Design of a (Bio)process to remove nitrogen oxides from flue gases

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TU Delft

Design of a (Bio)process to remove nitrogen oxides from f1ue gases CPD 3332

CPD NR 3332

Conceptual Process Design

Process Systems Engineering

DelftChemTech - Faculty of Applied Sciences Delft University of Technology

Design of a (Bio )process to remove nitrogen oxides from

f1ue gases

Authors

H.J. van Rein

B.B. Wolf

Y.J. Juan

M. van de Weg

J. Nijenhuis

Keywords

Telephone

0649337043

0611443414

0625290500

0641288563

0652678940

Student nr.

1251163

1252402

9413836

1197819

1251171

NOx, Denitrification, Absorbance, Flue gas, Fe(II)EDTA

2-,

BioDeNOx

Assignment issued

Report issued

Appraisal

Conceptual Process Design: Removal of NO. from flue gases

: 01-02-06

: 21-04-06

: 28-04-06

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Design of a (Bio)process to re move nitrogen oxides from flue gases CPD 3332

Preface

This report describes the final results of the Conceptual Process Design project "Design of a (Bio)process to remove nitrogen oxides from flue gases'. The course Conceptual Process Design (CPD) is part of the Master's Program 'Chemical Engineering' at Delft University of Technology, provided by the department Process and Product Engineering.

Generating the final result seems impossible without the help of people outside the CPD team. Of course we would like thank everybody contributing to this end product.

Especially we would like to thank Dr.Ir. C. S. Bildea for the huge amount of effort and patience he put into this project. Also we would like to thank our creativity coach M. de Niet who weekly took the time to listen to our problems and guiding us towards the right direction. Also mister G. Muyzer for giving information about the biochemical part of the process. Delft, April 21 th 2006 CPD 3332 HJ. van Rein B.B. Wolf Y.J. Juan M. van de Weg J. Nijenhuis

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Process Equipment Summary

COl NO Absorber POl Solution to ROl POB: Cool W COl in SOl Rotary filter ROl NO Bioreactor P02 Solution to COl POg: Cool W ROl Jacket S02 Filter

E01 Flue gas inlet cooler P03 Acid/Base in PlO: Cool W E02 in S03 Spare filter

E02 Flue 9as Feed cooler P04 Ethanol in : TOl Acid/Base

E03 Solutlon to ROl heater X01 : Transporter screw T02 Ethanol

E04 Bioreactor Jacket P06 Nutrition's in X02: Transporter screw T03 Nutrition's POl Fe(II)EDTA in X03: Clean gas Chimney T04 Fe(II)EDTA

-Designers

CPD-3332

Solid

Bleed~

Process Flow Scheme - Improved Solution Project

Proj. ID Number Completion Date

o

Stream number

Design of a (bio) process to remove nitrogen oxides from f1ue gases CPD-3332

2B April 2006

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Acid

_P03

Ethanol

_P04

P07

Cooling Water to E02

lf,;\

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1 Flue gas feed

POl

Cooling Water to E01

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Cooling water bioreactor

Process Equipment Summary

COl: NO Absorber

ROl: NO Bioreactor

E01

: Flue gas inlet cooler

E02 : Flue gas Feed cooler

E03

: Solution to ROl heater

E04

: Bioreactor Jacket

POl: Solution to ROl P08: Cool W COl in

P02: Solution to COl POg: Cool W ROl Jacket

P03

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Acid/Base in

PlO: Cool W E02 in

P04: Ethanol in

501: Rotary filter

502: Filter

503: Spare filter

TOl: Acid/Base

P06

:

Nutrition's in

P07: Fe(II)EDTA in

XOl

:

Transporter screw T02: Ethanol

X02: Transporter screw T03: Nutrition's

X03: Clean gas Chimney T04: Fe(II)EDTA

. .. ... -\ ....... . ... .

Designers

CPD-3332

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Biomass

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Process Flow Scheme - Improved Solution

Project

Proj. ID Number

Completion Date

c::>

Stream number

Design of a (bio) process to remove

nitrogen oxides from flue gases

CPD-3332

28 April 2006

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ERATUM

page 21

"Flue-gas < 101> enters the heat exchanger 1 (first unit) at a temperature of approximately 673 K and is cooled down with cooling water to a temperature of around 353 K - 373K. This is not a suitable temperature for the absorption of NO in water." should be changed in:

Flue-gas < 101> enters the heat exchanger 1 (first unit) at a temperature of approximately 673 K and is cooled down with cooling water to a temperature of around 423K. This is not a suitable temperature for the absorption of NO in water.

page 21

"The solidsjwastes in the bioreactor are removed <403> <404> from the bottom of the bioreactor." should be changed in:

The solidsjwastes in the bioreactor are removed <403> <404> after the rotary filter SOL page 37

Design criteria in table 5.7 should be changed from 0.1-20 micron to 100 micron. page 37

For generating high-pressure steam (heat exchanger E01), demi-water is used instead of river water.

Page 39

In figure 5.1 (electricity should be 1.450'103 kWhjt instead of 1.450'103 kWhjh) page 45

Table should be in paragraph 7.7

Pump [kWl P01 3.76 P02 25.8 P03 0.00031 P04 0.0189 P06 0.0050 POl 0.0050 P08 0.87 POg 0.057 P10 0.29

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Design of a (Bio )process to remove nitrogen oxides from f1ue gases CPD 3332

Stream numbers Pipe inside diameter rml

<503> 0.1382 <501> 0.1396 <801> 0.0000 <201> 0.0021 <301> 0.0169 <401> 0.0005 <901> 0.0005 <601> 0.0169 <701> 0.0136 <603> 0.0425

The pipe inside diameter is calculated for carbon steel pipe 64.

page 47

"Unfortunately not the who Ie process could be modelled, but only the absorber and bioreactor separately." should be changed in:

The complete process, absorber and bioreactor, is modelled simultaneously page 51

"Filtrate area of 10m3." should be changed in: Filtrate area of 10m2.

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Design of a (Bio)process to remove nitrogen oxides from flue gases CPD 3332

Summary

Nitric oxide (NO) and nitrogen dioxide (N02) are two greenhouse gases that contribute significantly to air pollution. These gases, often denoted as NOx gases, are side products of, for example, combustion of natural gas. In genera I industrial flue gas emissions are responsible for 17% of NOxjSOx released in the atmosphere. Approximately 24 million tons of NOx were released to the atmosphere from US sources during 1998. So the removal of NOx in particular is of great importance for creating an environmentally less damaging industry. Several processes are developed to remove NOx from (flue) gases. Two of the conventional post combustion techniques are Selective Catalytic Reduction (SCR) and Selective Non Catalytic Reduction (SNCR). In order to increase the sustainability of the flue gas scrubbing processes, bioprocesses can be introduced. The BioDeNOx process combines absorption of

NO

In an aqueous Fe(II)EDTA2- solution with biological reduction of the

absorbed NO in a bioreactor. The usage of Fe(II)EDTA2- can be explained by the formation of a stabie complex between the Fe(I1)EDTA2-and the NO. The biologica I reduction of NO to N2

gas takes place under thermophillic conditions. Ethanol is used as electron donor.

