Biomass Gasification and Hot Gas Upgrading in a Decoupled Dual

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Biomass Gasification and Hot Gas Upgrading in a Decoupled Dual Loop Gasifier Guangyong Wang, Shaoping Xu, Chao Wang, and Junjie Zhang Energy Fuels, Just Accepted Manuscript • DOI: 10.1021/acs.energyfuels.7b00782 • Publication Date (Web): 29 Jun 2017 Downloaded from http://pubs.acs.org on July 2, 2017

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Energy & Fuels

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Biomass Gasification and Hot Gas Upgrading in a Decoupled Dual Loop Gasifier

2

Guangyong Wang, Shaoping Xu*, Chao Wang, Junjie Zhang

3

State Key Laboratory of Fine Chemicals, Institute of Coal Chemical Engineering, School of

4

Chemical Engineering, Dalian University of Technology, No.2 Linggong Road, Dalian 116024,

5

China

6

*Corresponding author, e-mail address: [email protected] (SP Xu)

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Abstract: A decoupled dual loop gasifier (DDLG) has been developed in which biomass

9

gasification and hot gas upgrading are separated into two parallel loops and so that could be

10

optimized individually. In the gasification loop, the gasifier is so designed that the contact between

11

volatiles and char is restrained and therefore the steam gasification of char is enhanced. In the

12

upgrading loop, both desulfurizer and tar reforming catalyst are used for desulfurization and tar

13

destruction, respectively. As in-bed desulfurizer, an iron-bearing olivine supported ZnO

14

(Zn/olivine) was prepared and tested in a fixed bed reactor. H2S sorption over ZnO, adversely

15

affected by H2O, was accompanied by evident reduction of ZnO and vaporization of Zn at 550 oC.

16

By contrast, no obvious ZnO reduction was observed at the same condition over Zn/olivine. The

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reduction-resistance of Zn/olivine was illustrated by temperature programmed reduction and

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powder X-ray diffraction. In DDLG test with pine sawdust as feedstock and Zn/olivine+Ni/olivine

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as upgrading bed materials, a synergy was found between desulfurization and tar destruction. The

20

H2O-involved reactions such as steam gasification of char and steam reforming of

21

tar/hydrocarbons were intensified at elivated gasification temperature and in presence of

22

Ni/olivine. As a result, the decrease of H2O favored H2S sorption by Zn/olivine, which in turn

23

alleviated sulfur-poisoning of Ni/olivine. Under the gasifier temperature of 850 oC, steam to

24

biomass mass ratio (S/B) of 0.3 and upgrading reactor temperature of 600 oC, H2O and tar

25

contents were effectively decreased to 8.8% and 1.5 g/Nm3, respectively. In the 2 h test, during

26

which Zn/olivine experienced about 4 cycles of sulfidation/regeneration, H2S in product gas was

27

lowered to 1.7 ppmv.

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Keywords: Biomass gasification; Desulfurization; H2O conversion; ZnO; Olivine

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1. Introduction

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Synthetic natural gas (SNG) production from biomass with the advantages of sustainability and

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CO2 neutrality drew a lot of interest recently1-3. N2-free H2-rich syngas with 10-15 vol% CH4

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content produced by biomass gasification in dual bed gasification system could be a suitable

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feeding gas for bio-SNG production4-7. However, there are generally trace amounts of

35

sulfur-containing compounds and high molecular weight organic compounds (tar), and a large

36

amount of steam in the raw biogenous syngas. For example, in the raw gases from biomass

37

gasification in fluidized beds up to a few hundred ppmv of sulfur (mainly as H2S) and an average

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tar loading of about 10 g/Nm3 can be found8-12. These major impurities could deactivate the

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downstream methanation catalyst, particularly the Ni-based catalysts (by poisoning and coking),

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so they must be removed by gas-cleaning inside or downstream the gasifier.

41

Low temperature scrubbing downstream the gasifier is commonly adopted to remove tar,

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moisture and sulfur compounds13, 14, in which the syngas is cooled, cleaned at relatively low

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temperature and reheated before being sent to methanation reactor. Compared to the

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gasification-scrubbing-methanation process, an integrated gasification-methanation process with

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hot gas-cleaning is highly expected to significantly improve the efficiency of the SNG production

46

by avoiding cooling and reheating of the syngas10,

47

Specifically, sulfur removal could be achieved by high temperature sorption and tar could be

48

eliminated by reforming/cracking with the steam in the raw gas.

15

and the subsequent wastewater stream.

49

Incompatibility among the involved reactions is the main obstacle of the integrated

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gasification-methanation process. The biomass/char gasification is subject to the heat and mass

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transfer intensity between fuel particle and its surrounding heat carrier and gasification agent16, so

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that fine particles of the fuel and solid heat carrier and high velocity of the gasification agent are

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preferred. The raw product gas upgrading, on the other hand, desires ordinarily enough residence

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time of the raw gas on the catalyst or sorbent surface17, and consequently a low gas velocity is

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needed. In addition to their flow regime difference, there are remarkable difference in the

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operating temperature range between biomass/char gasification and raw gas upgrading.

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Biomass/char steam gasification with acceptable rate happens at about 800 oC or higher18.

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However, sulfur compounds elimination to sufficient low levels is difficult at such high 2


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temperature8, 19. The exothermic water gas shift reaction (WGS) is inhibited in this condition as

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well.

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Decoupling raw gas upgrading from biomass/char steam gasification seems to be necessary.

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Biomass steam gasification on dual fluidized bed gasifier was a preferential choice for bio-SNG

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production13, 14, during which gasification and combustion were decoupled into two reactors. The

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endothermic biomass steam gasification was balanced by the exothermic char air combustion with

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the help of circulating bed material, avoiding syngas consuming by the introduced O2 and the

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product gas dilution by the CO2 generated in combustion and the N2 introduced with air5, 20.

