DESIGN FOR THE PRODUCTION OF SPONGE

IRON BY DIRECT REDUCTION PROCESS AND

SIMULATION OF MOVING BED REACTOR

A PROJECT REPORT

Submitted by

ANJANA VEL S (RegNo.21906203005)

ASMA HANIF (RegNo.21906203008)

KRISHNA ARCHANA (RegNo.21906203017)

in partial fulfillment for the award of the degree

of

BACHELOR OF TECHNOLOGY

IN

CHEMICAL ENGINEERING

SRI VENKATESWARA COLLEGE OF ENGINEERING

ANNA UNIVERSITY:: CHENNAI 600 025

APRIL 2010

ii

ANNA UNIVERSITY: CHENNAI 600 025

BONAFIDE CERTIFICATE

Certified that this project report “DESIGN FOR THE PRODUCTION OF

SPONGE IRON BY DIRECT REDUCTION PROCESS AND SIMULATION OF

MOVING BED REACTOR” is the bonafide work of “ ANJANA VEL.S, ASMA

HANIF & KRISHNA ARCHANA.P ” who carried out the project work under

my supervision.

DR. R. PARTHIBAN MS.T.KAVITHA

HEAD OF THE DEPARTMENT SUPERVISOR

DEPARTMENT OF CHEMICAL ENGINEERING DEPARTMENT OF CHEMICAL ENGINEERING

SRI VENKATESWARA COLLEGE SRI VENKATESWARA COLLEGE

OF ENGINEERING OF ENGINEERING

PENNALUR, PENNALUR,

SRIPERUMBUDUR – 602 105 SRIPERUMBUDUR – 602 105

SIGNATURE OF INTERNAL EXAMINER SIGNATURE OF EXTERNAL EXAMINER

iii

CERTICATE

This is to certify that the project report titled ‘Design for the production of

sponge iron by direct reduction process and simulation of moving bed reactor’

submitted by Anjana Vel.S, Asma Hanif and Krishna Archana.P to Sri

Venkateswara College of Engineering, Anna University, Chennai in partial

fulfillment of the requirements for the award of the degree of Bachelor of

Technology is a bonafide record of work done by them under the supervision of

Prof. P.S.T.Sai. The contents of this thesis, in full or in parts, have not been

submitted to any other Institute or University for the award of any degree or

diploma.

Dr.P.S.T.Sai

Professor

Department of Chemical Engineering

Indian Institute of Technology Madras

Chennai – 600 036.

iv

ACKNOWLEDGEMENT

We would like to express our sincere gratitude to Professor P.S.T.Sai, Department of

Chemical Engineering, IIT Madras for giving us the opportunity to work at the

institute and for guiding us throughout the course of our project.

We sincerely thank Dr. R. Ramachandran, Principal and Dr. R. Parthiban, the Head of

the Department of Chemical Engineering, Sri Venkateswara College of Engineering,

for the encouragement given to us to carry out our project work and gain practical

knowledge in the field of study.

We are grateful to Ms.T.Kavitha, Senior Lecturer, Department of Chemical

Engineering, Sri Venkateswara College of Engineering, for her support and guidance

at each stage of our project. We are thankful to her for her suggestions and for

providing us with the necessary reference books and study material.

We thank our Project Coordinator, Mr. R. Govindarasu, Senior Lecturer, Department

of Chemical Engineering, Sri Venkateswara College of Engineering, for his

assistance.

We extend our thanks to Ms. Sheril Mary Mathew, M.Tech student, IIT-Madras for

her timely help that ensured us to finish our project in time.

Finally, we thank the faculty and staff of the Chemical Engineering Department, Sri

Venkateswara College of Engineering for their support.

v

ABSTRACT

In this project, the shaft furnace reactor of the MIDREX direct iron reduction

process is simulated. This is a counter current gas-solid reactor, which

transforms iron ore pellets into sponge iron.

Simultaneous mass and energy balance along the reactor leads to a set of

ordinary differential equation with two points boundary conditions. The iron ore

reduction kinetics was modulated with the un-reacted shrinking core model.

Solving the ODE system using Runge-Kutta method predicts the concentration

and temperature profiles of all species within the reactor.

The model was able to satisfactorily reproduce the data of a MIDREX [5]

plants: Siderca (Argentina). Also, it was used to explore the performance of the

reactor under different operating conditions. This capacity could be used for

design and control purpose.

Design of suggested equipments - moving bed reactor (Shaft furnace), Heat

exchanger (Shell and tube) and Reformer (fixed bed catalytic reactor) is carried

out.

vi

TABLE OF CONTENTS

ABSTRACT………………………………………………………………….v

LIST OF FIGURES…………………………………………….....................ix

LIST OF TABLES………………………………………………………...…x

LIST OF SYMBOLS……………………………………………………...…xi

CHAPTER NO. TITLE PAGE NO.