The process will work best with two reactor sections: An absorbtion section and a bioreactor section. For the absorption section a spray tower is selected due the low liquid fraction in the column. This is required due to large Hatta number for complexation and small Hatta number for oxidation. Which will result in a large selectivity for the absorption of NO. For the bioreactor a continuous stirred tank bioreactor (CSTB) is chosen. The main reason for choosing a CSTB is the ease in operation and the optimal growth of biomass in this type of reactor.

Models are developed taking into account the kinetics of all reactions and mass transfer resistances in both the phases. First models are set up describing the film layer. In this way the behavior of all relevant components are known. Conclusions can be made on whether reaction is taking place or just pure diffusion. Taking into account these conclusions, new models are derived. Analytical expressions are derived for NO and flux expressions for other components are linked to the NO expressions. In this way, less computational power is required to perform all calculations. Decreasing the computational power was necessary due to the lack of sufficient internal memory in the available Personal Computers. With these new flux expressions, new column equations are derived. These column equations are used to generate results. The main goal of the process is to reduce the NO concentration of flue gas with concentration of

100-300

o.arts per million (ppm) to

10

ppm. The process is simulated in matlab, where a modells written for the spray tower and the bioreactor. Spray tower models are set up for both the film layer and for the column (x and z direction). For the bioreactor simple molar balances are set up. This since up till now no detailed information is known about the 'chemica I behavior' of the biomass. With this model important parameters of the process can be calculated. All models are implemented with Damkohier numbers. With these numbers the reactor volume can be determined, if kinetic constants are known. With a FeEDTA concentration set on 39,9 moljm3 and a volume flow entering the absorber set on

0.030

m

3

_jsec.

_{The model meets the requirements.}

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Design of a (Bio )process to re move nitrogen oxides from flue gases CPD 3332

Table of content

1 INTRODUCTION ... 1

1.1 ENVIRONMENTAL BACKGROUND OF NOx ... 1

1.2 OTHER NOx REMOVING TECHNIQUES ... 1

1.3 BIOTECHNOLOGY FOR NOx REMOVAL ... 1

1.4 THE FE(Il)-EDT A COMPLEX ... 2

1.5 PROBLEM DEFINITION ... 3

1.6 AGREEMENTS WITH PROBLEM OWNER ... 3

2 PROCESS OPTIONS AND SELECTION ... 4

2.1 PHASE ANAL YSIS ... 4

2.2 ABSORPTION SECTION ... 6

2.2.1 Batch versus continuous ... 6

2.2.2 Hatta calculation ... ... 6

2.2.3 Absorber selection ...... 9

2.2.4 Conclusions ... ... 10

2.3 BIOREACTOR SECTION ... 11

2.4 BATCH OR CONTINUOUS OPERATION ... 11

2.4.1 batch process ... 11

2.4.2 Fed-batch operation ... 11

2.4.3 Continuo us operation ... 12

2.4.4 Applicationfor the CPD process ... 12

2.5 REACTOR TyPE ... 12

2.5.1 Continuous stirred tank bioreactors (CSTB) ... 12

2.5.2 Packed bed ... 13

2.5.3 Liquid fluidized-bed reactor ... 14

2.5.4 Airlijt. ... 15

2.5.5 Applicationfor the cpd project ... 16

2.6 BIOMASS ... 17

2.6.1 Denitrification ...... 17

2.6.2 Reduction of F e( 111 )EDTA· ... 18

2.6.3 Addition of sulfur compounds ... 18

2.6.4 Electron donors ... 19

2.6.5 Conclusion ...... 19

3 ~WS OF DESTC;N ... 20

3.1 DESCRIPTION OFTHE DESIGN ... 20

3.2 PROCESS DEFINITION ... 20

3.2.1 Process concept chosen .......... 20

3.2.2 Block scheme ... 21

3.2.3 Thermodynamic properties ...... 22

3.2.4 List of pure component properties and safety data ... 22

3.2.5 Feed composition ... 22 3.2.6 Purity of ethanol ... ...... 23 3.3 BASIC ASSUMPTIONS ... 23 3.3.1 Plant Capacity ... 23 3.3.2 Location ........... 23 3.3.3 Battery limit ... 23

3.3.4 Substances passing the battery limits ...... 24

3.3.5 Margin ...... 24

4 THERMODYNAMIC PROPERTIES AND REACTION KINETICS ... 25

4.1 GIBBSENTHALPY ... 25

4.2 ENTHALPIES AND HEAT CAPACITIES ... 25

4.3 PHYSICAL AND KINETIC PROPERTIES ... 26

4.3.1 Solubility Oxygen ... 27

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4.3.2 4.3.3 4.3.4 4.3.5 4.3.6 4.3.7 4.3.8 4.3.9 4.3.10 Solubility NO ... 28 Diffusivity NO ...... 29 Diffusivity F e( 11)EDTA 2-...•...•.•...•...•...••••..•...•.•...•..•.•••••...••••... ..•. 30 Diffusiviiy Oxygen ...... 31 Kinetics ... .... 31 kinetic constant NO ... 32 Equilibrium constant NO ...... 32

