Mattisson et al. (2003) define the

ratio of free oxygen or oxygen ratio in (3.1), where available oxygen is related to the

carriers oxidised mass.

ox

redoxO

m

mmR

−= (3.1)

Due to the thermal and mechanical forces long term stability of the oxygen carrier is

also important. In addition, it never should lead to agglomeration and carbon forming

side reactions.

In this paper experimental results about natural ilmenite as oxygen carrier are pre-

sented. In general, natural oxygen carriers have some advantages like low costs,

e.g. 120 €/t for ilmenite and its concentrates (Statistisches Bundesamt 2009). Com-

pared to Ni based materials most natural carriers like ilmenite do not have health

concerns.

Ilmenite has been the subject of various investigations, e.g. in Leion et al. (2008) as

well as in Berguerand and Lyngfelt (2008 a and b). It was found that agglomeration

occured with the use of ilmenite particles in the range of 90 and 250 µm at low su-

perficial velocities and temperatures above 600 °C. To avoid agglomeration a mini-

mum fluidisation velocity of 0.1 m/s is recommended.

According to Berguerand and Lyngfelt (2008 a) and Leion et al. (2008) oxida-

tion/reduction of ilmenite mineral involves three different levels of oxidation, see

Table 3.1. The raw material, FeTiO3, corresponds to the total reduced level of a

mixed phase of FeO + TiO2. Fe2TiO5·TiO2 is the most oxidised level corresponding to

the decomposed particular compounds of Fe2O3 + 2 TiO2 after oxidation at 800 °C

(Zhang and Ostrovski 2002). Leion et al. (2008) mentioned an intermediate level of

oxidation which corresponds to Fe3O4 + TiO2.

Table 3.1 Oxidation levels during reduction/oxidation of ilmenite, Berguerand and

Lyngfelt (2008 a) and Leion et al. (2008).

Level of oxidation Compound Corresponding compounds

Reduced (raw) FeTiO3 FeO + TiO2

Intermediate Fe3Ti3O10 Fe3O4 + TiO2

Most oxidised Fe2TiO5·TiO2 Fe2O3 + 2 TiO2

According to these oxidation levels, the available oxygen O2,avail during reduction is

given by the following equations.

avail,22232 O2

1TiO2FeO2TiO2OFe ++↔+ (3.2)

avail,2243232 O2

1TiO6OFe2TiO6OFe3 ++↔+ (3.3)

A theoretical oxygen ratio of RO = 0.05 (3.2) respectively RO = 0.017 (3.3) can be

obtained assuming complete oxidation and reduction steps. Due to impurities in

natural ilmenite, e.g. see Table 3.2, the actual oxygen ratio is around RO = 0.04 (3.2)

and RO = 0.0136 (3.3) when impurities content is around twenty weight percentage

and impurities are totally inert. However, it is obvious that impurities like sulphur are

involved, indicated by experimental SO2 - plots.

The used ilmenite was gratefully provided by ArcelorMittal Eisenhüttenstadt GmbH.

Three different particle size fractions of ilmenite are used: the fine fraction with parti-

cles in the range of 0.18 to 0.25 mm, the intermediate fraction in the range of 0.5 to

0.71 mm and the coarse fraction in the range of 1.0 to 1.4 mm. The mean bulk den-

sity is 4.64 kg/dm³ and the mean macroscopic porosity is 0.47.

Table 3.2 Composition of natural ilmenite (ArcelorMittal).

Component mass-%

dry basis Component

mass-%

dry basis

S 0.310 P 0.005

Fe 35.800 Na2O 0.740

(Fe)O 24.990 K2O 0.150

SiO2 8.720 Mg 0.095

Al2O3 5.000 Ti 31.400

CaO 1.300 V 0.320

MgO 2.700

4 Experimental section

4.1 Equipment

The experimental plant, Figure 4.1, was used for the analysis of heterogeneous gas-

solid systems at high temperatures and different flow- and chemical conditions,

Hamel (2001). For CLC - analysis the experimental arrangement had to be adopted,

above all the gas distributor had to be changed in order to handle oxygen carrier

particles. The core component is an electrically heated tube reactor with an inner

diameter of 0.0531 m applicable for a maximum bed height of about 0.15 m.

Figure 4.1 Experimental plant.

The fluidisation and fuel gases are supplied via mass flow controlers (mfc). A fixed

bed of incombustible material, such as ceramical particles, leads to a homogenous

and preheated mass flow into the tube reactor. Offgas is analised to determine the

main gas components like H2, CH4, C2H4, CO, CO2 and optional organic carbon. A

constant N2-flow as reference mass flow allows the balance calculation of the other

components. Therefore, the N2 is assumed to be totally inert. The measured gas

concentration profiles are numerically deconvoluted by a tanks-in-series model to

determine kinetic parameters.