1. INTRODUCTION

1.1 DIRECT REDUCTION OF IRON………..2

2. LITERATURE SURVEY

2.1 MIDREX………………………………….4

2.2 EQUIPMENTS INVOLVED……………..5

2.3 COMBUSTION TECHNOLOGY………..5

2.4 PROCESS FLOWSHEET………………...6

2.5 EQUIPMENT DETAILS

2.5.1 SHAFT FURNACE…………………7

2.5.2 REFORMER ………………………..8

2.5.3 HEAT RECOVERY UNIT………….9

vii

3. OBJECTIVE AND MODELLING

3.1 OBJECTIVES……………………………….11

3.2 MODELING

3.2.1 ASSUMPTIONS……………………....11

3.2.2 EQUATIONS INVOLVED…………...12

3.2.3 FORMULAE USED…………………..12

3.2.4 MASS AND ENERGY BALANCE

3.2.4.1 GAS PHASE ………….…….....14

3.2.4.2 SOLID PHASE……….…….......15

3.2.4.3 BOUNDARY CONDITIONS.....15

4. RESULTS AND DISCUSSIONS

4.1 EXTENT OF REACTION OF THE

REDUCING GASES………………………..19

4.2 TEMPERATURE PROFILES OF SOLID AND

GAS PHASE...................................................21

5. DESIGN OF EQUIPMENTS

5.1 DESIGN OF REFORMER.............................23

5.2 DESIGN OF HEAT EXCHANGER..............28

viii

6. POLLUTION AND SAFETY

6.1 ENVIRONMENTAL IMPACT.....................36

6.2 SAFETY REGULATIONS............................37

7. COST ECONOMICS ........................................39

APPENDIX : M atlab code for solving the set of simultaneous ODE .........47

REFERENCES..................................................................................................49

ix

LIST OF FIGURES

Figure No. Title Page No.

1 MIDREX – Flow sheet of direct reduction of iron............6

2 Shaft furnace geometry......................................................7

3 Fixed bed catalytic reactor.................................................8

4 Shell and tube heat exchanger.............................................9

5 Extent of reaction of H2 as a function of

shaft furnace length............................................................20

6 Extent of reaction of CO as a function of

shaft furnace length............................................................20

7 Temperature of gas as a function of shaft

furnace length.....................................................................21

8 Temperature of solid as a function of shaft

furnace length......................................................................21

9 Fixed bed catalytic reactor..................................................23

10 Plot of X vs

mr

1..............................................................................27

11 Single pass, parallel flow shell and tube heat exchanger....28

x

LIST OF TABLES

Table No. Title Page No.

1 List of equipments and their functions......................5

2 Operating conditions of Siderca plant......................16

3 Kinetics constants, effective diffusion

coefficient and heat transfer coefficient

used in the simulation................................................17

4 Calculation of rate constant of methane...................26

xi

LIST OF SYMBOLS

Ap pellet external area (cm2)

C reactor gas concentration (mol/cm3)

D effective diffusion coefficient (cm2/s)

Gm molar flow (mol/cm2s)

H reaction enthalpy (cal/mol)

h global heat transfer coefficient (pellets/gas) (cal/s/cm2/K)

k kinetics constant of the surface reaction (cm/s)

k g external mass transfer coefficient (cm/s)

L reduction zone length (cm)

M w molecular weight

n p number of pellets per unit volume (1/cm3)

R reaction rate (mol/cm3s)

^

R reaction rate per pellet (mol/s)

r0 external radius of the pellet (cm)

rc radius of the unreacted core (cm)

T temperature ( Co

)

u gas velocity (cm/s)

X extent of reaction/extent of reactant conversion (mol/cm3)

z space variable inside the reactor (cm)

xii

SUBSCRIPTS

atm atmosphere

i ith reaction

in reactor inlet

j jth reactant (gas or solid)

n gaseous reactant

rs reactive solid (Fe2O3)

ps product solid (Fe)

sol solid

g gas

GREEK LETTERS

α stoichiometric coefficient

ρ density of the solid reactant (g/cm3)

1

CHAPTER 1

INTRODUCTION

2

1.1 DIRECT REDUCTION OF IRON

In a direct reduction process [1], lump iron oxide pellets and/or lump iron ore,

are reduced (oxygen removed) by a reducing gas, producing direct reduced iron

(DRI). The reducing gas can be generated externally to the reduction furnace, or

can be generated from hydrocarbons introduced into the reduction zone of the

furnace. In the former case, the reducing gas is produced from a mixture of

natural gas (usually methane) and recycled gas from the reducing furnace. The

mixture is passed through catalyst tubes where it is chemically converted to a

gas that is rich in hydrogen and carbon monoxide.

Examples of processes that use variations of this general procedure include

Midrex and HYL. When the reducing gas is generated from hydrocarbons in the

reduction zone of the furnace, it is typically a rotary kiln furnace that uses

hydrocarbon fuels (primarily coal, but sometimes oil and natural gas) without

prior gasification in the reduction chamber. Examples include the ACCAR and

SL/RN processes.

Direct reduced iron is a virgin iron source that is relatively uniform in

composition, and virtually free from tramp elements. It is used increasingly in

electric furnace steelmaking to dilute the contaminants present in the scrap used

in these processes. It has an associated energy value in the form of combined

carbon, which has a tendency to increase furnace efficiency. For captive DRI

production facilities, there is the added advantage that the delivery of hot DRI to

the furnace can reduce energy consumption 16 to 20%.

3

CHAPTER 2

LITERATURE SURVEY

4

2.1 MIDREX

The charge is fed in continuously from the top of the furnace, passing uniformly

through the preheat, reduction and cooling zones of the furnace. The reducing

gas consists of about 95% combined hydrogen plus carbon monoxide. It is

heated to a temperature range of 1400° to 1700°F and is fed in from the bottom

of the furnace, below the reducing section. The gas flows countercurrent to the

descending solids. At the top of the furnace, the partially spent reducing gas

(approximately 70% hydrogen plus carbon monoxide) exists and is

recompressed, enriched with natural gas, preheated to 750°F, and transported to

the gas reformer.

The reformer reforms the mixture back to 95% hydrogen plus carbon monoxide,

which is then ready for re-use by the direct reduction furnace. In the cooling

zone, the cooling gases flow countercurrent to the DRI. At the top of the cooling

zone, the cooling gases exit and are sent to recycling, then return to the bottom

of the cooling zone. The cooled direct reduced iron (DRI) is discharged through

the bottom of the furnace, after which it is screened for removal of fines, and

treated to minimize the danger of spontaneous ignition during extended storage.

The reduced fines are briquetted to produce a usable DRI product.