Kinetic constant Oxygen ... 33

Summary table ... 34

S PROCESS STRUCTURE AND DESCRIPTION ... 3S 5.1 CRITERIA AND SELECTION ... 35

5.1.1 Heat exchangerISteam generator (EOl) ... 35

5.1.2 Flue gasfeed heat exchanger (E02) ... 35

5.1.3 Heat exchanger (E03) ... 35

5.1.4 Absorber(COl) .......... 36

5.1.5 Rotaryfilter (SOl) ............. ..... 36

5.1.6 Bioreactor (Ral) ... 36

5.1.7 Cartridge filter (S02/S03) ... 36

5.1.8 Chimney (X03) ... 37

5.2 PROCESS FLow SCHEME (PFS) ... 37

5.3 PROCESS STREAM SUMMARY ... 38

5.4 UTlLlTlES ... 38

5.5 PROCESS YlELDS ... 38

6 PROCESS CONTROL ... 40

6.1 CONTROL OBJECTIVES ... 40

6.2 CONTROLSTRUCTURE ... 40

6.2.1 Heat exchanger (EOI) ................ 40

6.2.2 Heat exchanger (E02) ........... 41

6.2.3 Heat exchanger (E03 ) ... 41

6.2.4 Absorber (Cal) ......... 41

6.2.5 Bioreactor (Ral) ... 41

6.2.6 Rotary filter (SOl) andfilter (S02/S03) ... 42

6.2.7 Clean gas chimney (X03 ) ...... 42

7 MASS AND HEAT BALANCES ... 43

7.1 BALANCEFOR TOTALSTREAMS ... 43

7.2 BALANCE FOR STREAM COMPONENTS ... 43

7 .3 EXPLANATION AND ASSUMPTIONS OF THE MASS BALANCE AND HEAT BALANCE ... 43

7 .4 EXPLANATION CALCULATION OF THE STREAMS IN THE ABSORBER ... .44

7 .5 EXPLANATION CALCULATION OF THE STREAMS IN THE BIOREACTOR ... .44

7 .6 EXPLANATION CALCULATION OF THE STREAMS IN THE HEAT EXCHANGERS ... .44

7.6.1 Heat exchanger EOI ... 44

7.6.2 Heat exchanger E02 ... 45

7.6.3 Heat exchanger E03 ... 45

7.7 PUMPS ... 45

7.8 HEATLOSS IN ABSORBER ... 45

8 PROCESS AND EQUIPMENT DESIGN •..•.•..•.••..•.•...•...••.•••••••••... 47

8.1 PROCESS AND EQUIPMENT DESIGN ... 47

8.1.1 Absorber ............... 47

8.1.2 Bioreactor ... 47

8.1.3 Heatexchangers ... 48

8.1.4 Rotary filter SOl ................. ... 51

8.1.5 Cartridge filter S02/S03 ... 51

8.1.6 Chimney (X03) ......... ... 51

8.2 INTEGRATION BY PROCESS SIMULATION: THE MATLAB MODEL. ... 53

8.2.1 Summary model development ... 53

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8.2.2 Basic assumptions ............... ...... 53

8.2.3 Dimensional input ............................ 53

8.2.4 General model derivation Film layer ... 54

8.2.5 Film model derivation NO ... 55

8.2.6 Molar balances of other liquid phase components ... 57

8.2.7 Film model results ... 58

8.3 GENERAL MODEL DERN A TION COLUMN ... 63

8.3.1 NO molar balance gas phase column ......... 64

8.3.2 Oxygen molar balance gas phase column ... : ... 65

8.3.3 General molar balance liquid phase column ................. 65

8.3.4 NO molar balance liquid phase column ... 65

8.3.5 Oxygen molar balance liquid phase column ............ ..... 67

8.3.6 Fe(II)EDTA2. molar balance liquid phase column ... 67

8.3.7 F e( II)EDTA _N02• molar balance liquid phase column ... ... 67

8.3.8 Fe(III)EDTA" molar balance liquid phase column ... 67

8.4 MODEL IMPLICATIONS / CALCULATIONS ... 68

8.4.1 Film model simplifications ... 68

8.5 COLUMN SIMPLIFICATIONS ... 71

8.5.1 NO gas phase column equation ... ... 71

8.5.2 Oxygen gas phase column equation ... ... 71

8.5.3 NO liquid phase column equation ... 72

8.5.4 Oxygen liquid phase column equation ......... ... 72

8.5.5 Iron(II)EDTA liquid phase column equation .................................. ...... 72

8.5.6 Iron(II)EDTA-NO liquid phase column equation ............... 73

8.5.7 Iron(III)EDTA liquid phase column equation ................. ... 73

8.5.8 Modeling Bioreactor ... 73

8.5.9 Overall molar balance ...... 73

8.5.10 Molar balance Iron(ll)chelate ............ 74

8.5.11 Molar balance Iron(lI)chelate-NO ... 74

8.5.12 Molar balance iron(lI)chelate ... 74

8.6 MODELING RESULTS ... 74

8.7 EQUIPMENTDATASHEETS ... 82

9 WASTES ... 88

9.1 GASEOUS WASTE STREAMS FROM THE ABSORBER AND THE BIOREACTOR ... 88

9.2 THE LIQUID BLEED ... 88

9.3 THE SOLID BLEED ... 89

10 PROCESS SAFETY ... 90

10.1 HAZOP ... 90

10.2 Dow F&EI (FIRE AND EXPLOSION INDEX) ... 90

10.3 PROCESS UNIT RISK ANALYSIS ... 92

10.4 CONCLUSION ... 93

11 ECONOMICS ... 94

11.1 INVESTMENT (ONCE-OFF) ... 94

11.2 QpERATING COSTS (ANNUAL) ... 95

11.3 INCOME ... 95

11.4 CASH~OW ... 96

11.5 ECONOMIC CRITERIA ... 96

11.6 COST REVIEW ... 96

12 CREATIVITY AND GROUP PROCESS TOOLS ... 98

13 CONCLUSIONS AND RECOMMENDA TIONS ... 100

13.1 CONCLUSIONS ... 100

13.2 THE STRENGTHS OF THE DESIGN ARE: ... 102

13.3 THE WEAKNESS OF THE DESIGN ARE: ... 102

13.4 RECOMMENDATIONS ... 102

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13.4.1 13.4.2 13.4.3 13.4.4 13.4.5 13.4.6 13.4.7

1ncreasing the Fe(ll)EDTA in the system . ......... 102

Decreasing temperature in the column ...... ...... 102

Changing volumetrie flow rate ...... ...... 102

Sensitivity analysis ............................ 103

Absorber sizing ........ . 103

Parameters / dimensional values ... 103

Bioreactor ................. ......... 103

LIST OF SYMBOLS ... 104

REFERENCES ... 106

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1 Introduction

1.1 Environmental background of NOx

Nitric oxide (NO) and nitrogen dioxide (N02) are two greenhouse gases that contribute significantly to air pollution. These gases, often denoted as NOx gases, are side products of, for example, combustion of natural gas. In general industrial flue gas emissions are responsible for 17% of NOxjSOx released in the atmosphere.1 Approximately 24 million tons of NOx were released to the atmosphere from US sources during 1998.2

₅₀

_{the removal of}

NOx in particular is of great importance for creating an environmentally less damaging industry. Several programs to reduce the NOx emission into the air are implemented. Titles 'I and IV of the 1990 Clean Air Act Amendments' reg u late NOx emissions from major sources. The goal of these programs was to reduce the NOx production with 2 million tons per year

below 1980 levels by the year 2000.3

1.2 Other NOx removing techniques

Several processes are developed to remove NOx from (flue) gases. Two of the conventional post combustion techniques are Selective Catalytic Reduction (SCR) and Selective Non Catalytic Reduction (SNCR). Major drawbacks of these techniques are that they produce secondary waste streams and it operates under extreme conditions.3 Another removal technique is the absorption of NOx by using an iron-complex. This can be done by means of different absorption columns and Gas Liquid contactors like packed columns and tray columns. As absorption liquids water, hydrogen peroxide and sodium hydroxide can be used.4_,_s

1.3 Biotechnology for NOx removal

In order to increase the sustainability of the flue gas scrubbing processes, bioprocesses can be introduced. The BioDeNOx process combines absorption of NO in an aqueous Fe(II)EDTA2

-solution with biological reduction of the absorbed NO in a bioreactor. The usage of Fe(II)EDTA2- can be explained by the formation of a stabie complex between the

Fe(II)EDTA2

- and the NO. The biological reduction of NO to N₂ gas takes place under

thermophillic conditions. Ethanol is used as electron donor. Regeneration of Fe(II)EDTA2 -,

which is oxidized to Fe(III)EDTA-, is simultaneously do ne in the bioreactor.6

The absorption of NO is performed by means of wet absorption. The absorption takes place according to:2

Fe(II) EDTA2

-

+

NO Fe(II)EDTA - N02

-The absorption liquid is denitrified in the reactor and the Fe(II)EDTA2- is regenerated using an

electron donor. If ethanol is used as an electron donor, the reaction becomes:2

This de-nitrification process is schematically represented in Figure 1.1

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11 BlOLOGICAL

Fe (L) "",,_ REGfHERATION

" \ /

);N

,

Figure 1.1 Schematic representatian af denitrificatian princip/t!.