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However, a series of reactions, such as biomass pyrolysis, char gasification, tar steam

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reforming/cracking and WGS, were still intertwined in the gasifier. As a result, insufficient tar

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elimination even with in-situ cracking/reforming catalyst21, 22, suppressed char steam gasification

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by the volatiles-char interactions23, 24, limited WGS reaction at relatively high temperature and

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short residence time25, etc. were inevitable. Especially, more than 90% of the steam used to

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fluidize the bed materials was not converted and the H2O content in raw gas reached up to

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35-60%9, 12, 26, which decrease significantly the overall thermal efficiency of the process. More

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importantly, the moisture was unfavorable for desulfurization over the commonly adopted

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ZnO-based sorbents, especially at the relatively high temperature needed for biomass

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gasification27,

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pyrolysis/gasification has been tried by introducing a reforming zone in the gasifier or a

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downstream reformer29-35. In this respect, a decoupled triple bed gasification (DTBG) system

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characterized by separated pyrolysis/gasification, tar cracking/reforming and char combustion

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reactions to benefit tar destruction, has been proposed in our previous studies6, 36, 37. In the DTBG,

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however, the improvement on tar elimination was still unsatisfactory because that the gasification

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and tar destruction were compromised in one loop and occurred over identical bed material.

83 84

28

.

To

optimize

tar

destruction,

decoupling

of

tar

destruction

from

On the other hand, hot biogenous syngas desulfurization with either in-situ or downstream desulfurizer has been paid little attention10, 15, 27.

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As indicated by various works on desulfurization of the hot gas from coal gasification, sulfur

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removal from the hot syngas could be achieved with reusable metal oxide, such as the oxides of

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Zn, Fe, Ce, Ca, Cu, and Mn19, 38-40. ZnO showed good potential for H2S removal and has been 3


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widely studied. However, it suffered from reduction with H2 and zinc volatilization over 550 °C8.

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ZnO-based multi-metal oxides with improving activity and extending operation temperature were

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then paid much attention41-43. Typically, zinc ferrite sorbents, i.e. with a crystallite spinel structure

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formed by ZnO and Fe2O344, 45, could effectively reduce sulfur compounds to less than 1 ppmv at

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450 oC46. H2S level to less than 5 ppmv over zinc ferrite in Lurgi gas at temperatures as high as

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760 °C was achieved during multiple sulfidation-regeneration cycles47. Volatilization of the metal

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zinc happens in a way similar to the vapor-phase transport of Ostwald ripening process48, which

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could be suppressed by using the support with higher total activation energies (determined by the

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intrinsic interaction between the metal and support) for both the metal adatoms and metal-reactant

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complexes49. More importantly, the interaction between active component and the support could

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help to increase the reduction temperature of the supported active component (ZnO) and thus

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inhibit the formation of volatilizable metal (Zn)50-53. So, intensifying the active component-support

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interaction could be another promising approach to improve the stability of ZnO54. In addition, the

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active components could also be utilized more efficiently by increasing its dispersion in a

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supported form and thus decreasing the diffusional path length of H2S. As it was reported that the

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sulfidation reaction stopped before the whole spherical ZnO-pellets were reacted at temperatures

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below 600 °C because of the grain boundary diffusion27, 55, 56. The support plays an important role

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in maintaining structural strength of the desulfurizer as well46.

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The relatively cheap and abrasion-resistant iron-bearing olivine, containing (MgxFe1-x)2SiO4 as

107

the main phase and small quantities of MgSiO3 and FeOx species, was widely used in dual bed

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gasification of biomass as active bed material or catalyst support for in-bed tar reforming4, 57.

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Thermal induced FeOx formation over calcined olivine was detected58-60. As the catalyst support,

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supported components (Ni, Fe and Cu) were found to be bonded to or even included inside the

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olivine structure during calcination, suppressing their reduction54,

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Fe-bearing olivine could be a potential support for ZnO, providing suitable iron source for zinc

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ferrite formation and possibly improving its stability as well.

57, 61

. Thus, the calcined

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In present work, a decoupled dual loop gasifier (DDLG) has been developed to facilitate

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biomass gasification and in-bed hot gas upgrading for qualified biogenous syngas production. By

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introducing an additional upgrading loop parallel with the gasification-combustion loop, raw gas 4


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upgrading,

i.e.

desulfurization,

could

be

optimized

independently

under

suitable

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reaction/regeneration temperature, gas-solid contact pattern and even the atmosphere. As in-bed

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desulfurizer, an iron-bearing olivine supported ZnO (Zn/olivine) has been prepared and tested in a

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fixed bed reactor. The abrasion-resistant and regenerable Zn/olivine was then adopted in DDLG in

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addition to the optimization of its reaction conditions.

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2. Experimental

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2.1 Desulfurization of syngas in a fixed-bed reactor

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Several metal oxides, i.e. CeO2, ZnO, calcined limestone (CaO) and calcined olivine (Fe2O3),

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and an olivine supported ZnO were used as H2S sorbent. The CeO2 and ZnO from Sinopharm

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chemical reagent Co., Ltd were used after dry-pressed into tablets which were then crushed and

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sieved into 20-40 mesh range. The limestone (20-40 mesh) came from Changxin Zhejiang and was

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calcined at 900 oC for 4.5 h. It contains 91.5 wt.% CaO. The olivine was received from the

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Chinese city of Yichang60 and calcined at 800 oC for 4.5 h. The olivine supported ZnO

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desulfurizer, i.e. Zn/olivine-750, has 6 wt.% zinc loading and was prepared by incipient wetness

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impregnation of the 800 oC calcined olivine with an aqueous solution of zinc nitrate, followed by

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calcination at 750 oC for 4.5 h. An olivine supported 6 wt.% nickel, i.e. Ni/olivine-350, was used

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as the methanation catalyst. It was prepared by incipient wetness impregnation of the 600 oC

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calcined olivine with an aqueous solution of nickel nitrate and calcination at 350 oC for 4.5 h60.