5

2.2 EQUIPMENT INVOLVED

In addition to its main vertical shaft furnace, the Midrex process utilizes the

following major pieces of equipment in its production process.

Table 1. List of equipments and their functions

NAME OF EQUIPMENT FUNCTION

Charge Feed System

It introduces the charge into the top of

the furnace.

Heat Exchanger

It preheats the gases prior to

reforming.

Reformer

It converts the natural and recycled

gases into the reducing gas.

Cooling Gas Scrubber

It recycles the cooling gases that exit

from the DRI cooling zone of the

furnace.

Top Gas Scrubber

It recycles the furnace exhaust gases

prior to combustion in the gas

reformer.

Ejector Stack

It rejects the scrubbed flue gases to the

atmosphere.

2.3 COMBUSTION TECHNOLOGY

The DRI processes outlined in the Process Equipment section were either

natural gas or coal-based. A combination of fuel feed at the charge end of the

reducing furnace, and fuel injection at various stages of the process were

employed. The combustion processes in these furnaces are highly dependent

upon the manner in which the fuels are introduced into the furnaces.

6

The Midrex gas-based process utilizes a vertical shaft furnace that has preheat,

reduction, and cooling zones. The reducing gas is ported into the reduction zone

through a bustle pipe that is located at the bottom of the reduction zone, while

the charge is introduced at the top of the furnace. Combustion occurs as the

reducing gases flow countercurrent to the charge.

The excess top gases are combusted as fuel for the reformer burners. The hot

flue gases from the reformer are passed through heat recuperators to preheat

combustion gases for the reformer burners, and also to preheat the process gases

before reforming. Usage of the heat recuperators has significantly raised the

efficiency of the process.

2.4 PROCESS FLOWSHEET

Fig. 1 MIDREX – Flow sheet of direct reduction of iron.

7

2.5 EQUIPMENT DETAILS

2.5.1 SHAFT FURNACE

Fig. 2 Shaft furnace geometry

The main function of the shaft furnace is to generate sponge iron from iron ore.

The solids flow downwards by gravity and the reducing gases flow upwards in

counter current, while the corresponding chemical transformations occur. As it

can be observed, the furnace consists of a vertical cylindrical container, with a

conic lower zone. The inner wall is covered with insulating materials resistant

to erosion. The reducing gases enter by the middle zone of the reactor through

the bustle, which consists of a channel with approximately 70 nozzles that direct

the gas towards the center of the solid bed. Immediately underneath, the upper

burden feeders are located. Following in descendent order one can found: the

wind boxes (which take the cooling gas that circulates around the lower conical

zone of the reactor), cooling gas distributor or (“inverted Christmas tree”),

which besides to inject cooling gases has the function to support most of the bed

weight.

8

Within the MIDREX Plant, the main unit operation is the MIDREX Shaft

Furnace. Inside the MIDREX Shaft Furnace, the iron ore is converted to DRI.

The spent reducing gases called recycled process gas, exit the top of the Shaft

Furnace where they are cleaned and cooled. The recycled process gas contains a

mix of CO2 and H2O, unreacted H2 and CO, and any inerts (N2, etc.)

circulating in the system. To obtain a high enough gas quality for reuse in the

Shaft Furnace, most of the CO2 is removed from the recycled process gas via a

vacuum pressure swing absorption (VPSA) unit. After this step, the process gas

is mixed with the fresh synthesis gas. At this point, the gas is near ambient

temperature. Therefore, the gas must be heated to about 900 C (1650 F) prior to

entering the Shaft Furnace, to ensure maximize reduction efficiency.

2.5.2 REFORMER

Fig.3 Fixed bed catalytic reactor

To maximize the efficiency of reforming, off-gas from the shaft furnace is

recycled and blended with fresh natural gas. This gas is fed to the reformer, a

refractory-lined furnace containing alloy tubes filled with catalyst. The gas is

heated and reformed as it passes through the tubes. The newly reformed gas,

9

(containing 90-92% H2 and CO) is then fed hot directly to the shaft furnace as

reducing gas.

Equations Involved:

HCOCOCH 22 224

HCOOHCH 3 224

2.5.3 HEAT RECOVERY UNIT

Fig.4 Shell and tube heat exchanger

The thermal efficiency of the MIDREX Reformer is greatly enhanced by the

heat recovery system. Sensible heat is recovered from the reformer flue gas to

preheat the feed gas mixture, the burner combustion air and the natural gas feed.

In addition, depending on the economics, the fuel gas may also be preheated. A

gas-gas shell and tube heat exchanger is designed.

10

CHAPTER 3

OBJECTIVE AND MODELLING

11

3.1 OBJECTIVES

1. To obtain the temperatures (solid and gas) and concentrations of gases

( 22 ,COOH , 2H , 4CH ,CO ) and conversion profiles along the length of

the reactor for different operating parameters (gas flow rate, solid flow

rate, initial gas temperature, gas composition) and geometric

parameters(diameter, length of the reactor).

2. To identify the operable conditions for the reactor.

3.2 MODELING

3.2.1 ASSUMPTIONS

In order to model the MIDREX shaft furnace, the following approximations are

considered:

(a) The iron ore pellet consumption is governed by the unreacted shrinking core

model.

(b) Mass and heat transfer resistances through the film around the solid particle

are negligible comparing with diffusional resistance inside the porous solid

(kg _ D/2/r0).

(c) Only steady-state operating conditions will be considered.

(d) Plug flow is assumed for gas and solid phase.