Since oxygen is present in the f1ue gas, it is therefore also in contact with the absorption

liquid, the oxygen will react with the iron. The oxidation of Fe(II)EDTA2- is as follows6:

This reaction is undesired because the Fe(III)EDTA- complex is not capable of absorbing any

NO. Regenerating of Fe(II)EDTA2-in the bioreactor is as follows6:

1.4 The Fe(II)-EDTA complex

Ethylenediaminetetraacetic acid or EDTA is a synthetic amino acid first used in the 1940'5 for

treatment of heavy metal poisoning. It is used extensively in a wide variety of industries (e.g.

metal, rubber, leather, textile, pulp and paper). The annual usage of EDTA in Western Europe from 1988 to 1994 was between 26000 and 30000 tonnes. Approximately 1200-1500 tonnes

of EDTA were sold in the Netherlands in the same period7,8. The EDTA molecule contains four

carboxylic groups. Further on, the molecule is highly labiIe (has a fast coordination rate), and forms a strong bond with divalent and trivalent metals using six coordination sites. This is

schematically drawn in Figure 1.29

o

0 • 11 11 • HO-C-CH2~ • . / CH2- C- OH /N-CH2CH2-N~" HO- C-CH2 CH2- C-OH • I1 11 •

o

0 Figure 1.2 EDTA

EDTA has chelating capacities, which means it can bind certain positively charged metals like Fe(II) or Fe(III). Figure 1.3 illustrates the six points at which the EDTA molecule coordinates with the

chelated metal. In this state, the Fe(II) chelated

ethylenediaminetetraacetic acid can rapidly react with absorbed NO gas to form a sta bie metal-nitrosyl complex. Denitrifying bacteria have shown to be able to reduce the ferric chelate as weil as the

nitrosyl complex back to the ferrous form. 10,7,9

Agure 1.3

Conceptual Process Design: Removal of NOx from flue gases

Che/ated EDTA

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Design of a (Bio )process to remove nitrogen oxides from flue gases CPD 3332 1.S Problem definition

The assignment "DESIGN OF A (BlO) PROCESS TO REMOVE NITROGEN FROM FLUE GAS ES" is intended to assess the feasibility and economics of a process for NO removal from flue gases. The main goal of the process is to reduce the NO concentration in the flue gas of 100-300 parts per million (ppm) to 10 ppm. The process uses Fe(II)EDTA2- as binding agent \

followed by mierobial denitrifieation. The turndown ratio of 50% must be feasible. The rate change of the feed flow will not exceed 10%/hr.

---~

The following eonstrains are known. (nominal values in bold)

• Flow rate 20000-40000 Nm3/hr • Temperature

400

°C (± 25°C) • Pressure 1.1 bar • Composition NOx

100-

300 vppm CO2 13 v/v% O2

2.0 -

5 v/v% • Particulates 5 mg/m3

General eonstrains of ecologieal, economie and social kind are as followes. No major changes in the local eeosystem are allowed. No toxie emissions in to air, water and land at the location. The price of product / produced energy must be eompetitive relative to existing technologies. (see chapter 1.2) The proeess has to be aeeepted by Duteh society and local communities.

1.6 Agreements with problem owner • The NOx can be seen as 100% NO.

• No sulfur oxides are present in the flue gas.

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2 Process options and selection

2.1 Phase analysis

Before choices can be made about the process equipment selection, it is important to know which phase transitions take place in the process. Wh en looking at the flue gas entering the process, the NO is in the gas phase. Therefore to remove the NO a GIG-interface, a

G/L-interface or a GIS-G/L-interface must be created. A salt like the Iron EDTA chelate complex does not exist in the gas-phase, therefore a GIG-interface is not possible and the two other options

(G/L-

or GIS-interface) remain.

The way to form a nitrosyl complex between NO and Fe(II)EDTA2- is given in the following

reactionl l:

NO

+

Fe(II)EDTA2- . . . Fe(II)EDTA-N02

-This nitrosyl complex can be formed in a system with gases and liquids, where the gaseous

~

c."

~\

NO is dissolved in an aqueous solution of Fe(II)EDTA2-. The aqueous NO will react with the

~

0-

ionic Fe(II)EDTA2 and forms the nitrosyl complex.

\) \

~

For a system with solid Fe(II)EDTA2-and gaseous NO it is not clear represented in literature if

_~ ""'\ it can form a nitrosyl-complex. Therefore in this state of the design, the two options

~

~~

\J' mentioned above will be taken into account.

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y

\.

D The conversion of nitrous oxide bounded to the Fe(II)EDTA2-to N2 is catalyzed by denitrifying

~ 'f- bacteria according the following reaction:

~~

_J

_~

6Fe(II)EDTA-N02-

+

electron donor

To catalyze the reduction of NO to N2, the denitrifying bacteria use an organic electron donor

(in this case ethanol). Since the microbial communities are living organisms, they need an aqueous environment for prevention of dehydration, supply of nutrients and removing of the waste material produced by the organisms. The denitrification will take place inside the cells of the bacteria. Therefore the Fe(II)EDTA-N02- has to make contact with the cell membrane

to make absorption in the cell possible. Since the cells live in an aqueous solution, the only way to make contact between the cell membrane and the Fe(II)EDTA-N02

- is to dissolve the

Fe(II)EDTA-N02- in the same solution where the bacteria lives.

As indicated above the denitrification part will take place in the liquid phase. Therefore two options for the process are recognized:

1. Removal of NO from the f1ue gas with a

GIL

-interface by using an aqueous solution of Fe(II)EDTA2-. And use a liquid phase for the denitrification of the nitrosyl-complex.

2. Removal of NO from the f1ue gas with a GIS-interface. To get the denitrifying bacteria in contact with the 'solid nitrosyl complex' it must be dissolved in water.

For the first option it can be remarked that the whole process can be executed in the liquid phase. The only phase transition that takes place is the transition from the gas phase to the liquid phase (dissolving NO). This instead of the second possibility where an extra operation step must be added; namely the dissolving of the nitrosyl-complex. Even 50, the process needs an additional step to regain solid Fe(II)EDTA2

- and to put it on a surface in the

absorber again. When the advantages and disadvantages of both options are compared, it can be concluded that the best way to remove NO from the f1ue gas is by creating a gas/liquid interface. Subsequently the denitrifying bacteria can convert the nitrosyl complex.

