Combustion Performance of Sewage Sludge in a Novel CLC System

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Combustion Performance of Sewage Sludge in a Novel CLC System with a Two-stage Fuel Reactor Jingchun Yan, Laihong Shen, Shouxi Jiang, Jian Wu, Tianxu Shen, and Tao Song Energy Fuels, Just Accepted Manuscript • DOI: 10.1021/acs.energyfuels.7b02493 • Publication Date (Web): 02 Oct 2017 Downloaded from http://pubs.acs.org on October 5, 2017

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Combustion Performance of Sewage Sludge in a Novel CLC System with a Two-stage Fuel Reactor Jingchun Yan1, Laihong Shen*1, Shouxi Jiang1, Jian Wu1, Tianxu Shen1, Tao Song2 1 Key Laboratory of Energy Thermal Conversion and Control of Ministry of Education, School of Energy and Environment, Southeast University, Nanjing 210096, China 2 School of Energy and Mechanical Engineering, Nanjing Normal University, Nanjing 210042, China

Email: [email protected] Abstract Chemical looping combustion (CLC) is a promising and efficient technology for sewage sludge (SS) combustion with carbon capture. CLC reactor configuration is crucial to the intimate contact between the solid phase and the gas phase species. This work proposed a novel CLC unit with a two-stage fuel reactor. Different from the conventional CLC unit with a single-stage bubbling bed as fuel reactor, two-stage fuel reactor design can make gas phase and solid phase be in adequate contact and achieve gas flow redistribution in the fuel reactor. On this unit, both cold and hot experiments were conducted. The gas-solid flow characteristics were studied on the cold model and the system was successfully commissioned and in stable operation. No gas leakage occurred between reactors. In the hot experiments, SS from the municipal wastewater treatment plant was chosen as solid fuel and hematite with a size range of 0.3-0.45 mm was used as oxygen carrier. To thoroughly evaluate the performance of this noval CLC system, the combustion compensation efficiency and carbon supplementation efficiency was proposed. The results showed that the two-stage fuel reactor design was beneficial to improving the carbon conversion efficiency of the whole system. The carbon conversion efficiency, carbon capture efficiency and combustion compensation efficiency increased within the temperature range of 800-900 ℃ , while the carbon supplementation efficiency decreased at high temperature. Besides, the effect of SS feed rate was also investigated. High SS feed rate resulted in lower carbon conversion efficiency due to the large bubbles produced during SS gasification process. At last, the phase characteristics of fresh oxygen carrier and oxygen carrier extracted from the both fuel reactor were detected. Fe3O4 was the main reduced phase of hematite in the FR. Key words: chemical looping combustion; two-stage fuel reactor; cold model; sewage sludge.

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1. Introduction Sewage sludge (SS) is formed during disposal of industrial or municipal wastewater, which refers to the combination of the liquid or water-carried residue. It is rich in nutrients, toxic compounds and pathogenic microorganisms

[1]

. The improper

disposal of SS will not only lead to health problems, but also result in unbalance of eco-system. As illustrated by Werther et al.

[2]

, the disposal route of SS contains

fertilizer using, land filling, dumping in the sea and incineration. Thermal treatment is a competitive method to dispose SS in the future with several advantages including diminishing the vast quantities of sludge, destructing toxic organic matters and minimizing odor generation. However, incineration of SS is inevitably faced with the same environmental and climate issues, such as emitting large volume of carbon dioxide and other greenhouse gases, as using fossil fuels. Even worse, thermal treatment of SS may pose heavy metal pollution related with its high heavy metal content, such as arsenic, chromium, lead, mercury and selenium etc.

[3]

. Once these

trace elements seep into underground reservoirs or infiltrate into soil, severe water pollution or soil contamination will happen and thus human health will be gravely threatened by their indirect effects. SS, by its definition, is considered as a waste but it is also regarded as a type of biomass resources according to its source and characteristics. Thus SS is likewise characterized by high organic contents and biological elements such as chlorine, nitrogen and sulfur. Incineration of SS at below 700℃ is good for the generation of dioxin on account of substitution of hydrogen on benzene by chlorine, of which 2,3,7,8-tetrachlorodibenzo-para-dioxin is considered as the most toxic matter

[4]

. Additionally, the nitrogen content of the SS is considerably

higher than that of other fuels. Above 1300℃, the emission of NOx and N2O are expected to be high, which is in agreement with the previous investigations [5]. As stated above, another main problem during incineration of SS is the greenhouse gases emission, of which carbon dioxide (CO2) is the accumulated maximum. CO2, which is believed to be mainly produced via human activities, accounts for 58.8 percent of all kinds of greenhouse gases, consequently, arousing climate change and presenting a great impact on the ecological environment [6]. In the ecosystem, carbon in the atmosphere is chemically stored in organism. As a kind of biomass, the production of CO2 through incinerating of SS is not regarded as additional carbon emission into the atmosphere. Thus if the CO2 generated during the SS combustion process is captured and storaged, negative CO2 emission will be achieved. Carbon Capture and Storage (CCS) is the only accessing way, for the time

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being, to mitigate greenhouse gases emission. However, traditional CCS technologies such as oxy-fuel, post-combustion and pre-combustion, remain some drawbacks like high CO2 separation cost and extra energy or equipment demand. For all the above reasons, chemical looping combustion (CLC) is a superior option for SS disposal. CLC is a promising and efficient alternative for combustion with carbon capture, which overcomes shortcomings of conventional CCS technologies due to the intrinsic separation of CO2 from the flue gases during combustion [8]. Furthermore, the dioxin and thermal NOx are significantly reduced due to the appropriate reaction temperature ranging from 800~950℃ during the CLC process. This is achieved by separating the combustion air and fuel into two reactors, the so-called air reactor (AR) and fuel reactor (FR) respectively, via an oxygen carrier. As depicted in Figure 1, the reduced state of the oxygen carrier MexOy-1 are oxidized by O2 in the AR according to the reaction (R1), after which the produced oxidized state of oxygen carrier MexOy are transferred into the FR to react with fuel according to the reaction (R2). The reproduced MexOy-1 then circulate back into the AR via a loop seal (LS). 2Me O + O → 2Me O + Q

(R1)

C H + 2n + mMe O (R2) → 2n + mMe O + nCO + mH O − Q

Figure 1. Prototype of chemical looping combustion process 3


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The combustion of SS using CLC technology based on hematite oxygen carrier was carried out in a 1 KWth continuous unit in our previous work [7]. It was found that the carbon conversion efficiency peaked out at less than 90% with the sludge flow rate of 100 g/h at 900℃. Moreover, the results showed that SS had a higher amount of unconverted carbon compared to Shenhua bituminous coal in a CLC process. Despite the effect of higher sludge feed rate on combustion performance was not conducted in this work, it could be predicted that this CLC unit would have a limited fuel handling capacity. Lower carbon conversion efficiency and combustion efficiency would be obtained at higher sludge feed rate. Therefore, a more efficient CLC unit with a higher handling capacity is necessary to be designed. The CLC reactor is the place where redox reactions occur. The proper design and operation of the reactor are essential for efficient conversion of fuel. In 2001, Lyngfelt et al.

