oxychlorination fluid bed reactor modelling...

OXYCHLORINATION FLUID BED

REACTOR MODELLING AND

SIMULATION

w w w . c f d o i l . c o m . b rw w w. c f d o i l . c o m . b r

Jeã Carlos Santos Moreira (RLAM/OT)

<[email protected]>

PVC Trade

The PVC world demand for the next 5 years is estimated in 10 million

ton but only 6,5 million tons will be added in new plant capacities.

In Brazil 560 thousand ton until 2010.

Experts forecast an yearly growth rate of 4,5% for the PVC world

consumption, mainly because in the emerging countries such as China,

India and Brazil. Projections for South America and Brazil indicate

w w w . c f d o i l . c o m . b r

India and Brazil. Projections for South America and Brazil indicate

yearly growth of 5% and 7%, respectively.

The South America PVC production capacity is only 1,4 million tons

a year and the per capita consumption is of 4 kg for inhabitant a year

in Brazil and 3 kg in Argentina. On the other hand, in Europe and in

the United States, 15 kg of PVC are used by inhabitant a year.

In China, where the market PVC expansion is one of the largest of

the world, the consumption is of 7 kg.

The goal was develop a model for the

simulation a fluidized bed oxychlorination

reactor in order to investigated through

the sensitivity analysis the effect of

different parameters which influence the

reactor performance and then apply

Problem characterization


reactor performance and then apply

optimization using a commercial software.

Oxyreactor Parameters

Frmf is the Froude number at minimum fluidization conditions.

characterizes the quality of the fluidização (<0,13 indicative of particle fluidization)

εmf uo/umf Hmf (m) H (m) Frmf dp (µm) deq (m)

0,48 436 8,1 11,1 0,00151 45 0,27


indicative of particle fluidization)

uo/umf is the ratio between the feeding speed and the

fluid surface velocity in the minimum fluidization

conditions. Indicates the potential of particle drag or the bed channeling

The two-phase fluidized bed

model (bubble – emulsion)

Fluid Surface velocity in

minimal fluidization condition:

Particle Fluidynamic


minimal fluidization condition:

µ

ρρ

1650

g)(du

gs

2

p

mf

−=


Bubble diameter:

4,0

mfo

2

m,eq )uu(D374,0d −= π

D

Z3,0

e)dd(dd−

−−=

4,0

mfooo,eq )uu(A374,0d −=


Relative velocity:

Absolute Velocity:

brmfob uuuu +−=

2/1

bbr )gd(711,0u =

Do,eqm,eqm,eqeq e)dd(dd −−=


Void fraction at minimal fluidization:

Catalytic bed Porosity:

021,0029,0

1586,0

=p

f

mfAr ρ

ρε

−=

2

3

)(µ

ρρρ pf

fp

dgAr


Bed height:

)1)(1(1 mfεδε −−−=

f

mf

f

mf

mf

f

1

1

L

L

ρ

ρ

ε

ε=

−

−=

Hypotheses

The two-phase fluidized bed is a valid model

The one-dimension model with heat and mass axial scattering is valid

The chemical reaction occurs only during the emulsion phase

The catalyst particle elutriation outward the bed is not being considered


The catalyst particle elutriation outward the bed is not being considered

The chosen reaction network is represented by five chemical reactions:

oxychlorination, dehydrochlorination of the 1,2-dichloroethene, oxidation

of ethene for CO and CO2, formation of chloride

Diffusivity and the coefficient of convection remain constant in the axial

direction

The reactor is adiabatic

Mass Balance

Bubble Phase:

Emulsion Phase:

−=

b

ib

e

ie

beb

ib

Q

Q

)k(A

dz

d

z

e

ie

F

iF

e

ie

b

ib ieQ

Q

Q

Q

σ−

−=

−

b

ibbe

iu

K )(=σ


Emulsion Phase:

i

z

e

ie

F

iFb

H

0

ibbeieFie r).1)(1(VdzeQ

N

Q

NA)k(NN i εδσ −−+

−+= −∫

i

z

e

ie

F

iFbbiF

F

eie r).1)(1(V)e1(

Q

N

Q

NAuN

Q

QN i εδσ −−+−

−+= −

Energy Balance

Bubble Phase:

)TT()H(dz

dTCu ebbbe

bpffb −−=ρ

FeFb TT;TT0z ==⇒=

∫∫zT

dTb


∫∫ =−

z

0

bbe

T be

bpffb dz)H(

)TT(

dTCu

b

F

ρ

z)H(TT

TTln.Cu bbe

Fe

bepffb −=

−−

ρ

z

Feeb e)TT(TT β−−−=pffb

bbe

Cu

)H(

ρβ =

Energy Balance

Emulsion Phase:

+−+−−− ∫=

dzA)TT()H()TT(CQ)TT(CQ beb

H

0z

bberefeo,f,po,fo,erefFF,f,pF,fF,e ρρ

0)TT(UAr)H()1)(1(V eww

n

1j

jrj =−+−−−+ ∑=

∆εδ


1j

∑=

ww

io

ii Ah

1

kL2

)D/Dln(

Ah

1

1UA

++=

πi

DiD

kNuh =4,05/4

DD PrRe023,0Nu =

66,0

g

p

8,0

pg

ps

23,0

ogp

43,0

g

pgg

p

g

RwC

Cud

K

C)1(

d

K.C.01844,0h

−=

ρ

ρ

µ

ρρε

Mathematic Model-Molecular Diffusion

The “i” component mass flow is given by the Fick´s Law

for molecular diffusion:

Diffusion coefficient for no polar gases (Equation of

Chapman-Enskog):

imii xcDN ∇−=


Chapman-Enskog):

To polar componentes:

(Brokaw)