The process consists of a chemical part and a biochemical part, the first contains the reactive absorption of NO in an aqueous FeII(EDTA) solution, the latter is the biological regeneration

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Design of a (Bio)process to remove nitrogen oxides from flue gases CPD 3332 of the iron chelate solution using denitrifying bacteria. Hence there are two options: The first option is to integrate both reactions in one system; the second option is to divide the process in two reaction sections. A disadvantage of integrating the process in one system is that both processes must be operated at similar conditions. This is not preferabie because the tem peratu re, for instance, has a profound effect on both the absorption rates and the selectivity. To promote NO absorption, it is advantageous to work at low temperatures. To get optimal conversion rates of the microorganisms, it is needed to work at optimal

conditions, which means that specified temperatures are also required. It is not very likely

th at the low temperatures in the absorption section correspond with the specified temperatures in the bioreactor.

Further on, some other parameters are important to obtain an optimal environment for the

microorganisms. The pH and the redox potential are of importance for the conversion rate of " -.

the microorganisms. When working with an integrated system it is not possible to optimize

0

one reaction without influencing the conditions for the other reaction. Therefore it is impossible to operate both reactions at optimum conditions. So when looking at these different parameters and the flexibility it is advantageous to work in a system where the absorber and bioreactor are separated.

An integrated system is more preferabie when looking at the spatial concerns. When using one reactor, it will be larger then a bioreactor of a separated process because a large part of the reactor will be taken by gas. Still there is no separated absorber, the total required space of an integrated process is a lot smaller compared with the divided process. Another advantage of an integrated process is that less (or non) circulation streams are needed. Looking at the required space and equipment (e.g. pumps and pipelines) an integrated process is preferred above a separated system.

Both systems have their advantages and disadvantages. The choice between both systems is based on the importance of every aspect. Therefore looking at the requirements of the process, the most important aspect is flexibility. The flue gas does not enter the process at

constant conditions (flow and composition) and if one of the both reaction steps does not

work optimal it must be possible to change one parameter without influencing the whole system. Because the process is new, at least at plant scale, the last point is of extra importance. The complications of the process at plant scale are not known and therefore the availability of changing and optimizing the process has to be as large as possible.

When looking at the discussion above, it can be concluded th at a divided process, with separated bioreactor and absorber, is the best choice. The process as it can be explained at this stage of the design is schematically drawn in Figure 2.1.

....

I

"

I

Absorption of NO Biochemical

with Fe(II)EDTA de-n itr ification

I

..

,..

I

Rgure2.1 simpte representatian af the process

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2.2 Absorption section

The absorption of NO into solution is a significant step in the BioDeNOx process due to fact that the biomass needs NO in solution in order to convert to N2 gas. NO enters the system in the gas phase and therefore needs to transfer into solution. The physical absorption of NO in water is not feasible for this particular application due to the low solubility of NO in water. Enhancing the absorption by using chemical reaction can be proposed in order to get sufficient cleaning of the flue gas. The 'chemical' flux of a component can be expressed as the product of the 'physical' flux times an enhancement factor due to chemical reaction (see

Eq.2.1)

Ja(Chemical) = Ja (physical)

*

Ea

Eq.2.1 The reactants (in this case NO and O2) still have to diffuse into the liquid where the reaction takes place simultaneously in series or parallel. For the reactor (absorber) selection it is of great importance to determine the rate-determining step.

2.2.1 Batch versus continuous

Since the volumetric flow rate of the gas feed is significant large continuous stream

(40000

Nm3/hr) a continuous system has the advantage over a batch system.

2.2.2

Hatta calculation

The type of absorber that is most suitable as gas-liquid contactor

I

reactor for the BioDeNOx

t\ _~ \ 1\ process depends on several parameters. One of them is the location of where the reaction

?)Y'

\DI

'takes place. By applying the film theory one can distinguish a film layer and a bulk in both

/ \.'( phases. Both film layers have ma ss transfer resistances but no convection, so all mass

\. r,<

y...1-

transfer is based on diffusion and reaction. Mass transfer diffusivity follow Fick's law.

~

Assumed is that the iron chelate complexes do not evaporate into the gas phase and

_ \ (Jc, ~ • therefore no reaction takes place in the gas phase. No mass transfer is taking place from the

""

\~.j1' liquid phase to the gas phase.

\J

î

~

In which reg ion the reaction takes place depends on the rate of mass transfer versus the rate

of reaction. Therefore the Hatta number is introduced. The value of Hatta is an indication of

)-'\ the location of reaction. Since both gases absorb simultaneously, competing in reaction with a

component (iron chelate) in the liquid, the reaction can be represented as followed:

NOG ~ 02G ~ NOL

+

02L

+

NOL 02L Fe(II)EDTA2- Fe(II)EDTA-N02- Fe(III)EDTA-Wanted Unwanted Introducing the Hatta number for a n,m-order kinetics reaction: 12

Ha

=

J.-

(_2_k

Den-IC

m)

k

L

n+l

n,m A A,I B

= mass transfer coefficient

= diffusion coefficient

= concentration gas phase component

=

concentration liquid phase component

(mis) (m2/s)

(moljm3)

(mol/m3)

Component A is the gas phase component and B denotes the liquid phase component.

Eq.2.2

With large mass transfer resistance relative to the reaction, the Hatta number is large. On the contrary with small mass transfer resistance and slow reaction the Hatta number is small. With large Hatta numbers all reactions take place in the film of the liquid. With small Hatta numbers the reactions take place in the bulk of the liquid.

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Reflecting to absorber designjselection, with large Hatta numbers (Ha >2) the most important factor is the interfacial area of the absorber. The liquid hold-up should be as small as possible because of the fact that it is functionless. With small Hatta numbers (Ha <0.2) the interfacial area is of no importance because all reaction takes place in the bulk of the reaction. Increasing the bulk volume increases the conversion.

Hatta number tor comp/exation

For a second order reaction like the reaction between Fe(II)EDTA2-and NO the Hatta number

becomes (n=l, m=l)

Ha

=f-~(kl

.

IDACB)

L

Eq.2.3

Where: k1,1 = kinetic constant

For the complexation reaction the Hatta number can be calculated. The diffusion coefficient of NO in Fe(II)EDTA2

- solution, the mass transfer coefficient of NO and the kinetic constants

are given in Table 2.1.

Tab/e2.1 Parameters far Hatta ca/cu/atian.

(T =

±

323K) r î Reference:

Diffusion Coefficient

8.610-*

9 _(m2_js_)~ _{F. Gambardella [10]}

Mass Transfer Coefficient ~ 3.9*10-4 (mjs i~ _ F. Gambardella [10] Kinetic constant

f

+:A..510*

5

(m3jmol.s)Q..J F. Gambardella et.al. [13]

~

The Hatta number corresponding to the above constants, and a Fe(II)EDTA2- solution of 50

moljm3, is significantly larger then two. Therefore the reaction is fast compared to mass transfer and takes place in the film layer of the liquid. Hatta as a function of iron chelate concentration is plotted below .