12

3.2.2 REACTIONS INVOLVED

)()(3

2)(

3

12232

gsg OHFeHOFe ………………………… (R1)

)()(3

2)(

3

1232

gsg OCFeCOOFe ……………………...… (R2)

The Water Gas Shift Reaction:

HCOOHCO 222

It must be noted that the WGSR is a linearly dependent reaction such that R1 −

R2 = WGSR. It is a very important reaction in the reductor reactor studied. So

even it was not chosen as one of the linearly independent reactions of the

system, it is taken into account implicitly, in the present analysis.

3.2.3 FORMULAE USED

The extent of reaction is defined in terms of concentration as

i

iij

o

jj XCC ………………………………………. (1)

for each species j and where ij is the stoichiometric coefficient.

From definition (1), it can be written for the gaseous phase as,

CCX HH

o

221

……………………………………….. (2)

CCX CO

o

CO

2 ………………...………………………. (3)

For the solid phase,

XCCXXX OFeOFe

o

rs 321][3

3232

……………..…. (4)

13

Considering the unreacted shrinking core model [6] and that the concentration

of reactive solid must be measured per unit of reactor volume, it is possible to

relate the radius of the unreacted core ( r c) with the solid conversion ( X 3

)

through equation (5).

)4

( 333

1

n

MXrr

p

woc ……………………….……………. (5)

where ro is the external radius of the pellet,

n pis the number of pellets per unit reactor volume,

M wand ρ are the molecular weight and the density of the reactive solid,

respectively.

The shrinking core model is used for the solid pellet. The corresponding

reaction rate expression per pellet is given by:

)///1(

4

2

2^

DrrDrk

cr

nocncn

ncR

where n = 1, 2 denotes H2 and CO, respectively.

The reaction rate per unit volume of reactor (R) is obtained through:

^

RR np

The solid molar flow (Gmsol

) is a function of X 3 and is related to shrinking

core radius, such as Gmsol

can be evaluated using expression

)/)(/(

)/)(/2/1(

)(333

333

MrrMr

MrrMrGrG

ww

ww

mmpsrs

psrs

solsol

pscorsc

pscorscL

c

14

The solid specific heat was calculated using:

)]([

)]([

333

333

rrr

rrCrCC

copscrs

cops

ps

pcrs

rs

p

ps

This expression considers the Cp weighted average of reactive solid (rs) and

product solid (ps) for any state of transformation given by the unreacted radius

(rc). Values of reaction enthalpies (H) and specific heats (Cp) were taken from

Perry’s Hand Book. The kinetics and diffusion parameters take as reference

those that were obtained from experiments performed in a laboratory gas-solid

fixed bed reactor at 900 ◦C with SAMARCO pellets. Kinetics coefficients

follow Arrhenius law. The dependency of the effective diffusion coefficients

with temperature is taken as TDi

75.1

3.2.4 MASS AND ENERGY BALANCE

3.2.4.1 GAS PHASE

0),,(

)(

21

TXXCG

TTAnTg

gsolppg

pm

h

dz

d

gg

0),(*31

^

1

1 XXRnX

pdz

du

0),(*32

^

2

2 XXRnX

pdz

du

15

3.4.2.2 SOLID PHASE

3.4.2.3 BOUNDARY CONDITIONS

0)(1

LzX ;

0)(2

LzX ;

0)0(3

zX ;

TTin

ggLz )( ;

TT atmsz )0( ;

The problem is solved making an attempt to predict TXX g,,

21 at z = 0 (shaft

furnace gas outlet) so that after solving the equations system [3] the boundary

conditions at z = L i.e.) 0)(1

LzX , 0)(2

LzX and TTin

ggLz )( are

satisfied.

0)),(),((*32

^

231

^

1

3 XXRXXRnX

u psol dz

d

i

soliisoligsolp

sol

psol

TXXRTHTTATXCXG

nTh

pmdz

d

solsol

0)],,()()([*),()( 3

^

33

16

Table 2. Operating conditions of Siderca plant

Gas

Gas flow rate 1,40,000 Nm3//h

Inlet composition (at z = L)

H2 52.9%

CO 34.7%

H2 O 5.17%

CO2 2.47%

CH4+ N2 4.65%

Inlet temperature 957 0 C (1230 K)

Solid

Production (Fe) 100 t/h

Mineral pellet density 3.4 g/cm3

Sponge iron density 3.1 g/cm3

Pellet ratio (r0 ) 0.5 cm

np 0.99 pellets/cm3

Reactor

Reaction zone length 1000 cm

Diameter 488 cm

17

Table 3. Kinetics constants, effective diffusion coefficient and heat transfer

coefficient used in the simulation

k1 0.225 exp (−14700/82.06/T) cm/s

k2 0.650 exp (−28100/82.06/T) cm/s

D1 1.467 × 10-6 × T0.75 cm2 /s

D2 3.828 × 10-6 × T0.75 cm2 /s

h 4 × 104 cal/cm2/s/K

18

CHAPTER 4

RESULTS AND DISCUSSION

19

The reduction zone of the shaft furnace of the MIDREX direct reduction

process was simulated. In order to do this, mass and energy balances were

solved for each phase in the counter current gas-solid reactor and the extents of

reaction for hydrogen and carbon monoxide and temperature profiles for gas

and solid are studied.

The resolution of the differential equations system allows knowing the

evolution of several variables throughout the reactor. The model satisfactorily

fit the data from MIDREX plant (Siderca SA in Argentina). In addition, they

allow exploring the behaviour of the reactor for different operating conditions.

The kinetics used in the simulations was found in a laboratory scale reactor.

Also some parameters were slightly adjusted for the plant in order to fit the

available plant data and taking into account the different type of iron ore pellets

and operation conditions.

4.1 EXTENT OF REACTION OF THE REDUCING GAS

It can be seen that the two extents of reaction are very similar even when the H2

concentration is greater than that of CO. This indicates that the different

concentrations used in the reducing gas allow that both reducing gases act

simultaneously throughout the entire reactor, removing the same amounts of

oxygen. Also it is clear that the CO is a better reducer than the H2 since with

smaller concentrations of CO similar reaction rates are achieved.