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The desulfurization of syngas with the H2S sorbent and the downstream methanation on a

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nickel-based catalyst were performed at 550 oC and atmospheric pressure in an electric heated

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fixed-bed quartz reactor with 8 mm inner diameter. The desulfurization performance of the

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H2S-sorbent was evaluated based on the catalytic methanation activity of the catalyst given that

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the H2S passed through the sorbent could definitely poison the catalyst62. For metal oxide sorbents

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comparison, 0.5 g sorbent was loaded in the up-layer combined with a feeding gas flow of 130

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ml/min; for sulfidation-regeneration evaluation, 1.0 g of the Zn/olivine sorbent and a feeding gas

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flow of 52 ml/min were adopted. During metal oxide sorbents comparison, relative lower sorbent

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weight and higher gas flow rate were adopted in order to obtain clear distinction in sulfidation

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over different sorbents within a short reaction time. The methanation catalyst in the down-layer

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was 1.0 g for both cases. The feeding gas was a mixture of H2 and CO2 (H2/CO2 = 4) with 500 5


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ppmv H2S, named as FG-1, or a mixture of H2 and CO (H2/CO = 1) with 500 ppmv of H2S, FG-2.

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They both were provided by Dalian Special Gases Co., Ltd. The gas was controlled by a mass

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flow controller to flow downward through the sequentially loaded H2S sorbent and methanation

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catalyst. To illustrate the influence of H2O on desulfurization, H2O injected by a syringe pump and

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preheated to 250 oC was added to the feeding gas before entering into the reactor. The influence of

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heat and mass transfer limitations in the mathanation catalyst bed has been proved to be

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negligible60. The concentrations of CO, CO2, CH4 and other gas components in the product gas

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were monitored using a GC7890Ⅱ gas chromatograph equipped with both TCD and FID

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detectors. The CO or CO2 conversion and CH4 selectivity were calculated based on the product

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gas composition. CO2 conversion decreased to 90% of the initial value was used as the threshold

156

to determine the breakthrough of sorbent sulfidation. Sulfur capacity, in (mmol of H2S)/(g of

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sorbent), was calculated based on sorbed H2S arose from the different feeding H2S until the

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breakthrough in the presence and absence of the H2S sorbent62. CO out

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CO conversion (%)=(1-

160

CO2 conversion (%)=(1- CO2 out )×100

CO in

) ×100

(1)

CO

(2)

2 in

161 162 163

CH4 yield

CH4 selectivity (%)= CO

reacted

CH4 yield

×100 or CO

×100

(3)

2 reacted

-Fig. 12.2 Biomass gasification and hot gas upgrading in DDLG

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As shown in Fig. 2, DDLG consists of two circulation loops, i.e. a gasification loop and an

165

upgrading loop with fine and coarse particles as circulating bed materials, respectively. Biomass

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gasification happens in the bubbling fluidized bed gasifier of the gasification loop. The hot syngas

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from the gasifier is further upgraded, typically, desulfurization and tar steam reforming, in the

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moving bed upgrading reactor of the upgrading loop, which also acts as an efficient particulates

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filter so that a dust-free product gas could be obtained. Both loops have an identical fast fluidized

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bed riser followed by a separator. The fine bed material, i.e. SiO2, with the residual char from the

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gasifier and the coarse bed material, i.e. desulfurizer and tar reforming catalyst, from the

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upgrading reactor are together transported by air through the riser into the separator. In the

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separator, the fine and coarse bed material particles from the riser are separated from the flue gas 6


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and then fluidized by the air flow introduced into the regenerator. The coarse particle drops down

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through the bottom tube of the separator into the regenerator and the fine particle outflows the

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separator into the combustor based on their different masses and terminal velocities. In the

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combustor, the char is burnt out with air and the fine bed material is heated and returned to the

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gasifier to provide heat the endothermic gasification needed. In the regenerator, the used

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sorbent/catalyst are refreshed with air and delivered back to the upgrading reactor. The quality and

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quantity of fine and coarse circulating bed materials in both loops are monitored. In this way, the

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intricate reaction network is decoupled and the gasification and the raw syngas upgrading could be

182

optimized independently.

183

-Fig. 2-

184

Schematic of the lab-scale DDLG facility is shown in Fig. 3. The gasifier is a bubbling fluidized

185

bed reactor with a down-zone of 56 mm i.d. and 80 mm height and an up-zone 98 mm i.d. and 190

186

mm heigh. Biomass was fed into the freeboard of the gasifier rather than into the bubbling

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fluidized bed material to diminish the contact between volatiles and char and thus to decrease the

188

inhibition of volatiles to the steam gasification of char23, 24 and improve carbon conversion. The

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upgrading reactor is a gas-solid countercurrent moving bed reactor with an i.d. of 123 mm and a

190

height 168 mm. The riser is a fast fluidized bed reactor with 20 mm i.d. and 2600 mm height. The

191

combustor and regenerator are moving bed reactors, with 80 mm i.d. / 140 mm height and 80 mm

192

i.d. / 190 mm height, respectively. All the reactors are made of 310S stainless steel and heated by

193

independent electrical furnaces to compensate heat loss. The circulating rate of fine bed material

194

and that of coarse bed material are controlled by the rotary valve between the gasifier and the riser

195

and that between the upgrading reactor and the riser, respectively.

196

-Fig. 3-

197

Silica sand (SiO2) with the particle size of 60-100 mesh (0.15-0.25 mm) was used as the fine

198

circulating bed material in the gasification loop. Silica sand of 20-40 mesh (0.38-0.83 mm) was

199

used as the coarse circulating bed material in the upgrading loop, incorporating desulfurizer and

200

tar reforming catalyst of the same particle size for hot raw gas upgrading. The olivine supported

201

with 6 wt.% zinc, i.e. Zn/olivine-850, and the olivine supported with 6 wt.% nickel, i.e.

202

Ni/olivine-850, were used as the desulfurizer and the tar reforming catalyst, respectively. The 7


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203

Zn/olivine-850 was prepared by incipient wetness impregnation of the 800 oC calcined olivine

204

with an aqueous solution of zinc nitrate followed by calcination at 850 oC for 4.5 h. The

205

Ni/olivine-850 was prepared by incipient wetness impregnation of the 1000 oC calcined olivine

206

with an aqueous solution of nickel nitrate followed by calcination at 850 oC for 4.5 h. The biomass

207

feedstock used in this work was pine sawdust from Dalian City, Liaoning Province and its

208

proximate and ultimate analyses are presented in Table 1. The pine sawdust was crushed and

209

sieved to 20-40 mesh and dried for 4 h at 105-110 °C before test.