[9]

at Chalmers University of Technology (CHALMERS) firstly proposed a

circulating system composed of two connected fluidized beds. After that, the reactors of variety ways were designed worldwide based on the circulating fluidized bed (CFB) principle, which was stated in a review by Juan Adanez et al.

[10]

. The published

results revealed that the most common design of single-stage FR is faced with some knotty problems such as insufficient gasification of solid fuel and low conversion efficiency of combustible gases. The solid fuel and oxygen carrier particles are mixed in a unique reactor where insufficient contact between produced syngas and oxygen carrier particles occurs. Simultaneously, the single-stage FR may suffer from elutriation of unburnt carbon fines, emission of combustible gasification products and recirculation of unburnt char from the AR [11-12]. Moreover, as for high volatile content of SS, the devolatilization and char gasification rate of SS are so fast that plenty of bubbles containing combustible gases are generated and grow up with the rise of bed height. This leads to the low conversion efficiency of fuel in the conventional CLC unit. In 2013 Joachim Werther and co-workers proposed a 25kWth CLC reactor consisting of a two–stage bubbling fluidized bed as the FR coupled with a circulating fluidized bed. Both cold and hot experiments were performed. In the hot experiment using ilmenite as oxygen carrier, the solid fuel was introduced into the lower stage reactor and CO2-concentrations in the dry FR off-gas of above 90 vol.% were achieved. This unit was designed to improve the carbon conversion efficiency via contact between the combustible gases and solids from the lower stage reactor and freshly oxidized state of oxygen carrier with a higher reactivity in the upper stage

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reactor. Although the design of two-stage FR has been proposed, the research of this novel unit is at its early stage, and there are still many problems demanding prompt solution. Accordingly, a new type of CLC system with a two-stage FR has been proposed to improve the performance of the overall CLC process in this work. On the basis of the cold model experimental results, this work reports on the primary experiment results operating with hematite as oxygen carrier and SS as solid fuel. The aim of this study is to obtain the characteristics and performance of the two-stage FR and dispose SS efficiently. 2. Experimental section 2.1 Bed materials and fuel preparation The bed materials in the CLC facility cold model are silica sands ranging from 0.1~0.45 mm and the inventory is around 10 kg. Natural hematite, provided by Nanjing steel manufacturing company, was sieved to the size range of 0.3~0.45 mm and then calcined in a muffle oven at 950℃ for 3h to improve the mechanical strength of the particles. The calcined hematite particles were sieved again to reduce the fine powder and to get uniform grain size of oxygen carrier. The inventory of oxygen carrier was about 12kg. The chemical compositions of the fresh oxygen carrier based on XRF (X-ray Fluorescence, ARL-9800, Switzerland) were showed in Table 1. Table 1 Elemental compositions of the fresh hematite oxygen carrier (wt.%) Fe2O3

SiO2

Al2O3

CaO

P2O5

TiO2

K2O

SO3

Others

83.25

7.06

5.33

0.23

0.29

0.09

0.03

0.25

3.47

De-watered SS supplied by Jurong wastewater treatment plant was used as solid fuel in the experiment. The samples were dried in a dry oven at 105℃ for 24h and then sieved to yield a size range of 0.3~0.45mm. The proximate and ultimate analyses of the SS, together with the lower heating value were illustrated in Table 2. Table 2 Proximate analysis and ultimate analysis of SS and content of carbon Proximate analysis(wt.%, ad)

SS

ultimate analysis(wt.%, ad)

M

V

FC

A

C

H

O

N

S

LHV (MJ/kg)

4.72

36.64

3.45

55.19

22.54

3.761

8.935

4.11

0.744

11.06

5


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2.2 Experimental facility 2.2.1 CLC facility cold model. Experimental tests of gas-solid flow characteristics were conducted on the CLC cold-flow model, as showed in Figure 2. The facility is composed of an AR, a two-stage FR (FR-I and FR-II), two cyclone separators (ARC and FRC), two LSs (ALS and FLS) and a syphon (SP). The AR is designed as a fast fluidized bed, which is a cylinder of 30 mm in inner diameter and 2700 mm in height. And the FR is excogitated as a two-stage rectangular bed on the basis of step oxidation of fuel, with a cross section of 100×50 mm2. The FR-II was 250 mm in height, while the FR-I was a spout-fluidized bed with a height of 1000 mm. The AR and two-stage FR are connected by ALS, FLS and SP, which are also circular columns with an inner diameter of 40 mm. All the pipes of the system are made of polymethyl methacrylate so that it is convenient to observe the gas-solid flow behavior in the system. Eight pressure taps are designed to measure pressure drop which is converted to digital signal via pressure transmitter collected by computer. Different fluidizing agents are injected into AR, FR, LSs or SP, and compositions of the exhaust gases are detected by a NGA 2000 type gas analyzer (Emerson Company, Bloomington, MN) to study gas leakage between reactors. Operating conditions and primary parameters in the cold test are as follows: The flow rate ranges of AR, FR, LSs and SP, which were obtained by theoretical calculation, are 2.0~5.0 m/s, 0.3~0.8 m/s, 0.2~0.5 m/s and 0.2~0.5 m/s, respectively. In order to ascertain how the gas from SP was distributed to AR and FR, SP was fluidized by CO2 while other reactors and LSs were fluidized by N2. Furthermore, for the purpose of inspecting the gas leakage between the AR and the FR, the AR and the FR were fluidized by O2 and CO2, respectively, while two LSs and the SP were fluidized by N2. The proportion of gas distribution of SP and the gas leakage between the AR and the FR were determined through detecting gas compositions of flue gas expelled from two cyclones. 2.2.2 Continuous interconnected fluidized bed CLC reactor The hot experiment facility was established with the same scale of the CLC facility cold model and formed a circle with end-to-end configuration, which was also depicted in Figure 2. Oxygen carrier particles in the fast fluidized bed were conveyed to the top of the bed by a current of air, separated by the attached cyclone, and then flowed into the FR-II through the ALS. The freshly regenerated oxygen carrier

6


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particles in the FR-II fell into the downcomer via FR cyclone after reacting with combustible gases released in the FR-I, and then flowed back into the FR-I through the FLS. In the end, the exhausted oxygen carrier particles from the FR-I were aerated inside the SP and passed into the AR, which constituted a circulation of oxygen carrier particles in the interconnected fluidized bed.