20

Fig.5 Extent of reaction of H2 as a function of shaft furnace length

Fig.6 Extent of reaction of CO as a function of shaft furnace length

21

4.2 TEMPERATURE PROFILES OF SOLID AND GAS PHASE

Fig.7 Temperature of gas as a function of shaft furnace length

Fig.8 Temperature of solid as a function of shaft furnace length

22

CHAPTER 5

DESIGN OF EQUIPMENTS

23

5.1 DESIGN OF REFORMER

Determining the catalyst weight necessary to achieve 80 percent conversion of

methane in a fixed bed reactor with a bulk density of 2.5/cm3

The reaction taking place inside the reformer is:

Fig.9 Fixed bed catalytic bed reactor

HCOCOCH 22 224

24

Solution:

Design equation

The design equation [9] for a tubular reactor that has catalytic reaction

occurring in it is given by:

mmo rdW

dXF

Neglecting pressure drop and catalyst decay we can integrate to obtain

X

m

mor

dXFW

0

Rate law:

mmcoco

mcomm

PKPK

PPkKr

1

Stoichiometry:

X

XP

X

XRTCRTCP moomomm

1

1

1

1

12.02*06.0 omoPy

XPP mom 1

XPP COmoCO 22

mm r

dW

dF

25

66.306.0

22.02 co

XPP moco 66.32

XPXRTCRTCP comocoomococo 22

0co

XXPco 168.02084.0

Note that moP designates initial methane pressure. Here, initial total pressure is

designated as oP .

On combining we get

X

COmom

mmocomomo

X

m

moXXPkK

dXXKPXKPF

r

dXFW

0

22

0 221

11

Parameter Evaluation

Substituting for 2, Hco PP and mP in

scatg

methanemol

PP

PPr

mco

mCO

m..

.

129.026.11

104.12

8

Yields,

min..

.

129.00848.1

166.3109238.5 6

catKg

methanegmol

X

XXrm

26

min7.178000

463082.0

084.0

0

0

mol

RT

PvCF mo

momo

min14

gmolXFF moco

796.07.17

14X

We can now proceed to solve for the catalyst weight necessary to obtain this

conversion by using

X

XXrm

129.00848.1

166.3109238.5 6

Numerical technique:

Table 4. Calculation of rate constant of methane

X mr

mr

1

0 1.998x10-5 50033.67

0.1 1.729x10-5 57834.22

0.2 1.476x10-5 67730.66

0.3 1.240x10-5 80636.1

0.4 1.019x10-5 98074.5

0.5 8.143x10-6 122792

0.6 6.2388x10-6 160285.1

0.7 4.4765x10-6 223385.8

0.8 2.852x10-6 350601.8

27

Fig.10 Plot of X vs

mr

1

Using numerical integration, by Simpson’s rule we have

)242424

387654321

1 ffffffffh

FW mo

350601.8)223385.8(4)160285.1(2)122792(4

)98074.5(280636.1467730.66257834.22450033.67

3

1.07.17

41064.17 Kg of catalyst.

Since the bulk catalyst density is 2.5 g/cm3, the reactor necessary volume is

.10056.7 3litresV

DESIGN SUMMARY:

Required weight of the catalyst 41064.17 Kg

Volume of the reactor litresV 310056.7)(

28

5.2 DESIGN OF HEAT EXCHANGER

A gas-gas exchanger [2] is to be designed to exchange heat between natural gas

and flue gases from the reformer. The natural gas is to be heated from 298K to

463K and the flue gas is to be cooled from 700K to 500K

Fig.11 Single pass, parallel flow shell and tube heat exchanger

Solution:

gT =427-25=402 oC

sT =227-190=37 oC

mT = (402-37)/ln(402/37)=153.00o C

)( mpTmq c

m=10000 kg/hr = 2.778 kg/s

cp= 1.151 kJ/kgK

29

Required heat duty is: q = 2.7778(1.151x103)(153)

= 489 kJ/s (i.e.KW)

TUBE AREA:

Dimensions of the tube are taken as per the standard specifications [7].

Dimensions of the tube:

O.D. =3.175cm=0.03175m

Number of tubes= 283

Length of tube =4.572 m

External surface of all the tubes per unit length of the tubes

= Πd x No.of tubes

=π(0.03175)283 =28.22m

Dimensions of the shell:

Shell side nozzles are 40.64cm in diameter

Flow area of the nozzles=Af= )(4

4064.02

=0.1297 m2

Shell side mass density, =0.525 m

kg3

Shell side velocity = )1297.0525.0(

7778.2

V=40.79 m/s.

30

TUBE SIDE FILM COEFFICIENT:

Sieder-Tate correlation:

w

Nu

14.0

33.08.0

PrRe027.0

Cp=1.29 kgK

kJ

µ=1.027x10-5 m

sN2

k=0.031 K

W

m2

k

Cp .Pr

031.0

)10027.1()1029.1( 53 rP =0.7

Tube side mass density, 0.717m

kg3

Cross sectional area of one tube= )0257.0(2

4

=518.75x10-6 m

2

Tube side mass flow rate=1.768 kg/sec.

Mass flow rate per tube= 1.768/283=6.25x10-3 kg/sec.

31

Tube side velocity, u= ./32.16)1075.5184411.2(

01864.06

sm

2.3017910027.1

)717.0)(32.16)(0257.0(..Re

5

ud i

Re=30179 > 10000

04.92)0.1)(7.0)(30179(027.0 33.08.0 uN

kNu dh ii

K

Wk

mdh

i

i 2)0257.0(

)031.0(04.9204.92

SHELL SIDE FILM COEFFICIENT:

w

se

kNu

GDDh eo

14.0

33.0

55.0

Pr.