.923.53~ 000 920 840 760 680 Ha(Cb) 600 520 440 360 280 ..292.046.₂₀₀L - - - - 1 _ - - - L _ - L _ - ' - - _ . . l - - _ l . . . - - - - 1 _ - - - L _ - L - - - - ' o 10 20 30 40 50 60

J

O

80 90 100 .10. Cb

l

\MM

S

}

~

.100.

Rgure2.2 Hatta as function of iron che/ate concentration (mali"";)

Hatta number tor oxidation

The main purpose for the absorber is the uptake of NO into the solution. Secondly the adsorption of oxygen should be minimized because it can oxidize Fe(II)EDTA2- into

Fe(III)EDTA- which is incapable of absorbing NO as result the efficiency of the column decreases. The best option for selecting a column is high NO adsorption selectivity with a small Hatta number for oxygen.

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~ - - - --- - --- ---.

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TU Delft

It is stated in several literature sources9,14 that the reaction kinetics between the EDTA complex and oxygen is more complex compared to the reaction between NO and the same complex.

The reaction order for oxygen is one, but for the process conditions range in which the BioDeNOx is proposed to work in (CFe(II)EDTA2- = 15-60 mOI/m3, pH=5-8, T=298-328K), the

reaction order of the Fe(II)EDTA2- is approximately two. The Hatta number for the oxygen reaction becomes:

Ha

=

~

'k

₁₂

D

_A

C/

KV'

L

Where: Kl,2 = kinetic constant

In which A is Oxygen and B is the Fe(II)EDTA2- complex. The data required for the Hatta calculation is as follows.

Tab/e2.2 Data of axygen and

Fe(II)EDTk-CT=

±

323K) Diffusion Coefficient CDA) 4.1 *10-9 Cm2/s)1O Mass Transfer Coefficient (KL) 1.8*10-4 (m/s)lO Kinetic constant (Kl 21

+

0.01 (m6/moI2.s)14

References

F. Gambardella [10] F. Gambardella [10] Wubs et.al. [14] The Hatta number in a concentration range of 0 - 100 is plotted in Figure 2.3 .

• 4~283. 5

Ha(Cb)

.0.045. 0 ' - " - - - " ' - - - ' - - - " ' - - - ' - - - '

o m ~ ffi m ~

~ ~ A

Figure 2.3 Hatta number far Oxygen

Conceptual Process Design: Removal of NOx from flue gases

Eq.2.4

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2.2.3

Absorber selection

Looking at the possible gas liquid contactors the following typical values are known.15

Table2.3 Typical values tor different reactors

Reactors a e l kl

[m2jm3]

L-]

[mjs]

Spray Tower 60 0.05 1.00E-04

Packed Bed 100 0.08 5.00E-05

Plate Tower 150 0.15 2.00E-04

Bubble column 200 0.9 2.00E-04

Statie Mixer 200 0.2-0.8 2.00E-04

Bubble Tank 20 0.98 2.00E-04

Agitated Tank 200 0.90 2.00E-04

Since the liquid hold-up has no function in the reactor (Ha»2), the liquid fraction (e l)

should be as small as possible.

Another consideration is the large gas throughput. A rule of thumb for packed bed absorber

states that the flow ratio between liquid and gas should be around ten. 15 This ratio is given

in equation 2.5.

)

P =lbar

Eq.2.5

When the flow ratio gets below 10, flooding can occur. The gas velocity prevents the liquid from falling down. When the flow ratio is above 10, the liquid velocity drops too fa st, which limits the contact time between gas and liquid.

The packed column is a counter-current flow device with liquid going down from top to Wffom and gasJvapor phase ascending from botlom to top. Columns are always designed to allow maximum gasjvapor flow capacity. In general the gas velocity is 20% below the maximum allowable value. The maximum allowable velocity is known as the flooding velocity. The flooding velocity is the ratio of the maximum volume flow rate of gas and the cross sectional area of the column.

Eq.2.6

Packed beds work best at these conditions.

A design tree for the absorber selection is shown in the figure below, which summarizes the first decision-making on the absorber.

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Design of a (Bio)process to remove nitrogen oxides from f1ue gases

Spray

tówer

Paeked

bed

Sta:tió

mIXer,

Spray

tower

Rgure2.4

Statie

mIXer

Paeked

,

bed

Plate

tower

Statie

mlXer

tower

Design tree for absorber se/edion

2.2.4 Conclusions

Bubble

tower

CPD 3332

Plàte

,

tow~r

After some simple calculations and following some basic rules of thumb the following

conclusion can be made. Since both main components, that effect the feasibility of the

absorption step of the BioDeNOx process, require a small liquid bulk volume. The following three options are possible: spray tower, packed bed or static mixer. During the design phase a definite selection between the three options will be made. The pros and cons are given in order of importanee in Table 2.4.

Tab/e 2.4 Pros and cons of severa/ options.

Spray Tower Plate tower Static mixer

1) Low liquid hold up

++

+

+/-2) Interfacial area

+/-

+

++

3) Feed ratio

++

4) Capacity

-

+

++

5) Low pressure drop16

++

+/-

-++

=

Excellent

+

=

Good

+/-

=

Medium

=

Bad

=

Terrible

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TU Delft

2.3 Bioreactor section

In this conceptual process design, microorganisms are involved to catalyze the de-nitrification of NO to N2 and the regeneration of the chelating agent Fe(II)EDTA2-. To obtain optimal

conditions and therefore optimal reaction rates for both reactions, the choice of the reactor is very important. The main function of a properly designed bioreactor is to provide a controlled environment in order to achieve the optimal growth andjor product formation. The performance of any bioreactor depends on many functions, including:

• Biomass concentration, which must remain high • Constant temperature

• Optimal pH and optimal temperature

• Optimal mixing or creation of the correct shears conditions. High shear rates may be harmful to the organism, but low shear rates mayalso be undesirable, because of unwanted f1occulation or growth of biomass on the reactor wall and stirrer17

The points above have to be taken into account when making decisions on the type of the bioreactor. But for a good decision-making, the characteristics of the processjreactions must also be taken into account. Some important aspects are listed below:

• The phases that will be present in the bioreactor are: (1) a solid phase of particulates and biomass, (2) a liquid phase of water and ethanol, with the reacting agents and nutrition for the biomass and (3) a gas phase with the products of the reaction (N2, and

CO2).

• The absorber will generate a continuous stream of water with the nitrosyl-complex that has to be treated in the bioreactor. And the absorber requires a continuous stream of Fe(II)EDTA 2-.

• Since oxygen will oxidize the Fe(II)EDTA2- complex, the oxygen level inside the

. bioreactor must be as low as possible

• There must be a possibility to remove solids from the system, otherwise there will be accumulation of solids

• A turndown ratio of 50% must be available. Therefore the reactor must be adjustable in capacity.