Figure 2. Prototype of the CLC facility cold model and the continuous interconnected fluidized bed CLC reactor

Two reactors were electrically heated to the target temperature in an oven controlled by a temperature programmer to compensate the heat loss during operation. And the reaction temperature was measured by seven K-type thermocouples, which 7


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were installed in the same places as the pressure taps 1-7. By virtue of a variable-speed screw feeder, SS particles were pneumatically conveyed to the bottom of the spout-fluidized bed with a N2 stream through a volumetric flowmeter. The operating conditions and original parameters were determined by the results obtained from the cold model experiment according to the relationship of thermal expansion between gas volume and temperature. The air flow of the AR was set as 1.5 m3/h. A steam flow, which was generated by heating water from a constant flow pump to boiling point and then introduced to the FR, was 0.24 kg/h. Two LSs and the SP were all fluidized by nitrogen and the flow rates of them were 0.36 m3/h (ALS), 0.24 m3/h (FLS) and 0.48 m3/h (SP), respectively. The fuel feed rate was variable and nitrogen for transferring fuel was 0.9 m3/h. The ambient temperature was approximately 25℃ and the temperature range of FR was 800~900℃. Once the reactor temperature was kept stable at the target value, the gaseous products from the outlet of the AR and the FR and the middle of the FR were firstly introduced into ash separators and desiccators and then collected by gas collecting bags. The compositions of gaseous products were also detected by a NGA2000 type gas analyzer.

2.3 Data evaluation in the CLC facility cold model The circulation index (CI), which is defined by the pressure drop measured between pressure taps located at the riser’s entrance and outlet multiplied by the actual gas volume flow at the outlet of the AR and measured temperature in the AR, is adopted in the experiment [13,14]. $ + 273 = ∆ × ,!"# × 273 The gas distributions of SP are calculated as follows:

'( = ) × * )

(E1)

(E2)

'(, = ,) × * ,) '( -'( = × 100% '( + '(, '(, -'(, = × 100% '( + '(, In the above equations, ) and ,) are the volume flow of the gas

(E3) (E4) (E5) at the

outlet of ARC and FRC, while * ) and * ,) are the concentrations of CO2 at the outlet of ARC and FRC. The gas distributions of SP are interpreted by volume flow and distribution proportion, as Eqs. (E2)-(E5) showed. 2.4 Data evaluation in the continuous interconnected fluidized bed CLC reactor 8


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The exit gas flow of carbonaceous gases from the FR and the AR can be calculated by N2 mass balance method:

-),,1 = -23 ,,1 ∙

5)63 ,,1 + 5)6,,1 + 5)78 ,,1

(E6)

1 − 95)63,,1 + 5)6,,1 + 5)78 ,,1 + 573 ,,1 :

-),, = -23 ,, ∙ -), = -23, ∙

5)63 ,, + 5)6,, + 5)78 ,, 1 − 95)63 ,, + 5)6,, + 5)78 ,, + 573 ,, : 5)63, 1 − 95)63 , + 563, :

(E7)

(E8)

where -),,1 , -),, and -), represent the exit gas molar flow of carbonaceous gases from the FR-I, FR and AR, respectively. -23,,1 and -23 ,, stand for the N2 inlet flow of the FR-I and FR, while -23 , means the N2 inlet flow of the AR, including the N2 in the air and influx from the SP. 5;,,1 and 5;,, (< = * , *, => , = ) are the measured gas concentrations (volume percent) from the FR-I and FR, and 5;, ( < = * , * ) are the measured gas concentrations (volume percent) from the AR, respectively. The carbon conversion efficiency of the FR-I (?),)!@,1 ) is defined as the fraction of total carbon contained in the SS converted to carbonaceous gases in the FR-I: -),,1 ?),)!@,1 = × 100% (E9) -),AB"CDE where -),AB"CDE is the total molar flow contained in the fed SS. The carbon conversion efficiency of the whole CLC system (?),)!@ ) is defined as the ratio of carbonaceous gases at the outlet of FR and total carbon contained in the SS:

?),)!@ =

-),,

× 100% (E10) -),AB"CDE where the symbol C in the ?),)!@ and ?),)!@,1 is an abbreviation on behalf of the meaning of Carbon. The gas fraction of each component (* , *, => and = ) is presented as the ratio of the gas species in the product gases: 5; H; = 5)63 + 5)6 + 5)78 + 573

(E11)

where 5; ( < = * , *, => , = ) is the volume fraction of each component (* , *, => and = ) in the flue gas. The carbon capture efficiency (?)) ) of the reactor system is evaluated as the ratio

9


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of aggregate amount of gaseous carbon leaving the FR to the total amount of carbonaceous gases leaving the FR and the AR, i.e. the overall reactor system. It reads:

?)) = where -),,

-),, (E12) -),, + -), and -), are derived from the integrations of Eqs. (7) and (8) over the

whole reduction period. Conventional data evaluation methodology is insufficient to thoroughly estimate the performance of CLC system with a two-stage FR. Therefore, new assessment criteria are necessary to be defined and proposed. The combustion compensation efficiency (?I,, ), which is an evaluation of favorable effect of the FR-II on combustible gases conversion neglecting fine char from the FR-I, is defined as

?I,, = 1 −

J∑ LMN O

,PP ,!"#

(E13)

J∑ LMN O

,P ,!"#

where JLMN O, ,!"# and JLMN O, P

PP ,!"#

(< = *, => , = ) are the volume flux

of gasification products (*, => , = ) measured from the free board of the FR-I and the outlet of FRC, respectively. They are calculated by N2 stream mass balance and gas concentrations (volume percent) of gas acquisition in the gas collecting bags. For the sake of representing the fraction of carbon gasified in the FR-II, carbon supplementation efficiency (ΩR,ST ) is proposed as

Ω),, =

[∑ VW ),; ],PP,!"# − [∑ VW ),; ],P ,!"# VW ) − [VW ) ];Y,!"#

(E14)

with VW ),; (< = * , *, => ) the molal quantity of * , * and => , VW ) the total molal quantity of carbon contained in the fed SS, [VW ) ];Y,!"# the molal quantity of carbon entrained with oxygen carrier to AR, which can be approximately calculated by the molal quantity of carbon dioxide at the outlet of AR.