210

-Table 1-

211

Prior to test, 3.0 kg of the fine bed material and 4.5 kg coarse bed material were added into the

212

reaction system. The bed materials were bubbling fluidized and fast fluidized/transported with air

213

in the gasifier and riser, respectively. The circulating rate of fine bed material was maintained to

214

be 3.0 kg/h and that of coarse bed material 8.0 kg/h. All of the reactors were electrically heated to

215

the desired temperature and then fluidization air of the gasifier was replaced by the preheated

216

steam. When the temperatures of all reactors reached steady, gasification began by feeding pine

217

sawdust into the gasifier at a rate of 0.3 kg/h. The product gas was extracted from the upgrading

218

reactor with the help of a vacuum pump and cooled in four sequential glycol-cooled (-12 oC)

219

condensers, in which condensable components and particulate matter were separated from the

220

permanent gas. The effluent product gas underwent further aerosol removal and was measured by

221

a wet type gas flowmeter and analyzed every 10 minutes by a gas chromatography GC-7890Ⅱ

222

equipped with a TCD and a FID. The flue gas from the separator was cooled down and the

223

entrained dust was scrubbed by a venturi gas scrubber. After about 1 hour, when the reaction

224

system reached a steady state (Fig. 10), tar and H2S were separately sampled via two sampling

225

points before the condensers. The tar sampling was performed by condensing and dissolving the

226

tar components out of the product gas with six 250 ml impinger bottles basing on the protocol

227

CEN/TS 1543963-65. The impinger bottles were located in a cooling bath cooled down to -12 °C by

228

a cryostat, with bottles 1-5 filled with toluene and bottle 1, bottle 5 and bottle 6 containing glass

229

beads. The liquid phases in the impinger bottles were unified after sampling and the aqueous

230

phase was separated from the toluene phase. The amount of water was then determined to

231

calculate H2O content in the product gas and water conversion. The tar was obtained by 8


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evaporating off the toluene solvent at 50 oC under reduced pressure. The H2S was sampled by

233

separating the condensable components from the permanent gas in the impinger bottle filled with

234

glass beads, cooled down to -12 oC. H2S in the effluent permanent gas was then measured by a

235

GC9790 equipped with a packed column (GDX-303, 2 m× 3 mm) and a FPD, with the accuracy of

236

0.1 ppmv. All the experiments were kept for at least 2 h. The general operation conditions are

237

summarized in Table 2.

238 239 240

-Table 2To evaluate the performance of the process, the parameters are defined by the following equations: Mass of carbon in product gas (kg)

241

Carbon conversion (%)= Mass of carbon in biomass fed into the system (kg) ×100 (4)

242

Dry gas yield (Nm3 /kgfuel, daf)= Mass of biomass of dry ash-free basis fed into the system (kg) (5)

243 244

Volume of dry product gas (Nm3 )

Tar content of product gas (g/Nm3 )=

Mass of tar in product gas (g) Volume of dry product gas (Nm3)

Mass of H2O in product gas (kg)

Water conversion (%)=(1- Mass of H O fed as gasifying agent (kg) )×100

(6) (7)

2

245 246

Cold gas efficiency (%)=

Lower heating value of product gas (MJ/Nm3 )×Gas yield (Nm3/kg) Lower heating value of biomass fed into the system (MJ/kg)

×100 (8)

2.3 Characterization of desulfurizer and catalyst

247

The desulfurizer was characterized by powder X-ray diffraction (XRD) and temperature

248

programmed reduction (TPR). The XRD was carried out on a Rigaku D/Max-2400 diffractometer

249

with a nickel-filtered Cu Kα radiation (0.15406 nm) and with a scanning rate of 4o per minute in

250

the 2θ range of 25-65o. The TPR was conducted on a Quantachrome ChemBET 3000

251

chemisorption analyzer with a TCD detector. Prior to a TPR measurement, the as-prepared sample

252

(0.2 g) was pretreated at 250 oC in Ar for 2 h, while 5 v/v% H2/Ar with the flow rate of 40 mL/min

253

was used during the TPR experiment.

254

3. Results and discussion

255

3.1 Desulfurization of syngas in fixed-bed reactor

256

Desulfurization of syngas FG-1 and FG-2 in a fixed-bed reactor was performed over different

257

metal oxide sorbents (ZnO, calcined limestone (CaO), CeO2 and calcined olivine (Fe2O3)), with

258

Ni/olivine-350 as the methanation catalyst. The CO/CO2 conversion and CH4 selectivity were 9


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259

determined after 0.5 h and are shown in Fig. 4. All the four sorbents exhibited distinct H2S

260

removal ability in FG-2 and postponed deactivation of the downstream Ni/olivine-350 by

261

sulfur-poisoning, compared with the Ni/olivine-350 exposed in FG-2 directly. In FG-1, however,

262

only ZnO showed relatively pronounced desulfurization performance and rapid deactivation of

263

Ni/olivine-350 by H2S happened with calcined limestone, CeO2 or calcined olivine as the H2S

264

sorbents.

265

-Fig. 4-

266

It has been suggested that H2O, as a product of the sulfidation reaction over metal oxides

267

sorbents, could limit H2S reaction with MeOx by two parallel routes: the shift of thermodynamic

268

equilibrium which promotes the hydrolysis of MeSx and the competitive strong adsorption of H2O

269

which blocks the diffusion path of H2S as well as the active sites of sulfidation27, 28, 66. The

270

potential H2O production by reverse water gas shift (RWGS) reaction in FG-1, as indicated in Fig.

271

5, could adversely affect sulfur removal. In FG-2, instead, much less H2O formed over the

272

sorbents (mainly by the sulfidation reaction itself), exerts limited influence on the desulfurization

273

performance. The superior desulfurization performance of ZnO compared with that of calcined

274

limestone, CeO2 and calcined olivine, especially in FG-1 atmosphere, indicates that ZnO could be

275

more suitable for deep sulfur removal from H2O-containing syngas.