36.0.

Triangular Pitch, Pt=0.04445m

ddPt

Do

e

o

]4/5.043.0[822

, md o

03175.0

03175.0

]4/03175.05.043.0[8 22

04445.0

De

=0.0364 m

Tube Clearance:

dP otC

= 0.04445 m - 0.03175 m = 0.0127 m

No. of baffles = 8.

Total length of the tube=4.572 m.

32

Distance between baffles, mmLB5715.0

8

572.4

Flow area of the tube bundle,Pt

BCDA

s

s

..

Shell diameter, mDs016.1

mAs

2166.0

04445.0

)5715.0()0127.0()016.1(

Shell side mass flow rate, m=2.7778 kg/sec.

s

kgm

mAG

s

s 273.16

166.0

7778.2

Shell side viscosity, µ= 3.16810x10-5

m

sN2

19222)10168.3(

)73.16()0364.0(.Re

5

GD se

93.70)0.1()65.0()19222(36.0 33.055.0 uN

93.70.

k

Dh eo

K

W

mho 2

13.1090364.0

056.093.70

Also, using jH factor, we know that jH=70 from the jH vs Re graph.

651.0056.0

5^10*188.3*3^10*151.1.Pr

k

Cp

33

The fouling factors on each shell and tube side are 0.00025K

W

m2

468.93

100025.000025.0

4.271

1

1

U

U=67.18K

W

m2

Surface area required= mT m

U

q 256.47

153*138.67

74.488541

)(

Available area=57m2

% Excess area = %84.19100*56.47

56.4757

For Baffles with 25% cut:

jH=87

0.10364.0

056.08765.0Pr

33.0

14.0

33.0

w

Kj

Dh

e

H

o=116.16

K

W

m2

The fouling factors on each shell and tube side are 0.00025K

W

m2

16.116

100025.000025.0

4.271

1

1

U

34

U=78.16K

W

m2

Surface area required= mT m

U

q 288.40

15316.78

74.488541

)(

Available area=57m2

% Excess area = %43.3910088.40

88.4057

DESIGN SUMMARY:

Shell side:

Shell diameter, mDs016.1

Overall heat transfer co-efficient U=67.18K

W

m2

Surface area required= 47.56 m2

% Excess area = 19.84%

Tube side:

O.D. = 0.03175m

Number of tubes= 283

Length of tube = 4.572 m

Overall heat transfer co-efficient U=78.16K

W

m2

Surface area required= m2

88.40

% Excess area = 39.43%

35

CHAPTER 6

POLLUTION AND SAFETY

36

6.1 ENVIRONMENTAL IMPACT

Worldwide, there is an increasing emphasis on environmental issues and the

steel industry is under intense scrutiny. The emissions from integrated steel

mills, especially sinter plants and coke ovens, are a particular concern. As a

member of the global community, Midrex along with its parent company, Kobe

Steel, recognizes the importance of the environment and the limits to natural

resources. Through the years they have been proactive in reducing the

environmental impact of their iron and steelmaking processes.

Their focus is on avoiding pollution rather than controlling and treating it, with

a goal of zero emissions and waste. Steelmakers cannot continue to landfill

wastes and treat gaseous emissions at the "end of the pipe." Most MIDREX DR

Plants are designed for 100% recycle of process gases and water.

Reducing energy use goes hand-in-hand with environmental benefits because it

automatically leads to lower emissions. In the US it now takes about one-half

the energy to produce a ton of steel as it did in 1975. State-of-the-art MIDREX

Plants are extremely energy efficient, with natural gas consumptions as low as

2.3 net Gcal/t DRI.

MIDREX Plants are designed to minimize air, water, and noise pollution.

Emission levels meet all applicable 1998 World Bank standards. Carbon

dioxide emissions are becoming a concern and in this regard, the MIDREX

Plant steelmaking has an advantage versus the traditional blast furnace route. In

some cases, a MIDREX Plant facility has only one-third the carbon emissions

per ton of steel as a Blast Furnace complex.

37

Another environmental issue is the disposal of metal-bearing wastes, both

ferrous and non-ferrous. Integrated steel facilities and mini-mills are finding it

more and more difficult to dispose wastes from iron and steelmaking processes.

These wastes include iron oxide screenings, mill scale, and bag-house dust. In

addition, there are millions of tons of metal-bearing waste materials stockpiled

around the world from various mining and processing operations. With the

commercialization of the FASTMET Process, there is now a cost-effective

means of dealing with these wastes.

Steel mill wastes such as BF dust, BF Sludge, BOF dust, Sinter dust, EAF dust,

mill scale, and etc., will be processed by FASTMET plant. Benefits from this

application will be:

Elimination of waste disposal cost and landfill liability.

Waste changed to a quality source of iron (DRI).

6.2 SAFETY REGULATIONS

These are some of the safety measures adopted in the Midrex Plant:-

Coke oven doors, jambs and other equipment should be designed so as to

minimize the occurrence and magnitude of leaks.

Leaks from coke oven doors, lids, and other equipment should be

eliminated or reduced through a comprehensive operation and

maintenance programme designed for that purpose.

A periodic air monitoring system should be instituted, in order to

establish “regulated areas” where the exposure limit for coke oven

emissions is exceeded.

38

A respiratory protection programme should be instituted for workers in

regulated areas.

Coke oven workers should receive regular medical surveillance,

particularly focusing on the early detection of cancer, with appropriate

follow-up.

Mobile coke oven machines should be designed for safe entry and exit,

and provided with travel alarms. Windows should be kept clean and free

of obstructions. Where necessary, cameras or other devices should be

installed to allow the operator to see all sides of the machine.