3. Results and discussion 3.1 Operation behavior of the CLC facility cold model 3.1.1 Pressure profile The pressure drop, which originates from wall and inter-particle frictions, extrusions, constrictions, bends, acceleration of particles and the relative velocity between gas and solids, is the crucial indicator reflecting operating status of the

10


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system [15]. The pressure profile presented in Figure 3, is measured under certain working conditions which is showed in Table 3. These working conditions were determined from theoretical calculation and a large number of experimental results. The system maintained stable operation under these operating parameters. Table 3 Working conditions and parameters of bed pressure drop experiments Number

uAR,in （m/s）

uFR,in （m/s）

1

2.554

0.361

uSP（m/s） uALS（m/s） uFLS（m/s） 0.332

0.309

0.287

9 8 7

Pressure drop (kPa)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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∆p76

6 5 4

∆p13

3 2

∆p54

1 0 0

500

1000

1500

Time (s)

Figure 3 Pressure drop measured in the CLC facility cold model

As Figure 3 Shows, the pressure drops across both the FR-I (∆p76) and the FR-II (∆p54) are quite stable and fluctuate in uniform. The pressure drop depends on the voidage, gas and solids density

[12]

. It can be observed that ∆p76 is higher for the

larger total solids inventory. This is beneficial to sufficient reaction between solid fuel and oxygen carrier due to the prolonged residence time of solid fuel particles and gasification products. The mean solid concentration at the upper part of the AR is lower than that at the lower part of the AR. But the solid concentration increases at the location closed to the inlet of cyclone, which can be observed through the AR made by polymethyl methacrylate and the pressure drop profile across the AR. This is because of the sharp bend design of cyclone, where the solids have collision with wall, lose energy and are accumulated there. As a result, Strong fluctuations in the pressure drop across the AR (∆p13) are visible.

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Energy & Fuels

3.1.2 Influence of superficial gas velocity of SP on pressure drop and circulation index. The SP is the key device connecting the AR and the FR. In order to figure out the pivotal role of SP on the performance of the CLC facility, the superficial gas velocity of SP was increased from 0.265 to 0.354 m/s with the other arguments intact. The influences of superficial gas velocity of SP on performance are shown in Figure 4. 7

24

3

20

-1

∆p76

Circulation index CI (kPa⋅m ⋅h )

22

6

Pressure drop ∆p (kPa)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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5

CI

18

4

16

3

14

∆p13

2

12 10

∆p54

1

8

0

6 0.26

0.28

0.30

0.32

0.34

0.36

Superficial gas velocity of SP uSP,in (m/s)

Figure 4 The influences of superficial gas velocity of SP on pressure drop and circulation index

It can be seen that larger pressure drops ∆p13 and ∆p54 are achieved with the increasing of superficial gas velocity of SP whereas the pressure drop ∆p76 decreases, which can be explained by an improvement of circulation rate of the system. The increase of superficial gas velocity of SP leads to an augment of circulation index due to the increasing oxygen carrier throughput of SP. The circulation rate of the oxygen carrier is crucial to the chemical looping process and it must be sufficiently high to supply demanded oxygen and heat to the reactions in the FR. The experiments conducted in the CLC facility cold model provide guidance for operation of the hot unit, despite the circulation index is used for qualitative analysis. If the circulation rate in the hot unit is quite low, the inventory of oxygen carrier and/or the superficial gas velocity can be increased to settle the problem. However, their regulation capacity is limited, because the increase of inventory may lead to a higher pressure drop in the reactors which results in extra energy consumption, and higher superficial gas velocity may change the fluidization status.

3.1.3 Gas distributions of the SP and gas leakage between reactors 12


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The distribution ratios of fluidization gas in the SP flowing into the AR and the FR are necessary to be ascertained for accurately calculating the N2 mass balance. Thus, experiments are conducted under the working conditions shown in Table 4 and the results are depicted in the Figure 5. Table 4 Working conditions and parameters of gas distribution experiments Number

uAR,in （N2,m/s）

uFR,in （N2,m/s）

uSP （CO2,m/s）

uALS （N2,m/s）

uFLS （N2,m/s）

1

2.554

0.361

0.265-0.354

0.309

0.287 65

0.8

PSP-FR

55

0.6 50

VSP-FR

0.5

45

PSP-AR , PSP-FR (%)

60

0.7

VSP-AR , VSP-FR (m3/h)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

PSP-AR

0.4

VSP-AR

40

0.3 0.26

0.28

0.30

0.32

0.34

35 0.36

Superficial gas velocity of SP uSP,in (m/s)

Figure 5. The influence of superficial velocity of SP on the gas distributions of SP

It is apparent that gases from the SP distributed into both the AR and FR rise with the increase of superficial velocity of the SP on account of the increasing volume of input gas. Nevertheless, the distribution ratio of fluidization gas flowing into the FR decreases due to more gases mixed up with bed material particles entrained to the AR. As stated above, it is not hard to find that the gas distributions of SP have a close relationship with particle circulation rate. Gas leakage between reactors, which is hazardous to the chemical looping process, is inevitable in the interconnected fluidized bed configuration. Air leakage from the AR into the FR will result in a dilution of the flue gas stream with N2, increasing extra cost of CO2 separation and casting away the advantages of CLC technology. The combustible gases leaking from the FR into the AR may cause CO2 release into the atmosphere, reducing the carbon capture efficiency of the CLC process and pose dangerous consequences even worse [17]. Therefore, a SP is designed to create a gas barrier in order to minimize gas leakage. Experiments on gas leakage