276

-Fig. 5-

277

The influence of H2O content on desulfurization over 0.5 g ZnO at 550 oC was further studied

278

with FG-1 and 0-30% more H2O as feeding gas. The H2O feeding was stopped at 90 minutes. As

279

shown in Fig. 6, in the initial 90 minutes, CO2 conversion and CH4 selectivity decrease with the

280

increase of H2O content in the feeding gas. After that, the methanation performance of the catalyst

281

experienced 10% H2O feeding could be recovered, but those experienced more H2O feedings

282

could not. The more the steam feeding, the worse the recovered catalytic performance of the

283

catalyst. The results indicate that deep sulfur removal from feedgas containing up to 10% H2O

284

could be achieved over ZnO, while H2O contents as high as 30% could remarkably deteriorate the

285

desulfurization performance at 550 oC and result in a rapid deactivation of the downstream

286

Ni/olivine-350. Meanwhile, some white powder material was observed to deposit on the

287

downstream inside wall of reactor tube after the experiments, suggesting that ZnO was partially 10


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Energy & Fuels

288

reduced to metallic zinc by H2 and the produced Zn underwent vaporization and condensation in

289

the reactor67.

290

-Fig. 6-

291

As mentioned above, ZnO showed relatively good potential for H2S removal, nevertheless, it

292

worked unfavorably in H2O-rich atmosphere and was prone to reduction and loss in hot syngas as

293

well. Such defects limit its application for deep sulfur removal from the conventional hot

294

biogenous syngas with H2O contents upto 35-60%, especially as the in-bed desulfurizer.

295

3.2 Characterization of Zn/olivine and its desulfurization performance in fixed-bed reactor

296

To suppress ZnO reduction, it was supported on an iron-bearing olivine calcined at 800 °C for

297

4.5 h. The dehydration and extraction of iron oxides from the olivine structure onto surface during

298

calcination58 develops the porous structure of the olivine. Specifically for the olivine used here,

299

800 °C was shown to be the optimal calcination temperature to improve the iron oxides

300

formation60. The thermal induced FeOx shells could improve the dispersion of the supported ZnO

301

and provide opportunities for the interactions between iron oxides and ZnO as well. The XRD spectra of the 800 oC calcined olivine supported ZnO after further calcination at 750

302 303

o

304

Besides the main olivine phase58 maintained, the new peaks appeared at 31.7° and 34.4° (2θ)

305

could be attributed to the main reflection of ZnO45. The relative intensity of the peaks at 29.8° and

306

35.6° (2θ) increase a lot, indicating the formation of ZnFe2O4 whose main diffraction peaks are

307

somewhat overlapped with that of the olivine phase45, 68.

308

C for 4.5 h (Zn/olivine-750) compared with that of the calcined olivine is shown in Fig. 7.

-Fig. 7-

309

The TPR profiles of the calcined olivine and the Zn/olivine-750 are presented in Fig. 8. The

310

calcined olivine exhibits a broad reduction peak above 350 °C, which is mainly attributed to the

311

reduction of the free iron oxides associated with the olivine phase. It corresponds to about 3 wt.%

312

Fe0 in the reduced olivine calculated from the hydrogen consumption during TPR experiments, as

313

1 mole H2 could reduce 1/3 mole Fe2O3 which is the main form of reducible iron in the calcined

314

olivine60. A similar reduction peak starting at 350 °C and ending at 850 °C appears for the

315

Zn/olivine-750, with an additional reduction peak at higher temperature. The new peak (>850 °C)

316

is in good accordance with the reported ones during reduction and modification of ZnFe2O469. 11


Energy & Fuels

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

317

Notably, the potential reduction peak of ZnO is not so apparent67, 70. The slight decrease of H2

318

consumption below 850 oC of the Zn/olivine-750, compared with that of the calcined olivine,

319

could attribute to the strong interaction between ZnO and the support and the formation of

320

reduction-resistant structure (ZnFe2O4, as mentioned above), which may suppress the reduction of

321

the supported ZnO47, 54, 57, 61.

322

-Fig. 8-

323

Sulfidation and regeneration behaviors of the Zn/olivine-750 were evaluated in FG-1 at 550 oC

324

with the spent sorbent regenerated in air flow at 750 or 850 oC (Fig. 9). Meanwhile, sulfur

325

capacity of the Zn/olivine-750 was determined. The breakthrough point for the fresh

326

Zn/olivine-750 was 2.25 h, corresponding to a H2S capacity of 0.16 mmol/gZn/olivine. After

327

regeneration at 750 oC for 1 h, the sulfidation time till breakthrough of the refreshed

328

Zn/olivine-750 extended to 3.25 h, indicating that the sulfur removal ability was maintained and

329

even developed. To investigate the effect of regeneration condition, the second and the third

330

regeneration were conducted at 750 oC for 0.5 h and 850 oC for 0.5 h, respectively. As shown in

331

Fig. 9, after the second regeneration, the working sulfidation time was shorted to 2.75 h, indicating

332

that the regeneration at 750 oC for 0.5 h was insufficient. After the third regeneration, possibly

333

because of the improved regeneration rate at elevated temperature, the sulfidation time was

334

extended to 4 h, corresponding to a H2S capacity of 0.28 mmol/gZn/olivine. Obviously, the

335

Zn/olivine-750, underwent repeated sulfidation-regeneration cycles, was well recovered after

336

regeneration in air at 850 oC for 0.5 h.

337

Despite of the varied regeneration conditions, the reactivity of the Zn/olivine-750 appears to be

338

improved with the sulfidation-regeneration cycling, suggesting that the sulfidation of this sorbent

339

can be improved by activation treatments38. The enhanced desulfurization performance of the

340

refreshed Zn/olivine could mainly result from the thermal induced structural modification during

341

regeneration, i.e. iron thermal transformation in the olivine grain60 and possibly an improved

342

interaction between ZnO and olivine. Furthermore, the white material evolved during the

343

sulfidation of the ZnO pellets at 550 oC was not observed after desulfurization over the

344

Zn/olivine-750, implying an improved reduction-resistance of the supported ZnO as indicated by

345

the TPR test. 12


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Energy & Fuels

346 347

-Fig. 93.3 Biomass gasification and hot gas upgrading in DDLG

348

Low H2O conversion could greatly decrease thermal efficiency of the overall gasification

349

process and increase the cost, remaining a weakness of biomass steam gasification in dual bed

350

gasifier12. Furthermore, high H2O content is detrimental to the downstream utilization of the

351

as-produced hot gas, typically, deteriorating the hot gas desulfurization behavior as indicated

352

above. Therefore, less H2O feeding for bed material fluidization and intensifying H2O conversion

353

were attempted in DDLG to decrease the H2O content in the biogenous syngas.