Workers exposed to hot surfaces or radiant heat from open ovens should

be provided with appropriate protective equipment, and covered by a heat

stress prevention programme.

Measures should be taken to eliminate or reduce the escape of hazardous

substances during maintenance operations, when samples must be taken

for laboratory analysis, and during barge, truck and railcar loading.

39

CHAPTER 7

COST ECONOMICS

40

EQUIPMENT COST [8]

S.No Equipment Quantity Rate

(Rs)

Total cost

(Rs)

1 Shaft Furnace 1 50,00,000 50,00,000

2 Reformer 1 30,00,000 30,00,000

3 Heat exchanger 1 10,00,000 10,00,000

TOTAL COST 90,00,000

DIRECT COST ESTIMATION AMOUMT (Rs.)

Purchase Equipment Cost 90,00,000

Installation Cost 36,00,000

Instrumentation Cost 9,00,000

Piping cost 13,50,000

Electrical Cost 40,50,000

Building Cost 40,50,000

Land Cost 9,00,000

TOTAL AMOUNT (Rs.) 2,38,50,000

41

INDIRECT COST ESTIMATION

AMOUNT (Rs.)

Engineering and supervision 27,00,000

Construction Expenses 9,00,000

Contractors Fee 11,92,500

Contingency 19,08,000

TOTAL COST (Rs.) 58,00,500

TOTAL CAPITAL INVESMENT:

Fixed Capital = Total Direct Cost +Total Indirection Cost

= Rs. 2,96,50,500

Working Capital = 25% Fixed Capital

= Rs. 74,12,625

Start up Capital = 10% Fixed Capital

= Rs. 29,65,505

Total Capital = Fixed Capital + Working Capital + Start up

Capital

= Rs. 3,73,59,630

42

DIRECT PRODUCTION COST

Raw Material Cost per Kg

(Rs.)

Quantity per year

(Kg)

Amount

(Rs.)

Iron Ore 10 43,50,000 4,35,00,000

Coal 4.5 15,00,000 67,50,000

Lime stone 2 18,000 36,000

Filter aid 0.3 36,00,000 10,80,000

TOTAL AMOUNT (Rs.) 5,13,66,000

OPERATING LABOUR COST

S.No. Designation No. Salary per

month (Rs.)

Total

(Rs.)

1 General Manager 2 50,000 1,00,000

2 Production Manager 3 20,000 60,000

3 Engineers 6 15,000 80,000

4 Supervisors 9 10,000 90,000

5 Skilled Workers 20 6,000 1,20,000

6 Clerks 5 6,000 30,000

7 Safety Manager 2 20,000 40,000

Total Cost (Rs.)5,20,000

43

COST OF UTILITIES

UTILTY COST PER YEAR (Rs.)

Electricity 50,00,000

Steam 25,00,000

For a year,

Operating Labour Cost = 5,20,000 x 12 = Rs. 62,40,000

Clinical cost (5%OLC) = Rs. 3,12,800

Maintenance and repair = 5% of the Fixed Capital

= Rs. 14,82, 525

Depreciation of

Equipments

= 10% Fixed Capital

=Rs. 29,65,050

Insurance Tax = 1% Fixed Capital

= Rs. 2,96,505

Plant Overhead Cost (POC) = 40% OLC +Supervision Cost + Maintenance

Cost

= Rs. 66,78,525

General Expenses = 10% POC

= Rs. 6,67,825.5

44

REVENUE

The market cost of 1ton of Iron is approximately around Rs. 30,000

For a year Rs. 90,00,000 ( this inclusive of all taxes)

PAY BACK PERIOD

=(2,96,50,500+29,65,505)/( 29,65,050+90,00,000)

= 2.725 years

=2 years and 9months

Historically, gasifier/DRI plant combinations have not been built because they

have been uneconomical relative to alternative methods of producing iron. This

is primarily due to the significant capital costs associated with building both a

gasification unit and most of a standard DRI Plant. The high capital charge has

negated the cost benefits associated with using coal as the primary fuel source.

The following table provides an approximate breakdown of the cost inputs to

produce DRI, utilizing a coal gasification unit as the synthesis gas source.

Unfortunately, in the U.S., a total cost of US$ 130 /t is high relative to currently

operating natural-gas based DRI Plants, coal based iron-making plants, and high

quality scrap steel. A good target total cost is US$ 120/t or less. Thus, cost

savings must be found for a gasification based DRI plant to be viable in the U.S.

However, in other parts of the world, the gasification/MIDREX Plant

combination can be more cost competitive.

45

This is due to a variety of reasons, including:

(1) lower cost iron ore and coal,

(2) use of low cost petroleum refining by-products

(3) lower cost construction labor,

(4) a shortage of existing local steelmaking capacity,

(5) environmental pressures to close existing coal based iron-making plants.

Some or all of the above competitive benefits for the gasifier/DRI Plant

combination are found in the major steelmaking regions of China, India, S.

Korea, Brazil, South Africa, and Western Europe. In these countries, the

availability of low cost electricity for use in the steel mill is critical. If low cost

fuels are available, then the gasification unit could be sized to also make enough

synthesis gas for an adjacent power plant to produce the necessary low cost

electricity.

The keys to making a gasifier/MIDREX Plant a viable option is:

Where available, utilize excess synthesis gas production from a separate

gasification-based power plant or chemical plant. Depending on local

conditions, if an acceptable quality synthesis gas is available “across-the-

fence” for less than about US$ 3.00/MMBtu, then a DRI plant may be a

viable option.

Minimize the capital cost of the gasification plant.

46

If possible,

- No desulfurization system

- No particulate scrubbing system

- No hydrocarbon removal system

- No gas expander

Maximize integration of the gasification and MIDREX unit operations.