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Energy & Fuels

between reactors are conducted under working conditions shown in Table 5. As shown in Figure 6, the SP placed between the two reactor units can dramatically prevent gas mixing, and the gas leakage from the AR to the FR can be reduced to less than 1% while the gas leakage from the FR to the AR can be as far as zero. Table 5 Working conditions and parameters of gas leakage experiments Number

uAR,in （O2,m/s）

uFR,in （CO2,m/s）

uSP （N2,m/s）

uALS （N2,m/s）

uFLS （N2,m/s）

1

1.965-2.751

0.361

0.332

0.309

0.287

2

2.554

0.278-0.389

0.332

0.309

0.287

1.0

1.0

0.8

VAR-FR

0.8

VFR-AR

VAR-FR

0.6

0.6

VFR-AR

0.4

0.4 3

VAR-FR , VFR-AR (m /h)

VAR-FR , VFR-AR (m3/h)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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0.2 0.0 -0.2 -0.4

0.2 0.0 -0.2 -0.4

-0.6

-0.6

-0.8

-0.8 -1.0

-1.0 2.0

2.2

2.4

2.6

0.28

2.8

0.30

0.32

0.34

0.36

0.38

0.40

Superficial gas velocity of FR uFR,in (m/s)

Superficial gas velocity of AR uAR,in (m/s)

(a)

(b)

Figure 6 Gas leakage between reactors versus

superficial gas velocity of (a) AR and (b) FR

3.2 Pressure drops variation during combustion experiments The pressure drops in Figure. 7 reveals that pressure drops fluctuate stably when steady operation has reached. It is visible that there is a higher pressure drop in the FR-I while lower pressure drop is obtained in the FR-II, which resembles the configuration observed in the cold model. As discussed in section 3.1.1, this is because the FR-I has a larger oxygen carrier inventory which keeps a high level of dense bed in favor of prolonging the residence time of solid fuel particles and oxygen carrier. The mixing of oxygen carrier and solid fuel as well as its gasification products is thus able to be intensified and redox reactions can be carried out adequately. Highest pressure drop between the pressure tap 5 and 6 is obtained due to the existence of a mass of oxygen carrier. The gauge pressure at the bottom of the downcomer, p8, is also measured to monitor the height of bed materials in the long down pipe. The results in Figure. 7 (b) indicate that the pressure fluctuation in the downcomer is extremely stable during the steady operation period, which indirectly manifests that a satisfactory circulation of oxygen carrier is proceeding and the 14


Page 15 of 28

normal operation of CLC process is achieved. 8 8

Air reactor (∆p13)

7

∆p65

4

Fuel reactor stage I (∆p76)

2

Pressure drop (KPa)

6

6

Pressure drop (KPa)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

5 4 3 2 Downcomer (p8)

0

1

Fuel reactor stage II (∆p54) 0

-2 0

500

1000

0

1500

200

400

600

800

1000

Time (s)

Time (s)

(a)

(b)

1200

1400

1600

1800

Figure 7 Pressure drops measured during operating of CLC facility with hematite as oxygen carrier at 900℃. (a) Pressure drops between the pressure tap 5 and 6 and across the FR-I and the FR-II. (b) Pressure drops across the AR and the gauge pressure at the bottom of the downcomer.

3.3 Effect of reaction temperature In the process of SS chemical looping combustion, the solid SS particles were pneumatically transmitted into the FR at the bottom of the FR-I, where solid fuel particles were promptly heated up to the FR temperature and gasification began to happen simultaneously according to reactions [(R3)-(R5)]. The reactions between syngas and oxygen carrier particles taking place in the FR-I during the whole process which were shown as reactions (R6), (R7) and (R10), were endothermic reactions [7,18-20]

. Therefore, the oxygen carrier also served as heat carrier, which absorbed the

heat released from the reactions (R11) taking place in the AR and provided heat in the FR. Sewage sludge pyrolysis:

Sweage sludgedefNg → volatiles CO, H , CH> + char

(R3)

Char gasification:

char principally C + H O → CO + H

(R4)

char principally C + CO → 2CO

(R5)

Water-gas shift (WGS) equilibrium:

CO + H O → CO + H

(R6)

Reverse reaction of methanation:

CH> + H O → CO + 3H

(R7)

CO + 3Fe O → 2Fe O> + CO

(R8)

Combustion reactions:

15


Energy & Fuels

H + 3Fe O → 2Fe O> + H O

(R9)

CH> + 12Fe O → 8Fe O> + CO + 2H O

(R10)

Regeneration reactions:

4Fe O> + O → 6Fe O

(R11)

The circulation of hematite was insufficient to achieve autothermal process on account of short of heat supplied from regeneration reactions and heat dissipation to the ambience, thus an electric heating furnace was applied to maintain the desired temperature. In this section, the effect of reaction temperature varied in the range of 800-900℃ on the performance of the CLC process was investigated, keeping the SS feed rate at 400g/h.

3.3.1 Gas yield As shown in Figure 8, CO2 was the dominating component in both fuel reactors during the reduction period. However, its concentration decreased with temperature from 95.70% at 800℃ to 94.84% at 900℃ in the FR-I, but increased from 96.71% at 800 ℃ to 97.65% at 900 ℃ in the FR-II. CO and CH4 were also detected as unconverted products from incomplete oxidation by hematite. The rise of temperature from 800 to 900℃ resulted in an increase of the volume fractions of CO and CH4 in the FR-I, but reverse trend of fractions of CO and CH4 was observed in the FR-II. Furthermore, the fractions of combustible gases (CO and CH4) in the FR-I were a bit higher than those in the FR-II at each temperature point during the temperature rising process. 100

100

98

98

Gas concentrations in the FR-II (%)

Gas concentrations in the FR-I (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 16 of 28

CO2

96 94

CH4

4 3 2

CO

1 0

96

CO2

94

4 3

CH4

CO

2 1 0

800

820

840

860

880

900

800

Reaction temperature (℃)

850

900


(a)

(b)

Figure 8 Gas concentrations of (a) FR-I and (b) FR-II as a function of reaction temperature

The variations of the gas fraction were the consequence of combination of competition among different reactions, thermodynamics and kinetics of chemical reactions and gas-solid mixing. The devolatilization of SS would be well proceeded