2.4 Batch or continuous operation

The points above will be considered further in this chapter, but first a choice will be made which operation mode will be used:

1. Discontinuous operation (batch process) 2. Semi-continuous operation (fed-batch) 3. Continuous operation18

2.4.1

batch process

Because there are no incoming and out going f1ows, all the substrates are in the medium from the beginning. Therefore, their initial concentrations are quite high. After inoculation of the biomass, the cells are growing uncontrolled until essential nutrition's are exhausted or the accumulation of inhibiting products ceases the growth. The advantage of batch operation is the low effort for process control. This because it can be applied advantageously when high substrate concentrations have no negative effect on the desired biologica

I

reactions. Batch operation is also preferred when producing various products and when handling small amounts.19

,18

2.4.2

Fed-batch operation

When operating in a fed-batch mode, medium components are continuously fed to the reactor, but no medium is taken out. The fed-batch mode of operation is mostly used when the substrate concentration must be kept low for optimum growth or product formation. Both, batch and fed-batch operation modes are preferred to avoid problems with strain stability and sterility that may arise during prolonged cultivations. The disadvantage of the

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TU Delft

previous discontinuous modes of operation is the low productivity, mainly by two reasons: after each reaction section (cultivation), the reactor has to be emptied, cleaned and refilled again. This causes long unproductive intervals. Furthermore, since the reactor is inoculated with low cell densities, the maximum cell concentration and production rate is only reached close to the end of the cultivation. lS,19

2.4.3

Continuous operation

In a continuous cultivation there is a permanent inflow of substrate and outflow of medium. The cell concentration can be maintained at high values comparable to the maximum values in batch operation. Similar to the fed-batch mode, continuous operation provides good possibilities to control the biological reaction by setting a proper residence time via the flow rate. The advantage of a continuous process is that it usually can reach the highest productivity, since the repeated cycle of reactor propagation is not required. It will have a large scope for mechanization and automation. On the other hand, failures in the equipment, infection by other microorganisms, or aging effects of the cell population limit the maximum operation time.1S,19

2.4.4

Application tor the CPD process

A batch or fed-batch operation process is mostly preferred when working with unstable/non steriIe microorganisms. The microorganisms used in this conceptual process design are taken from, for instance, water treatment plants. In these plants the circumstances are not steriIe and the microorganisms will not be renewed very often or not at all.

A batch process is only favorable when high initial concentrations can be obtained. In this case the concentration of the nitrosyl complex is not very high. Therefore a batch operation would not be the best option.

The absorber will deliver and require a continuous liquid stream. Therefore the bioreactor needs at least a production rate equal to the absorption rate. Batch or fed-batch would be possible, but than more reactors or large storage vessels are needed. It is not sure if this will contribute to a cost effective design.

50

when looking at the advantages and disadvantages of the three operation modes, it can be concluded th at a continuous operation mode would be most suitable for the CPO design.

2.5 Reactor type

To obtain optimum reaction conditions it is important to choose a specific type of reactor. Below four typical bioreactors will be describes and subsequently a conclusion will be made which reactor is best for the de-nitrification process.

2.5.1

Continuous stirred tank bioreactors (eSTS)

The most straightforward continuous bioreactor is the continuous stirred tank bioreactor. The catalyst is suspended in a large tank through which substrate flows, and is retained within the reactor by filtration, subsequent sedimentation, magnetic forces or by being attached to the stirrer paddies. The CSTB has the dual advantages of low capital cost and low operating costs. The height-to-diameter ratio of the vessel can vary between 2: 1 and 6: 1, depending largelyon the amount of heat to be removed, and the stirrer may be top or bottom driven. A CSTB is fitted with baffles, which prevent a large central vortex being formed and improve mixing. Usually vessels with a diameter less than 3 mare fitted with four baffles. Larger vessels are fitted with six or eight baffles. If heat removal seems to be a problem, which can occur in vessels greater than 100m3, up to twelve baffles can be used, and coolant passed through the baffle. Good mixing can be obtained and thus diffusional limitation can be reduced to a minimum.

ca

refu I consideration has to be given to agitator design within a bioreactor, because it controls the operation of the bioreactor. This because the mixing patterns will determine the conversion rates. Some types of agitators are given in Figure 2.5

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Design of a (Bio)process to remove nitrogen oxides from f1ue gases

fllx-BLADE

DISC TURBINE

Figure2.5

BIX-BLADE 45°

PITCH EO TURBINE

different types of agitators

PADDLE

A[JDITATDR MIG-ADITATOR

CPD 3332

The turbine stirrer, being easy to construct and having a high power number, is the most widely used type of stirrer. The other types are less intensively applied. A eSTB has a typical low void fraction (The ratio of the volume taken up by air spaces (the voids) to the total volume of a material) therefore in comparison with other reactors a bigger volume is needed to have a comparable conversion. In practice it is probably uneconomic to aim for conversion factors of more than 90% with this type of reactorP/O/l A very good mixing can be obtained in a eSTB thanks to different agitators th at can be applied. When a choice is made about the agitator type, it should be taken into account th at the stirrer does not harm the biomass and also good mixing is obtained. It is important for the reaction and biomass to obtain a certain redox potential in the reactor; this will depend on the pH and the ethanol concentration. Therefore easily adding of ethanol and aCid/base must be possible, which is easy when operating the continuous stirred bioreactor, the pH can be influenced. Thanks to good mixing the pH will be the same in the whole reactor. Further on, the reactor can be fitted with a bleed to re move solids that are present in the system. The oxygen level in the reactor shall not differ a lot from other anaerobe reactors. When the flue gas stream is reduced to 50%, the amount of NO in the gas stream will increase. To obtain the required level of NO concentration in the outlet, the capacity of the eSBR must therefore be higher at low flow rates of flue gas. For this type of reactor it will not be a problem to increase the capacity, hence flexible design is possible. Because the reactor is easy accessible, adding of microorganisms will not be aproblem.

2.5.2

Packedbed

A packed bed reactor is a vessel with a fixed bed that contains the catalyst. In a packed bed reactor the catalyst will be immobilized on the bed. The catalyst can be in the form of microorganism or non-microorganism. The particIe size of the catalyst is mostly bigger than normal with a size of 1 and 3 mm22

• Particles can

be

in different shapes like spherical,

cylindrical or other shapes. This reactor can handle gas and liquid flows with counter and co-current flows, and is easily to adapt to a process that is continues. The running costs are very low compared to a high purchase price. Because of the low void fraction there is a smaller volume needed comparable to other reactors. In normal operation the reactor is totally used and the microorganism live in a stabie and wetted environment. In a case of a turn down ration of 50% the stabie environment is not possible22• Turn down ratio of 50% will give

uncontrolled effects on temperature, reactivity, growing of microorganism and pH. The oxygen level will not differ a lot from other anaerobe operated reactors. High conversions are reached easily because of high surface area. Disadvantages are the biocatalyst th at is immobilized. This can block the beds when microorganism is growing. Growing can cause channeling or blocking the total reactor. This reactor is not easily accessible and therefore the reactor has disadvantages like difficult control of the environment in the reacto~o. Especially pH but also temperature control will give problems. Because the mixing takes place by the stream it is not easy to get a homogeneous mixture in the reactor. To avoid this problem the reactor can be made larger or contain a recycle stream. This measurement will give more problems than that is solves. But an advantage of this is th at the dilution rate is very low.