354

The pressure drop across the gasifier at different H2O feeding rates, as one of the most

355

important gas-solid flow features, could help to determine the flow regime71. Variation of pressure

356

drop with increasing H2O feeding at 800 oC indicates that the bed material (SiO2 of 60-100 mesh)

357

starts to fluidize with H2O feeding rate of about 60 g/h. H2O feeding rates with the interval of 90

358

g/h to 300 g/h, corresponding to S/B of 0.3-1.0, were used in this work. Preliminary tests have

359

proved that the segregation of char in the gasifier was not apparent in this interval. The relatively

360

low superficial gas velocity in the gasifier, approaching to the minimum fluidization velocity of

361

the bed material, benefits the contact between gasification agent and biomass/char, as the mean

362

bubble diameter in the fluidized bed increases with the superficial gas velocity which would

363

deteriorate gas-solid contact25. In addition, excessive unreacted H2O effluent can be alleviated, as

364

it was reported that H2O conversions in the gasifier were generally maintained to be less than 10%

365

with S/B of 0.4-2.6 and the feeding H2O was significantly beyond the consumed12. In a typical

366

gasification test, the pressure drop and the product gas compositions as a function of time on

367

stream were shown in Fig. 10. The stably pulsing pressure drop indicates that the bed material was

368

sufficiently fluidized and bubbled vigorously. The gas composition reached a steady state after

369

about 1 hour.

370

-Fig. 10-

371

The optimization of biomass gasification over DDLG was then performed at varied S/B and

372

gasifier temperature. In these tests, SiO2 was used as bed materials in both loops and the

373

upgrading reactor temperature was maintained to be 350 ºC to avoid volatiles cracking and

374

condensation. 13


Energy & Fuels

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

375

Effect of S/B on biomass steam gasification is shown in Fig. 11. Carbon conversion and gas

376

yield increase with the S/B rising up to 0.7 and then decrease. Tar content decreases slightly and

377

product gas composition remains almost constant with the increase of S/B. The increasing steam

378

addition could promote char steam gasification at low S/B, which is responsible for the

379

improvement of both carbon conversion and gas yield. Increasing S/B to 1.0 from 0.7, however,

380

deteriorated the gasification performances, due to the unfavorable steam-char contact in the

381

gasifier at higher superficial gas velocity and heat removal by the increasing H2O effluent. H2O

382

conversion decreases monotonously with the increase of S/B, in accordance with the published

383

work72. H2O conversion remains below 18%, suggesting that the feeding steam was still evidently

384

exceeded the needed even with S/B as low as about 0.4, i.e. the lower limit of the recommended

385

S/B12.

386

-Fig. 11-

387

The gasifier temperature could greatly affect the biomass steam gasification, as indicated in Fig.

388

12. Carbon conversion, H2O conversion and gas yield increase with raising gasifier temperature,

389

while tar content shows reverse tendency, contributed by the enhanced char steam gasification and

390

steam reforming/cracking of tar at higher temperature. H2 concentration increases with the

391

increase of temperature and reaches to 36.8% at 800 oC, and then remains more or less unchanged.

392

Lower heating value (LHV) of the product gas increases slightly. From the point of view of

393

improving H2O conversion as well as product gas yield, higher gasifier temperature is preferred.

394

Nevertheless, it is very difficult to work in the gasifier above 900°C for biomass gasification with

395

pure steam12.

396

-Fig. 12-

397

Notably, the tar contents in the raw biogenous syngas from the gasifier of DDLG are of 100

398

g/Nm3 magnitude, remarkably exceeded the reported values from fluidized bed gasifiers11, 12 and

399

close to those from the updraft (counterflow) gasifiers11. This could be due to the specific design

400

of the gasifier. As indicated in Fig. 3, biomass was fed into the freeboard rather than the

401

commonly way into the bubbling fluidized bed material. Biomass underwent rapid pyrolysis upon

402

touching the hot bed material in surface layer of the bubbling fluidized bed. The as-produced

403

pyrolysis volatiles outflowed with the effluent steam and syngas from char steam gasification, 14


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Energy & Fuels

404

rather than flowed into the deeper zone of the fluidized bed. The nascent char from biomass

405

pyrolysis, instead, sank into the fluidized bed and mixed with the bed material sufficiently

406

(confirmed by the insignificant segregation of char in bed material). Consequently, in-situ tar

407

elimination was suppressed by the short residence time of the volatiles and the limited contact

408

between volatiles and the hot bed material (char and SiO2), resulting in tar-rich product gas. In

409

contrast, steam gasification of the nascent char dispersed in the fluidized bed and the derived H2O

410

consumption benefited from the restrained volatiles and char interactions, as the volatiles were

411

suggested to inhibit steam gasification of char23, 24. As expected, carbon conversion reached to

412

81.6% at the gasifier temperature of 850 oC and S/B of 0.4, approaching to or even exceeding the

413

results of the ordinary dual fluidized bed gasifiers7. Meanwhile, H2O conversion of 20.6% was

414

achieved over DDLG, corresponding to 27.0% H2O content in the raw gas which is less than the

415

published data from biomass steam gasification in fluidized bed10, 12.

416

In the configuration of DDLG, the as-produced volatiles is delivered to and upgraded in the

417

upgrading reactor under optimized condition. The improvements of the steam involving reactions

418

during volatiles upgrading, such as tar/hydrocarbon steam reforming and WGS, could promote

419

H2O conversion further and thus facilitate desulfurization of the raw syngas.