16


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Energy & Fuels

even at a relative lower temperature (800℃) compared to char gasification.[7,21,22] As for SS with with most carbon being in the form of the volatile, a large proportion of SS fed into the FR-I was pyrolyzed rapidly leading to a formation of large gas bubbles or even a jet at the feeding point

[15]

. The contact between combustible gases and hot

oxygen carrier was thus restrained, and the SS particles which were not ready for pyrolysis were entrained by gas bubbles, bypass the bed and finally were gasified in the freeboard in the FR-I. As a consequence, the gasification products had no access to contacting with oxygen carrier and therefore cannot be converted to CO2 and water leading to an increase of CO2 in the FR-I, which was in agreement with the experimental consequences listed in Ref. [23, 24] The unconverted gasification products from the FR-I entered into the FR-II through the gas distributors and mixed up with oxygen carrier particles from AR. These oxygen carrier particles were freshly oxidized and had a higher reactivity, and no large bubbles are generated due to the limited gasification of fuel particles in the FR-II. Thus the fractions of CO and CH4 in the FR-II decreased with the rise of reaction temperature. After reaction the reduction state of oxygen carrier is generated and its reactivity decreases. This can be explained by the oxygen-carrying capacity of different states of oxygen carrier. The more oxygen is released in the early stage, the less oxygen is released in the later presenting lower reactivity. In fact, there is an interaction between the reduction gases concentrations in the FR-I and the reductive degree of oxygen carrier in the FR-II. The large amount of combustible gases leads to the deep reduction of oxygen carrier in the FR-II. And the deep reduction state of oxygen carrier with lower reactivity flowing from the FR-II into the FR-I results in the high concentrations of combustible gases in turn. In a manner of speaking, both the fuel type and the two-stage FR design lead to this special phenomenon. It is worth noting that the fraction of H2 was not detected in both FR-I and FR-II during the process and it was possibly due to the higher reducibility of H2 compared with CO and CH4 [7,21].

3.3.2 Carbon capture efficiency rss During the CLC process, it is inevitable that part of carbon which was incompletely converted to carbonaceous gases through gasification was entrained into the AR with the external circulation of oxygen carrier. Carbon in the AR was further oxidized to CO2 by oxygen in the air, which led to a small fraction of CO2 in the AR flue gas.

17


Energy & Fuels

Figure 9 exhibited the carbon capture efficiency as a function of the reaction temperature. The carbon capture efficiency increased from 97.04% at 800℃ with the rise of reaction temperature and it reached 98.93% at 900℃.The carbon capture efficiency demonstrated a high level on the whole and the reason of this can be explained by reaction kinetics process and specific design of the CLC facility. Firstly, with the increase of the reaction temperature, the char gasification process was enhanced as mentioned before and therefore the residual carbon in the FR-I was decreased. Consequently, carbon carried to the AR was reduced simultaneously which resulted in the increase of carbon capture efficiency. Moreover, different from the single-stage FR design, the unconverted fine carbon particles in the FR-I were pneumatically transported upward to the FR-II for further gasifying. This led to a decrease of residual carbon in the FR-I and carbon transferred to the AR, which was also responsible for the reduction of CO2 fraction at the outlet of AR and the increase of carbon capture efficiency. Thus it can be seen that both high temperature and two-stage FR design are beneficial to the prospective carbon capture efficiency. The reaction temperature above 900℃ can be suitable for SS combustion in this CLC system. 99.5

Carbon capture efficiency (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 18 of 28

99.0

98.5

98.0

97.5

97.0

96.5 800

820

840

860

880

900


Figure 9 Carbon capture efficiency of SS at different reaction temperature

3.3.3 Carbon conversion efficiency rs,stu and carbon supplementation efficiency vw,xy Figure 10 illustrated the effect of reaction temperature on the carbon conversion efficiency for SS. In the temperature range of 800-900℃, ?),)!@ increased with increasing the reaction temperature. 18


Page 19 of 28

Owing to the intense heat and mass transfer between fuel and oxygen carrier under high temperature, the SS gasification, especially char gasification process, was dramatically enhanced to generate more carbonaceous gases. Besides, the high temperature especially above 800℃ also improved the reaction rate between oxygen carrier and reduced gases, thus promoting the gasification of char

[22-24]

. Eqs. (9) and

(10) were qualitative description of fuel conversion degree without considering the unconverted carbon transferred to the AR and carried out with ash in the flue gas. A higher reaction temperature is necessary to attain higher carbon conversion efficiency due to the reduction of residual carbon in the ash. 100

Carbon conversion efficiency (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

95 90

ηc,con

85 80 75 70

ηc,con,I

65 60 800

820

840

860

880

900


Figure 10 Carbon conversion efficiency of SS at different reaction temperature

Figure 11 presented the carbon supplementation efficiency ΩR,ST versus the reaction temperature. It is noticed that higher carbon supplementation efficiency was obtained at lower temperature and decreased with the rise of reaction temperature above 850 oC. This is owing to the insufficient char gasification in the FR-I at lower temperature. And the unconverted fine char was further gasified to carbonaceous gases in the FR-II. As stated above, higher temperature contributed to a higher char gasification rate, especially above 850 ℃. Accordingly, char gasification fraction in the FR-II decreased, which was responsible for the low carbon supplementation efficiency at high temperature.

19


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Carbon supplementation efficiency (%)

Energy & Fuels

Page 20 of 28

22

20

18

16

14

12 800

820

840

860

880

Temperature of fuel reactor (

900

)

Figure 11 Carbon supplementation efficiency of SS at different reaction temperature

3.3.4 Combustion compensation efficiency rz,{| During the CLC experiments with SS as solid fuel, the combustion compensation efficiency as a function of reaction temperature was shown in Figure 12. It can be clearly seen that the combustion compensation efficiency substantially increases with the increasing reaction temperature and achieves maximum efficiency to 46.5% at 900℃. The conversion of combustible gases is not very obvious at lower temperature because the gasification process is not so notable and the reaction between oxygen carrier and combustible gases is at a relatively low rate. In the FR-I, a spouted fluidized bed, the produced gasification products could not be converted completely with insufficient fresh oxygen carrier. With the increase of reaction temperature, the gasification process is enhanced and thus combustible gases production augment which can also be seen in Figure 8. The freshly oxidized state of oxygen carrier in the FR-II has an increasing reactivity with the rise of reaction temperature, leading to a sharply increase of combustion compensation efficiency.