4.2 Procedure

A total mass of 0.3 kg ilmenite with one of the mentioned particle size fractions is

arranged in a fixed bed. The superficial gas velocity is just below minimum fluidisa-

tion velocity. Figure 4.2 shows one experimental run. After a short period of inertisa-

tion, reduction takes place by injecting methane with a mole fraction of 0.08 for ten

minutes. Finally, the oxygen carrier is oxidised by a mixture of oxygen and nitrogen

until the original molar fraction of oxygen is obtained.

Figure 4.2 Molar fraction of the dry outlet gases as a function of time.

4.3 Data evaluation

The reaction rate r (4.1) in relation to the amount of oxygen available in the oxidised

carrier is defined as the time derivative of the oxygen conversion X during reduction

(4.2), Cho et al. (2004).

dt

dXr = (4.1)

redox

red

mm

mmX

−

−= (4.2)

The time derivative of X can be obtained from the oxygen released during reduction

(4.3) regarding the oxygen containing components CO, CO2 and SO2 detected in the

dry gas sample.

)xxx2x2(n)mm(

M

dt

dXOHCOSOCOout

redox

O222 ++⋅+⋅⋅⋅

−= & (4.3)

MO means the oxygen mole weight, outn& the total mole flow at the reactor outlet and

xi is the molefraction of component i. OH2x was calculated based on the methane

conversion.

According to the definition of oxygen ratio (3.1) the mass of available oxygen is given

by eq. (4.4) with RO=0.04. The mass of the reduced oxygen carrier mred can easily be

obtained by determining the weight of the raw sample.

−+−⋅=−=

O

redredox0,OR1

11mmmm (4.4)

The CO2-yield 2COγ (4.5) is a characteristic quantity for fuel (methane) conversion.

42

2

2

CHCOCO

COCO

xxx

x

++=γ (4.5)

Both, reduction rate and CO2-yield, have to be plotted as a function of temperature,

particle size and fluidisation velocity during reduction period.

5 Results and discussion

Two characteristic values demonstrate the way of oxygen carrier working, i.e. the

maximum value of overall reduction rate rmax and the mean value rstat of a stationary

interval at the end of the reduction period, Figure 5.1.

Figure 5.1 Reduction rate r as a function of oxygen conversion X.

Results for maximum and stationary reduction rates as a function of temperature for

the three different particle sizes with a normalised superficial gas velocity, i.e.

u/umf = 0.03 are shown in Figure 5.2. For all tested particle sizes and temperatures

the reduction rates vary between 0.0002 and 0.0005 1/s. This corresponds to results

obtained from Mattisson et al. (2001), where the reduction rates lie between 1 and

8 %/min using iron oxide. Increasing the temperature from 800 to 900 °C the reduc-

tion rates are almost doubled. The difference between rmax an rstat decreases with a

rise in temperature. The variation in the measuring points for one specific tempera-

ture is propably influenced by the number of reduction/oxidation cycles, as stated by

Leion et al. (2008).

Yet, no reliable correlation between reactivity and particle size is known as hardly

any work has been carried out on different particle sizes, (Johansson 2007). Right

now it appears as though different particle sizes do not have a significant impact on

reduction rate. Temperature influence on reduction rate ist obvious. A maximum

CO2-yield up to 95 % was found at a temperature of 900 oC, Figure 5.3. The influ-

ence of the particle size seems to be neglectable in this context.

Figure 5.2 Maximum rmax and stationary rstat reduction rates as a function of

temperature for different particle sizes. Normalised superficial gas

velocity: u/umf = 0.03.

Figure 5.3 Maximum and stationary CO2-yield 2COγ as a function of tempera-

ture for different particle sizes. Normalised superficial gas velocity:

u/umf = 0.03.

From experiments accomplished at 900 °C with intermediate sized particles, the

impact of varying superficial velocities is determined, Figure 5.4.

Figure 5.4 Maximum rmax and stationary rstat reduction rates as well as maxi-

mum and stationary CO2-yield 2COγ as a function of superficial gas

velocity/minimum fluidisation velocity u/umf.

Increasing velocity leads to an increasing reduction rate and unfortunately to a sub-

stantial decrease of fuel conversion. This decrease of CO2-yield is caused by a

greater amount of fuel which has to be converted with the same amount of oxygen

carrier but also by a less efficient contact between the particles and gas at higher

velocities, (Johansson et al. 2006). Therefore, the bed height has to be increased in

a technical arrangement while increasing the gas velocity.

Finally, it was found that at all experimental runs no agglomeration occurred during

reduction although having a bed velocity below 0.1 m/s. Agglomeration occurred only

during oxidation after a long-term reduction cycle, indicated by a sudden rise in

pressure drop over the bed.

6 Acknowledgements

The present work has been carried out at the Department for Energy- and Environ-

mental Process Engineering at the University of Siegen (Germany). The work was

aided by the doctoral scholarship programme of the DBU (www.dbu.de). The used

samples of ilmenite were gratefully provided by ArcelorMittal Eisenhüttenstadt

GmbH.

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