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- - - - -- - -- - -- -- - - - -- - - .

~

TU Delft

Lacking of stirrers in the reactor will prevent the microorganism being damaged. Also the replacement of immobilized catalyst is difficult and takes a lot of time to replace. Because of hindering trough the packed bed of the gas or liquid the reactor builds up a pressure drop. A high pressure drop will ask for more pump capacity to get the liquid trough the packed bed.

2.5.3

Liquid fluidized-bed reactor

Fluidized beds are very common in the chemica I industry. Also in the biotechnology the fluidized beds are lately used more and more. The majority of applications in biofluidization involve the use of a liquid phase to fluidize the solid carrier usually in the form of beads to which a pure enzyme is immobilized. The three most important areas of application for biofluidization are:

1. for enzymes immobilized on a solid matrix

2. for pure cultures of cells immobilized on a solid matrix 3. for various types of effluent treatment

There are various types of bioreactors th at can be used. An overview of the most common types of fluidized beds are given in Figure 2.6:23

Phase combination G-S L-S G-L-S G-L-S @ G

t

L L G , j

I

G

t

5 5 Continual phase G L L G Discontinuous phase S S G.S L,S

Figure2.6 different types of fluidized beds

Since the liquid phase is the continuous phase in the conceptual process design, the first and the latter type cannot be applied. Therefore the second and third type will only be considered here. A solid-liquid fluidized bed is characterized by its high mass transfer/heat transfer rates and uniform distribution of solids. Further on, large surface area is available for biological growth and for contacting with the liquid or gas phases and finally there are economie advantages in the use of biofluidized bioreactors. There are, however, also a number of disadvantages that may make its use undesirable: Back mixing of reactants and products will occur while they are absorbed on the moving catalyst. This results in 1055 of yield and a lowering of conversion. Also the pressure drop of fluidized beds is relatively high in comparison with other forms of equipment. This may make their use uneconomical. And finally, erosion of the vessel walls and internals may occur as a result of the solids motion.

This may make maintenance costs higher than in fixed bed units.

When applying a fluidized bed in the conceptual process design, it will probably not give problems in operating the temperature/heat removal and pH. This mainly because the homogeneity and good mixing properties in a fluidized bed. When looking at the availability for reducing the flue gas stream to 50 % of its capacity, it gives problems in a fluidized bed.

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Design of a (Bio)process to re move nitrogen oxides from flue gases CPD 3332 For instance, the properties of a f1uidized bed are strongly dependant on the flow of the continuous phase. When the velocity of the liquid is increased or decreased, several different

flow regimes can be observed.23,24

2.5.4

Air/ift

The airlift reactor is a modified bubble column reactor characterized by three distinct parts

namely: riser, gas-liquid separator, and downcome~5. The downcomer can be internalor

external looped (Figure

2.7i

6• Both options do not have difference in operation. Both will

have the same pressure drop th at is depended on the height of the reactor. An airlift reactor

is typically used for slow oxidation reactions25. The airlift reactor uses aeration for mixing, fed

into the bortom!7, and has the advantage that there is no mechanical stirring device in the reactor. This means low shear, and therefore minimal changes for harming the microorganisms. Gas that enters the reactor will leave the reactor at the top, and liquid, which is fed on the same place as the aeration gas, liquid can also be taken out at the top of the reactor. Microorganism can stay in the reactor if the flow is good controlled. These kinds of reactors are characterized by low capital costs mainly because of their simple mechanical configuration. Also good mixing is a benefit of the reactor. Considerable backmixing in both

gas and liquid phases, high pressure-drop and bubble coalescence can be disadvantageous27.

The mixing stream will normally consist of air. But in the conceptual process design the oxygen level must be as low as possible. When this reactor is used in the conceptual process design, another gas, like N2, must be used as mixing gas. Even when the volume of the

vessel is very high the vessel remains very good mixed without large power requirements.28

Therefore the reactor is flexible enough if the turn down ratio is 50%. Airlift reactors have an

excellent heat exchange and temperature control!7. Due to good accessibility of the reactor the replacement of the microorganisms will not give problems. A disadvantage of the airlift reactor is th at because the microorganisms circulate through the bioreactor, the conditions change, and it is impossible to maintain consistent levels carbon source and nutrients for the biomass throughout the vessel!7. The separation of gas from the liquid is also not very efficient when foam is present.

Rgure2.7 (a) Intemal air/ift loop reador,(b) and extemal air/ift loop readoi6

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Design of a (Bio)process to remave nitrogen oxides from f1ue gases CPD 3332

2.5.5

Application for the cpd project

To make a choice which reactor can be applied for the conceptual process design, all the points mentioned above will be taken into account. With own reasoning skilIs and knowledge of the reactors, the pros and cons of the reactors are indicated. The most important aspects for the design are given in Table 2.5. The points are classified in order of importance.

Table2.5 Pros and cons of the discussed reactors

Type of reactor CSTS Fluidized Air lift reactor Packed bed bed

Usability

j

controllability

++

+j-

+

+j-Pressure drop

++

--

+j-

--capitaljoperation costs

+

-

++

-capacityjturn down ratio

+

-

+

-Mixing

++

+

-Heat removal

+

++

-When looking at Table 2.5, it appears that the continuous stirred tank bioreactor will be the best option for the conceptual process design. With the option usabilityjcontrollability is meant the control in temperature, pH and redox potential. This holds the ease of addingjmixing of ethanol, acid and nutrients for the bacteria. Because the environment for the biomass is very important, this point is the first in order in importance. To generate optimal reaction rates the microorganisms have to life under optimal conditions.

A large pressure drop over the bioreactor will contribute to high operating costs for the bioreactor. The pressure drop in a CSTB is almost negligible, and fits therefore the best in the conceptual process design.

With the aspect capacity, the turn down ratio of 50 % is taken into account. It appeared here that the CSTB is the reactor with the highest flexibility in operation. Due to a good choice in stirring devices, the CSTB is rated as the reactor with best mixing applications.

The heat removal in a CSTB will give no problems. When a lot of heat is generated, extra baffles can be installed with coolant passing through. From table 2.5 it appeared th at an airlift reactor looks also favorable for th is process, but the fact that another gas for mixing the stream is needed, will give problems in for instance cost when N2 is used. This, and the

pressure drop, is the main argument for choosing the continuous stirred tank bioreactor above the airlift reactor.