20


Page 21 of 28

50

Combustion compensation efficiency (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

45 40 35 30 25 20 15 10 5 800

820

840

860

Reaction temperature (

880

900

)

Figure 12 Combustion compensation efficiency of SS at different reaction temperature

3.4 Effect of SS feed rate The effect of the ratio of SS to oxygen carrier in the course of CLC process was also investigated by varying the SS feed rate from 300-600 g/h, keeping reaction temperature at 900℃.

3.4.1 Gas compositions of both two-stage FR and AR The influence of SS feed rate on gas compositions of both two-stage FR and AR was displayed in Figure 13. As we can see, the concentrations of CO and CH4 increased with the SS feed rate in both FR-I and FR-II whereas the CO2 fraction decreased. In the AR, the CO2 concentration slightly increased while the O2 composition in the flue gas decreased from 17.49 % to 9.71 %. SS particles were pneumatically transported into the bottom of FR-I, where SS pyrolyzed to combustible gases and char showed as (R3). And then a series of reactions (R4)-(R7) happened to produce syngas. When the feed rate of SS was small, the majority of carbonaceous gases can be converted to CO2 resulting in the high concentration of CO2 and low fractions of CO and CH4，because sufficient lattice oxygen was supplied by oxygen carrier to carbonaceous gases to CO2. With the further increasing of feed rate, the gasification products increased and lattice oxygen provision was becoming so insufficient that part of gasification products cannot be converted to CO2, therefore leading to a descent of carbon dioxide. Due to inadequate gasification in the FR, the redundant char was entrained to the AR, where char burnt with oxygen in the air to form carbon dioxide. Consequently, the oxygen concentration decreased and the CO2 fraction increased with the rise of SS feed rate, as shown in Figure 13(c). Furthermore, the ascensional range of carbonaceous gases 21


Energy & Fuels

and descending range of CO2 in the FR-II were smaller than those in the FR-I, respectively, which can also prove that the FR-II has a favourable effect on the carbonaceous conversion. 100

100

95

95

90

90

85

Gas concentration in the FR-II (%)

Gas concentration in the FR-I (%)

CO2

80 75 70

CO 20 15 10

CH4

5 0 300

350

400

450

500

550

CO2

85 80 75 15

CO 10

CH4

5 0

600

300

350

400

450

500

Sewage sludge feed rate (g/h)


(a)

(b)

550

600

18 16

Gas concentration in the AR (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 22 of 28

14

O2

12 10

0.2

CO2 0.1

0.0 300

350

400

450

500

550

600


(c) Figure 13 Gas compositions in the (a) FR-I, (b) FR-II and (c) AR as a function of SS feed rate.

3.4.2 Carbon conversion efficiency Carbon conversion efficiency of FR-I and the whole CLC system were investigated to evaluate the effect of SS feed rate on the performance of CLC process. As indicated in Figure 14, the carbon conversion efficiency of FR-I and the whole CLC system both decreased smoothly with the increasing of feed rate. Higher carbon conversion efficiency was obtained at a relative lower SS feed rate. As stated above, the devolatilization rate of SS was quite fast at 900 ℃. With the further increasing of SS feed rate, devolatilization of SS occurred to generate a large volume of volatile gases which formed plenty of large bubbles in the bed. The superficial velocity increased in the meanwhile. Part of carbon were thus entrained out of the reactor and cannot be converted to syngas due to the limitation of residue time of char in the FR. However, the char carried into the FR-II can be partially gasified due to the high reactivity of oxygen carrier. The carbon conversion efficiency of the whole CLC

22


Page 23 of 28

system was improved compared to that in the FR-I, which showed the superiority of the two-stage FR design.

95

Carbon conversion efficiency (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

ηc,con

90

85

80

75

ηc,con,I

70 300

350

400

450

500

550

600


Figure 14 Effect of SS feed rate on carbon conversion efficiency.

3.5 Characteristics of the oxygen carrier in the two-stage FR It is believed that the performance of SS CLC process in the two-stage FR unit not only depends on the design of the CLC facility, but is also closely associated with the characteristics of the oxygen carrier. Therefore, to explain the effect of oxygen carrier reactivity on the performance of whole CLC process, the particle sample was extracted from the FR-I and FR-II using hematite as oxygen carrier after 8 h of continuous operation at the reaction temperature of 900 ℃ and the SS feed rate of 400g/h, respectively. The x-ray diffraction (XRD) patterns of fresh oxygen carrier samples and used ones in both FR-I and FR-II are shown in the Figure 15. Compared with fresh oxygen carrier, it is observed that the major crystalline phase identified in two samples from the FR was Fe3O4, whereas a portion of Fe2O3 was detected in the sample extracted from the FR-II. As seen from the pattern of sample in the FR-I, the diffraction peak of Fe2O3 decreased apparently while the diffraction peak of Fe3O4 increased in reverse, which was due to the further reduction of Fe2O3 to Fe3O4 in the FR-I. Additionally, minor of FeO was detected in the FR-I, which should be caused by the deep reduction of Fe3O4. This is because the reduction rate for the reaction from Fe2O3 to Fe3O4 is fast, but the subsequent reduction to FeO or Fe is very slow according to the some presented papers

[7, 31-33]

. The oxygen-carrying capacity of

Fe3O4 is weaker than that of Fe2O3, which had been proved by experimental results in the previous literature [34]. Thus it is sufficient to note that the oxygen carrier in the FR-II has a higher reactivity than that in the FR-I and thus results in higher carbon conversion efficiency of whole CLC process. Moreover, the diffraction peaks of

23


Energy & Fuels

quartz (SiO2) and aluminum oxide (Al2O3) were detected in the oxygen carrier samples from both FR-I and FR-II. The presence of these minerals was likely from the component of hematite as well as the SS ashes produced during the CLC process. For further study, the predominant crystalline compound was Fe2O3 and trace of SiO2 was detected in the fresh oxygen carrier. Moreover, the diffraction peaks of SiO2 and Al2O3 were more remarkable in the FR-II. We may infer that part of SiO2 and Al2O3 were from SS ashes because more ashes would be wafted to the FR-II by the fluidizing agent on account of their light weight. As reaction time went on, a series of redox reactions between the SS ash and oxygen carrier would occur causing a negligible contribution to the sintering, agglomeration and inactivation of oxygen carrier [19,20,28-30]. (A) Fresh OC

(B) OC in the FR-II

1

2

Intensity (CPS)

1

Intensity (CPS)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 24 of 28

1 1 1

1

1 1 4 1 4 5

1

10

20

40

1 2

1

50

2

1

4

30

5

15

21

1 25

4

1

2

(C) OC in the FR-I

1 1

4

2

60

70

1

2

80

90

10

20

30

2

1

40

2θ (°)

2θ (°)

(a)

(b)

2

2

2 1

1 4

4

50

25

2 22

3

60

70

80

Figure 15 XRD analysis of (a) fresh oxygen carrier and (b) used oxygen carrier in the FR-I and FR-II after 8 h of operation: 1: Fe2O3; 2: Fe3O4; 3: FeO; 4: SiO2; 5: Al2O3.