Utilize low-cost petroleum refining by-products where available.

Build an integrated mini-mill complex, which includes a power plant.

47

APPENDIX 1

The code [10] for solving the set of simultaneous ordinary differential equations

with the given two point boundary conditions so as to predict the extent of

reaction of the reducing gases and the temperature profiles of the reactor is

given below:

function dydz=final(z,y)

dydz=zeros(size(y));

x1=y(1);

x2=y(2);

Tg=y(3);

x3=y(4);

Tsol=y(5);

%Number of pellets per unit volume(1/cm^3)

np=0.99;

%Gas velocity(cm/s)

u=(208.02*(298/Tg));

%External radius of the pellet(cm)

ro=0.5;

%Kinetics constant of the surface reaction(cm/s)

k1=0.225*exp(-14700/(82.06*Tsol));

k2=0.650*exp(-28100/(82.06*Tsol));

%Effective diffusion co-efficient(cm^2/s)

D1=(1.467*(10^-6))*(Tg^1.75);

D2=(3.828*(10^-7))*(Tg^1.75);

%Reactor gas concentration(mol/cm^3)

CH21=((1.4*0.529)/(82.06*Tg));

CCo1=((1.4*0.347)/(82.06*Tg));

%Radius of the unreacted core(cm)

rc=(0.125-(1.45*x3))^0.33;

fprintf('rc=%5.5f\n',rc);

%Specific heat(cal/deg.mole)[4]

cprs=(24.72+(0.01604*Tsol)-(423400*Tsol^-2));

cpps=(4.13+(0.00638*Tsol));

Cpsol=((cprs*3.4*rc^3)+(cpps*3.1*(0.125-

rc^3)))/((3.4*rc^3)+(3.1*(0.125-rc^3)));

Cph2=6.62+(0.00081*Tg);

48

Cpco=6.60+(0.00120*Tg);

Cpg=(((CH21-x1)*Cph2)+((CCo1-x2)*Cpco))/((CH21-x1)+(CCo1-x2));

%Pellet external area(cm^2)

Ap=3.142;

%Global heat transfer co-efficient (pellets/gas)(cal/s.cm^2.K)

h=4*10^-4;

%Molar flow rate(mol/cm^2.sec)

Gmg=((9.605*10^-3)*(298/Tg));

Gmsol=(9.283*10^-4)*(((0.5*(rc)^3*.02125)+((0.125-

(rc)^3)*0.055))/(((rc)^3*0.02125)+((0.125-(rc)^3)*.055)));

%Solid velocity(cm/sec)

Usol=0.0479;

%Reaction enthalpy(cal/mol)

H1=8371.765+(-0.5133*Tsol)+(-4.44*10^-3*0.5*Tsol^2)-

(1.1433*10^5/Tsol);

H2=-2393.473+(2.596*Tsol)+(-5.04*10^-3*0.5*Tsol^2)-

(1.05*10^5/Tsol);

%Gas phase:

dydz(1)=(1/u)*((np*4*pi*(rc^2)*(CH21-x1))/((1/k1)+(rc/D1)-

((rc^2)/(ro*D1))));

fprintf('X1=%15.9e\n',dydz(1));

dydz(2)=(1/u)*((np*4*pi*(rc^2)*(CCo1-x2))/((1/k2)+(rc/D2)-

((rc^2)/(ro*D2))));

fprintf('X2=%15.9e\n',dydz(2));

dydz(3)=((np*Ap*h*(Tsol-Tg))/(Gmg*Cpg));

fprintf('Tg=%10.5f\n',dydz(3));

%Solid phase:

dydz(4)=(np/Usol)*(((4*pi*(rc^2)*(CH21-x1))/((1/k1)+(rc/D1)-

((rc^2)/(ro*D1))))+((4*pi*(rc^2)*(CCo1-x2))/((1/k2)+(rc/D2)-

((rc^2)/(ro*D2)))));

fprintf('X3=%15.9e\n',dydz(4));

dydz(5)=(np/(Gmsol*Cpsol))*((Ap*h*(Tsol-

Tg))+((H1*(4*pi*(rc^2)*(CH21-x1))/((1/k1)+(rc/D1)-

((rc^2)/(ro*D1))))+(H2*(4*pi*(rc^2)*(CCo1-x2)/((1/k2)+(rc/D2)-

(rc^2)/(ro*D2))))));

fprintf('Tsol=%10.5f\n',dydz(5));

end

49

REFERENCES

1. Daniel R. Parisi, Miguel A. Laborde, Modeling of counter current moving

bed gas-solid reactor used in direct reduction of iron ore, Chemical

Engineering Journal 104 (2004) pp.35–43.

2. Donald Q.Kern, ‘Process Heat Transfer’, McGraw Hill International

Edition.

3. J.R. Dormand, P.J. Prince, A family of embedded Runge-Kutta formulae,

J. Comp. Appl. Math. 6 (1980) 19.

4. O.Kubaschewski, C.B.Alcock, ‘Metallurgical Thermochemistry’,

Pergamon Press.

5. W.D. Munro, N.R. Amundson, Solid-fluid heat exchange in moving beds,

Ind. Eng. Chem. 42 (1950) 1481.

6. Octave Levenspiel, ‘Chemical Reaction Engineering’, John Wiley &

Sons.

7. R.H.Perry, D.W.Green, ‘Perry's Chemical Engineers' Handbook’,

McGraw-Hill.

8. Peters And Klaus D. Timmerhaus., ‘Process Economics and Plant

Design’, McGraw Hill International Edition.

9. H.Scott Fogler, ‘Elements of chemical reaction engineering’, Prentice-

Hall of India Pvt. Ltd.

10. William J.Palm, ‘Introduction to Matlab 7 for engineers’, McGraw-Hill.