2/1

BAD

2

AB

2/33

ABM

1

M

1

P

T10bxD

+= −

Ωσ

∂+=

*

2

AB*

DDT

196,0ΩΩ

Mathematic Model-Molecular Diffusion Coefficient

Stefan-Maxwell:

The effective diffusivit is obtained by considering αi = 0

(diffusion controlled by chemical reaction over the

)NyNy(D

1

Ny)y1(N

D

ijji

n

ij j,i

n

ij

jiii

m,i

−

−−

=

∑

∑

≠

≠


(diffusion controlled by chemical reaction over the

surface of the catalyst):

Then,

∑ ∇−= iiefiiii ycDNyN α

∑≠=

−

=n

ji1j i

j

j

i

j,i

i

ief

N

N

y

y1

D

y

1D

Kinetic Model

For the main reaction a kinetics model following the

Langmuir-Hinshelwood theory was used. For the

remaining 4 reactions, empirical kinetics models

proposed by Gel'Perin et al. (1979a) were used.

OHClCHCH2HCl1/2OHCCuCl2 + →++


Langmuir-Hinshelwood

kr = 269exp(-37,8/RT) ; Ka = 0,63

OHClCHCH2HCl1/2OHC 2222C240T

CuCl2

242 o+ →++

=

Ea

cEar

CK1

CCKkr

+=

Algorithm

METHODS:

Newton-Raphson

Finite Differences

Gauss

ESTIMATE VALUES

Te e e(I)

CALCULATE THE MASS BALACE

FUCTIO FOR THE BUBBLE AD

EMULSIO PHASE (OLIEAR

ALGEBRAIC EQUATIOS SYSTEM)


SIM

Gauss

O

O

e(I)(k+1) - e(I)(k) < EMAX ?

PRIT Te, e(I)

ED

YES

YES

CALCULATE THE EERGY BALACE “F”

FUCTIO AD EMULSIO PHASE

TEMPERATURE

IS THE EERGY FUCTIO

BALACE F = 0 ?

Results – Conditions for Model

Reactor pressure: 3,8 kgf/cm2g

Temperature of the emulsion phase: 240 ˚C

Feed temperature: 154 ˚C

Temperature of the cooling water: 200 ˚C


Parameters Plant ModelRelative

Deviation %

Residence time, s 35,35 35,27 -

Ethene conversion, % 96,39 94,39 2,08%

Ethene gain, % 97,61 95,99 1,65%

Results – Emulsion PhaseTemperature

The temperature rise reduces the selectivity of the reaction.

Such reduction in the selectivity as well the temperature rise

can be observed by the higher formation of CO and CO2.

0,95

1,00

0,96

0,98

1,00

Eth

en

e C

on

ve

rsio

n


0,80

0,85

0,90

0,95

0,86

0,88

0,90

0,92

0,94

0,96

200 210 220 230 240 250E

the

ne

Co

nv

ers

ion

Eth

en

e G

ain

Emulsion phase temperature, ˚C

Ethene Gain

Ethene Gain (Plant)

Ethene Conversion

Results – Bubble Diameter

When the bubble diameter

increases mass transfer

between the bubble phase

and the emulsion

decreases. Besides, an

increase in the bubble 0,93

0,94

0,95

0,96

Eth

en

e C

on

ve

rtio

n

XET


increase in the bubble

diameter favors a higher

transportation of non-

reacted gas through the

bed, once reagents will not

meet the catalyst particles.

0,90

0,91

0,92

0,35 0,45 0,55 0,65 0,75 0,85

Eth

en

e C

on

ve

rtio

n

deq (m)

Results – Bed Height

0,93

0,94

0,95

0,96

Eth

en

e C

on

ve

rsio

n

zmin

Only about 5% of the catalyst is

used for the reaction. Light

variations in bed height at

minimum fluidization have

outstanding effects over the

conversion up to 8 m. Beyond

this value, any trials to increase


0,91

0,92

0,93

2 4 6 8 10

Hmf (m)

Eth

en

e C

on

ve

rsio

n

this value, any trials to increase

the catalyst concentration in the

bed to promote increases in the

ethene conversion, become

economically not feasible. For the

expanded bed, this point

corresponds to 9 m height.

Results – Residence Time

0,98

0,95

0,96

Decreasing in the gas surface velocity, the residence time

increases and the average bubble size decreases as the

interface mass and energy transfer increases. An increase in

residence time will produce a reduction in selectivity what can

be explained by the parallel reactions.


0,86

0,90

0,94

0,87

0,88

0,89

0,90

0,91

0,92

0,93

0,94

0,95

10 20 30 40 50

Eth

en

e G

ain

Eth

en

e C

on

vers

ion

Residence time (s)

Ethene Conversion

Ethene Gain

Final Comments

The pseudo-homogeneous model hypothesis is valid and it was possible to

conclude that the diffusion process does not pose as a limiting phase for this

system.

The optimized values for the residence time tend to be closed to 24 s for systems

with recycle. Non-converted ethene may be reprocessed.

Variations in bed heights have effect on reaction gains up to 9 m; above this

value, investments in catalyst quantity increase become economically


value, investments in catalyst quantity increase become economically

impracticable as a way to increase ethene conversion.

Develop a simulation applying multi-dimension model (CFX, Fluent, H):

Varying the distributor design to control bubbles diameters by limiting their size

and thus, enlarging the feasible area for mass and energy transfer between

phases. Besides, the larger reagents volume, confined inside the bubbles, favors

the passage of gas through the bed without its contact with the catalyst particles.

Some accessories such as baffles can be used for reduction of bubbles diameter.

Analyzer effect of PSD (multi-phase).

.

oxychlorination fluid bed reactor modelling...

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