420

Thus, Ni/olivine-850 was included in the upgrading loop, acting as tar reforming catalyst, along

421

with the Zn/olivine-850 desulfurizer, for in-bed hot gas upgrading. Meanwhile, both the spent

422

Zn/olivine-850 and the deactivated Ni/olivine-850 underwent regeneration in the regenerator at

423

850 oC, as the sulfurized Zn/olivine-850 could be well recovered at 750-850 oC in air flow and the

424

potential carbon and even sulfur deposits over Ni/olivine-850 were suggested to be removed at

425

about 850 oC14.

426

As shown in Table 3, tar in the raw syngas was partially reformed over SiO2+Ni/olivine at the

427

upgrading reactor temperature of 600 oC. The catalytic activity was not good enough, mainly due

428

to sulfur-poisoning of nickel. Meanwhile, H2O conversion decreases a lot compared with that over

429

the inert bed material, because Ni/olivine-850 as an oxygen carrier is prone to be reduced by H2 to

430

produce H2O. In comparison, with the introduction of Zn/olivine-850 additionally, tar and gaseous

431

C2+ hydrocarbons were mostly eliminated. CH4 decreases and H2/CO ratio increases distinctly.

432

63.8% H2O conversion and H2O content as low as 8.8% were achieved. The results confirm that 15


Energy & Fuels

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

433

tar/hydrocarbon

434

SiO2+Ni/olivine+Zn/olivine and consequently H2O conversion was promoted as expected. In the 2

435

h test, during which Zn/olivine-850 experienced about 4 cycles of sulfidation/regeneration, the

436

H2S in product gas was lowered to 1.7 ppmv.

437 438

steam

reforming

and

WGS

were

Page 16 of 41

remarkably

intensified

over

-Table 3Conclusions

439

Biomass gasification and hot gas upgrading are decoupled in the DDLG system basing on their

440

different requirements on flow regime, operating temperature, etc. Both desulfurizer and tar

441

reforming catalyst are needed for desulfurization and tar destruction during in-bed hot gas

442

upgrading.

443

As in-bed desulfurizer, the Zn/olivine was prepared and tested in a fixed bed reactor. H2S

444

sorption over ZnO, adversely affected by H2O, was accompanied by evident reduction of ZnO and

445

vaporization of Zn at 550 oC in H2/CO2/H2O/H2S. By contrast, no obvious ZnO reduction was

446

observed at the same condition over Zn/olivine-750. The result confirmed the improved

447

reduction-resistance of the olivine supported ZnO as indicated by the TPR test, due to the

448

interaction between ZnO and the support and additionally the formation of ZnFe2O4 between the

449

supported ZnO and the thermal induced iron oxides on olivine. Meanwhile, H2S was efficiently

450

removed over Zn/olivine-750 at 550 °C and the spent Zn/olivine-750 could be well recovered at

451

750-850 °C during repeated sulfidation-regeneration cycles.

452

In DDLG test with pine sawdust as feedstock, optimization of the biomass gasification itself

453

showed that low S/B and high gasifier temperature benefited H2O conversion and thus diminished

454

the adverse effects of H2O on H2S sorption. With the introduction of Zn/olivine-850 and

455

Ni/olivine-850 as upgrading bed materials, a synergy was found between desulfurization and tar

456

destruction during hot gas upgrading. The H2O-involved reactions such as steam reforming of

457

tar/hydrocarbons and WGS were intensified in presence of Ni/olivine-850. As a result, the

458

decrease of H2O favored H2S sorption by Zn/olivine-850, which in turn alleviated sulfur-poisoning

459

of Ni/olivine-850. Under the gasifier temperature of 850 oC, S/B of 0.3 and upgrading reactor

460

temperature of 600 oC, H2O content and tar were effectively decreased to 8.8% and 1.5 g/Nm3,

461

respectively. In the 2 h test, during which Zn/olivine experienced about 4 cycles of 16


Page 17 of 41

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Energy & Fuels

462

sulfidation/regeneration, H2S in product gas was lowered to 1.7 ppmv.

463

Acknowledgments

464

This work is supported by the Natural Science Foundation of China (No. 50776013) and the

465

National High Technology Research and Development Program of China (No. 2008AA05Z407).

466

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pore structure of sorbents containing zinc ferrite. Energy & Fuels 1997, 11, (4), 887-896.

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in Ni/Al2O3 for syngas methanation at high temperatures. RSC Advances 2015, 5 (14),

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sorbents performances in the desulfurization process in link with biomass gasification. Fuel

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Processing Technology 2014, 118, 254-263.

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zinc ferrites as regenerable sorbents for hot coal gas desulfurization. Solid State Ionics 2000, 138

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gasification in a dual fluidized bed. Applied Catalysis B Environmental 2012, 121, 214-222.

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behavior of iron-bearing olivines (FexMg1-x)2SiO4 used as catalysts for biomass gasification.

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impregnated activated carbons for the production of SNG. Fuel Processing Technology 2015, 138

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project ERK6-CT1999-20002 (Tar Protocol).

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L. Guideline for sampling and analysis of tar and particles in biomass producer gases.

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ECN-C-02-090. URL.

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66. Novochinskii, II; Song, C. S.; Ma, X. L.; Liu, X. S.; Shore, L.; Lampert, J.; Farrauto, R. J.

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Low-temperature H2S removal from steam-containing gas mixtures with ZnO for fuel cell

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application. 1. ZnO particles and extrudates. Energy & Fuels 2004, 18 (2), 576-583.

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High-temperature reactivity of mechanosynthesized zinc ferrite. Solid State Ionics 1997, 101,

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varying column shapes. Powder technology 2013, 246, 561-571.

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665

URL.

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24


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Energy & Fuels

667

Figure captions

668

Fig. 1. Schematic of the fixed-bed test.

669

Fig. 2. Concept of the DDLG system.

670

Fig. 3. Schematic of the lab-scale DDLG facility.

671

Fig. 4. Methanation activity of Ni/olivine downstream different H2S sorbents at 550 oC in (a) FG-2

672

and (b) FG-1.

673

Fig. 5. Product gas distribution for RWGS reaction of FG-1 over ZnO at 550 oC.