4. Conclusion A new type of CLC system with a two-stage FR was proposed. The gas-solid flow characteristics cold experiment and SS treatment using chemical looping combustion technology were performed in this CLC unit. Some conclusions can be obtained: 1. Stable pressure fluctuations in the reactors were obtained in the cold experiments. The circulation rate of the oxygen carrier increased with the rising of superficial gas velocity of SP, which is the key device of the whole CLC system. Gas leakage between reactors was hazardous to the chemical looping process, however the syphon placed between the two reactor units can dramatically prevent gas mixing and the gas leakage. The cold experiment verified the feasibility of the performance of the continuous CLC unit.

24


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Energy & Fuels

2. SS combustion using hematite as oxygen carrier were conducted in the continuous CLC unit with a two-stage FR. With the rise of reaction temperature, the variation of gas concentration in the FR-I and FR-II showed different tendency. The volume fractions of CO and CH4 increased and the volume fraction of CO2 decreased with the increase of reaction temperature in the FR-I, but reverse trend was shown in the FR-II. The carbon capture efficiency increased from 97.04% at 800℃ to 98.93% at 900℃, and the carbon conversion efficiency of the whole system increased from 87.80% at 800℃ to 90.74% at 900℃. The combustion compensation efficiency was also increased with the rise of reaction temperature while the carbon supplementation efficiency declined due to the decrease of unconverted fine char at high reaction temperature. Lower carbon conversion efficiency was obtained at high SS feed rate. 3. The characteristics of fresh oxygen carrier and the oxygen carrier in the two-stage FR were detected. The XRD patterns showed no ash-related problems and the crystalline composition of oxygen carrier extracted from the FR-I and FR-II testified the redox reactions occurred during the sewage sludge CLC process. Overall, the novel CLC unit with a two-stage FR possesses excellent performance compared with previous CLC facilities.

Acknowledgements The authors gratefully acknowledge the support of this research work by National Key R&D Program of China (2016YFB0600801) and the National Natural Science Foundation of China (Grants 51476029, 51276037, 51561125001 and 51406035).

Reference [1] George, Tchobanoglous, L. B. Franklin, and H. David Stensel. "Wastewater engineering: treatment and reuse." Edition 4th. Metcalf and Eddi Inc (2003). [2] Werther J, Ogada T. Sewage sludge combustion[J]. Progress in energy and combustion science, 1999, 25(1): 55-116. [3] Werle S, Wilk R K. A review of methods for the thermal utilization of sewage sludge: The Polish perspective[J]. Renewable Energy, 2010, 35(9): 1914-1919. [4] Poland A, Knutson J C. 2, 3, 7, 8-Tetrachlorodibenzo-p-dioxin and related halogenated aromatic hydrocarbons: examination of the mechanism of toxicity[J]. Annual review of pharmacology and toxicology, 1982, 22(1): 517-554.

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[5] Murakami T, Suzuki Y, Nagasawa H, et al. Combustion characteristics of sewage sludge in an incineration plant for energy recovery[J]. Fuel Processing Technology, 2009, 90(6): 778-783. [6] Olabemiwo F A, Danmaliki G I, Oyehan T A, et al. Forecasting CO2 emissions in the Persian Gulf States[J]. Global Journal of Environmental Science and Management, 2017, 3(1): 1-10. [7] Niu X, Shen L, Gu H, et al. Characteristics of hematite and fly ash during chemical looping combustion of sewage sludge[J]. Chemical Engineering Journal, 2015, 268: 236-244. [8] Mattisson T. Materials for chemical-looping with oxygen uncoupling[J]. ISRN Chemical Engineering, 2013, 2013. [9] Lyngfelt A, Leckner B, Mattisson T. A fluidized-bed combustion process with inherent CO 2 separation; application of chemical-looping combustion[J]. Chemical Engineering Science, 2001, 56(10): 3101-3113. [10] Adanez J, Abad A, Garcia-Labiano F, et al. Progress in chemical-looping combustion and reforming technologies[J]. Progress in Energy and Combustion Science, 2012, 38(2): 215-282. [11] Thon A, Kramp M, Hartge E U, et al. Operational experience with a system of coupled fluidized beds for chemical looping combustion of solid fuels using ilmenite as oxygen carrier[J]. Applied Energy, 2014, 118: 309-317. [12] Coppola A, Solimene R, Bareschino P, et al. Mathematical modeling of a two-stage fuel reactor for chemical looping combustion with oxygen uncoupling of solid fuels[J]. Applied Energy, 2015, 157: 449-461. [13] Berguerand N, Lyngfelt A. Design and Operation of a 10kWth Chemical-Looping Combustor for Solid Fuels–Testing with South African Coal[J]. Fuel, 2008, 87(12): 2713-2726. [14] Berguerand N, Lyngfelt A. The Use of Petroleum Coke as Fuel in a 10kWth Chemical-Looping Combustor[J]. International Journal of Greenhouse Gas Control, 2008, 2(2): 169-179. [15] Thon A. Operation of a System of Interconnected Fluidized Bed Reactors in the Chemical Looping Combustion Process[M]. Cuvillier, E, 2014. [16] Kunii D, Levenspiel O. Fluidization engineering[M]. Elsevier, 2013. [17] Hossain M M, de Lasa H I. Chemical-looping combustion (CLC) for inherent CO2 separations—a review[J]. Chemical Engineering Science, 2008, 63(18): 4433-4451.

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Energy & Fuels

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Combustion Performance of Sewage Sludge in a Novel CLC System

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