674

Fig. 6. Methanation activity of Ni/olivine downstream ZnO in FG-1 with varied H2O addition: (a)

675

CO2 conversion and (b) CH4 selectivity.

676

Fig. 7. XRD spectra of the 800 oC calcined olivine and Zn/olivine-750.

677

Fig. 8. TPR profiles of the 800 oC calcined olivine and Zn/olivine-750.

678

Fig. 9. Breakthrough curves of the fresh and regenerated Zn/olivine-750 in FG-1 with 10 vol.%

679

H2O addition: (a) CO2 conversion and (b) CH4 selectivity.

680

Fig. 10. Product gas compositions and the pressure drop of the gasifier as a function of time,

681

biomass feeding rate 0.3 kgfuel,db/h, gasifier temperature 800 oC and S/B 0.7.

682

Fig. 11. Effect of S/B on biomass steam gasification with gasifier temperature at 800 oC.

683

Fig. 12. Effect of gasifier temperature on biomass steam gasification with S/B of 0.4.

25


Energy & Fuels

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

684

Table captions

685

Table 1. Proximate and ultimate analyses of the pine sawdust.

686

Table 2. General operation parameters of the DDLG system.

687

Table 3. Key parameters and results for biomass gasification with raw syngas upgrading over

688

varied bed materials.

689 690

26


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Energy & Fuels

691 692

Fig. 1. Schematic of the fixed-bed test.

693

27


Energy & Fuels

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

694 695

Fig. 2. Concept of the DDLG system.

696

28


Page 28 of 41

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Energy & Fuels

697 698

Fig. 3. Schematic of the lab-scale DDLG facility.

699

29


Energy & Fuels

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

700 701

Fig. 4. Methanation activity of Ni/olivine downstream different H2S sorbents at 550 oC in (a) FG-2

702

and (b) FG-1.

703

30


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Energy & Fuels

704 705

Fig. 5. Product gas distribution for RWGS reaction of FG-1 over ZnO at 550 oC.

706

31


Energy & Fuels

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

707 708

Fig. 6. Methanation activity of Ni/olivine downstream ZnO in FG-1 with varied H2O addition: (a)

709

CO2 conversion and (b) CH4 selectivity.

710

32


Page 32 of 41

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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

711 712

Fig. 7. XRD spectra of the 800 oC calcined olivine and Zn/olivine-750.

713

33


Energy & Fuels

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

714 715

Fig. 8. TPR profiles of the 800 oC calcined olivine and Zn/olivine-750.

716

34


Page 34 of 41

Page 35 of 41

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

717 718

Fig. 9. Breakthrough curves of the fresh and regenerated Zn/olivine-750 in FG-1 with 10 vol.%

719

H2O addition: (a) CO2 conversion and (b) CH4 selectivity.

720

35


Energy & Fuels

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

721 722

Fig. 10. Product gas compositions and the pressure drop of the gasifier as a function of time,

723

biomass feeding rate 0.3 kgfuel,db/h, gasifier temperature 800 oC and S/B 0.7.

724

36


Page 36 of 41

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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

725 726

Fig. 11. Effect of S/B on biomass steam gasification with gasifier temperature at 800 oC.

727

37


Energy & Fuels

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

728 729

Fig. 12. Effect of gasifier temperature on biomass steam gasification with S/B of 0.4.

730 731

38


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Energy & Fuels

732

Table 1. Proximate and ultimate analyses of the pine sawdust. Proximate analysis (wt.%, ad)

733

Moisture

Ash

Volatile

8.26

0.61

78.40

Fixed carbon 12.73

Ultimate analysis (wt.%, daf) Carbon

Hydrogen

Oxygen1)

Nitrogen

Sulfur

47.75

6.98

44.84

0.07

0.36

1) by difference

734 735

39


Energy & Fuels

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

736

Page 40 of 41

Table 2. General operation conditions of the DDLG system. Fine bed material

60-100 mesh SiO2

Fine bed material inventory (kg)

3.0

Fine particles circulating rate to biomass feeding rate (kg/kg)

10 20-40 mesh Zn/olivine,

Coarse bed material

Ni/olivine and SiO2

Coarse bed material inventory (kg)

4.5

Coarse particles circulating rate to biomass feeding rate (kg/kg)

27

Biomass feeding rate (kgfuel,db/h)

0.3

Steam to biomass mass ratio (S/B) (kgH2O/kgfuel,db)

0.3-1.0

o

Gasifier temperature ( C)

700-850 o

Upgrading reactor temperature ( C)

350-600

o

Combustor temperature ( C)

850

o

Regenerator temperature ( C)

850

o

Riser temperature ( C)

850 o

Steam generator temperature ( C)

800

Operating pressure

atmospheric

737 738 739

40


Page 41 of 41

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

740

Table 3. Key parameters and results for biomass gasification with raw syngas upgrading over

741

varied bed materials. Feeding rate (kg/h) S/B o

Gasifier temperature ( C) Fine bed material o

Upgrading reactor temperature ( C) Coarse bed material composition (wt.%) 3

Gas yield (Nm /kgfuel, daf) 3

LHV (MJ/Nm ) Cold gas efficiency (%)

0.30

0.30

0.30

0.4

0.3

0.3

850

850

850

SiO2

SiO2

SiO2

350

600

600

SiO2+Ni/olivine

SiO2+Ni/olivine+

(88/12)

Zn/olivine (76/12/12)

1.03

1.14

1.51

14.6

12.8

11.4

SiO2 (100)

81.2

78.7

92.9

3

Tar content (g/Nm )

61.2

23.2

1.5

H2O conversion /%

20.6

10.5

63.8

H2O content /%

27.0

24.0

8.8

H2S concerntration /ppmv

146.1

147.1

1.7

H2

34.9

41.0

47.3

CO

33.3

28.5

28.5

CH4

10.1

8.4

6.3

CO2

17.2

19.4

17.1

C2H4

2.7

1.6

0.3

C2H6

0.6

0.5

0.3

C3H6

1.0

0.5

0.1

C3H8

0.2

0.1

0.1

Dry gas composition /vol.%

742 743

41


Biomass Gasification and Hot Gas Upgrading in a Decoupled Dual

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