Industrial Catalytic Reactor Simulation from Lab-Scale ... · Industrial Catalytic Reactor ......

Laboratory for Chemical Technology, Ghent University

http://www.lct.UGent.be

Industrial Catalytic Reactor Simulation from Lab-Scale Intrinsic

Kinetics

Joris W. Thybaut

Methusalem Advisory Board Meeting, Ghent, Belgium, 19 June 2014

chemical kinetics in reaction engineering

from atom (nm) to full process (m)

intrinsic vs ‘industrial’ catalytic kinetics

bulk catalyst

CRs CRi

CPs CPi

CPb = CPs = CPi

CRb = CRs = CRi

intrinsic kinetics measurement

• well defined flow pattern

• isothermicity• no or well controlled

concentrationgradients

• gas/three-phaseconditions

• 1 to 10g of catalyst• high p and T

methodology

– validation of data available from industrial / pilot plant studies

– operatingconditions

– reaction pathway– detailed kinetics – model

construction– evaluation

estimated parameters

– phase effects– formation of

new compounds– solvent

adsorption– assessment of

liquid phase non ideality

extended set of gas phase

experimentation

limited liquid phase

experimentation

three-phase industrial reactor

simulation

Berty Reactor

Robinson Mahoney Reactor

industrial trickle bed

reactor

outline

• introduction• gas-to-liquid phase kinetics

– methodology/concepts– pyridine HDN

• multiphase catalytic reaction engineering– isomerization synergy in hydroconversion– complex mixture and phase effects

• multi-scale reactor modeling– ethylene oligomerization

• conclusions6

gas versus liquid phase thermodynamics

• ideal gas phase– identical molecular interactions– partial pressures as activity measure

• non ideal liquid phase– molecule dependent interactions– liquid hydrocarbon mixtures with

dissolved light gases– Peng-Robinson equation of state– fugacties as activity measure

�� ,

��

rate equations

• gas phase

• liquid phase– substitution of pi by fi according to

– phase effects in the adsorption step– gas-liquid equilibrium 8

��→�� ,��∗� ��

�"#$�

��%&��

��

��→�' � ��%&��∗�

� � 1 (�� ( ��%&�� ( ��)��)

� � �� ( ��*��*

∗ � ∗,�+� / � ( �*�-

*�-�.

��

��/ ��

0 1 � 1, …, 3

4� � 4∗�� 4∗��

4� � 4∗��

gas and liquid phase experimentation

gas phasea liquid phaseb

Reactor type Berty type (CSTR) Robinson Mahoney (CSTR)Temperature range (K) 573 – 633 543-613Pressure (MPa) 1.5 – 4.0 6.0 – 8.0

H2/pyridine (mol/mol) 80 – 600 10 – 15Space time (kgcat.s/mmol) 0.36 – 1.8 0.65 – 3

Solvent/pyridine (mol/mol) 40 20 – 40

Solvent n-hexane Halpasol© C11-C13 n-alkanes

a R. Pille and G. Froment, Stud. Surf. Sci. Catal. 106 (1997) 403-413.b C. Raghuveer et al. Fuel 125 (2014) 206-218

NH2 SH

-H2S+H2

+H2-NH3-H2S+H2

gas phase reaction network

additional reactions at

liquid phase conditions

gas and liquid phase parameters

Parameter A a / Ka Ea / -∆H

kp⇄pp 33 ± 7 41 ± 10

kpp→pa 4148 ± 802 185 ± 19

kratio 1x10-3 d 100 ± 46

kPentylPP->PP+C5 1x104 d 67.8 ± 16.2

Kp 61 ± 15 110b

Ksolvent 0.02c 39.8 ± 3.0

K1-pentylpiperidine 135c 81.8 ± 3.64

Kpp 85c 117 ± 3

KNH3 30c 76 ± 2

KH2 0.2c 104 ± 3

KH2S 21c 129 ±3Units : k [mol/hr-kgcat], Ea [kJ/mol), K [bar-1] and ∆H [kJ/mol]a at average temperature, i.e., 599 Kb Calculated, assuming similar loss of mobility as piperidinec Calculated from estimated adsorption enthalpy and best fitting adsorption entropy.d (s-1 bar -1 ) Value of pre-exponential factor (Dumesic 1993, The microkinetics of heterogenous catalysis)

Component ∆Sadsorption(J/mol-K)

ammonia -98

solvent -98

1-pentylpiperidine -130

piperidine -158

hydrogen -187

hydrogen sulphide

Decreasing mobilityon catalyst surface

gas and liquid phase simulation

Symbols: experimentally obtained values, ♦ : Pyridine, ■ : Piperidine, •: 1-Pentylpiperidine,▲: C5 hydrocarbons

▬▬ : Pyridine, ▬▬ : Piperidine, -- : 1-Pentylpiperidine, ••• : C5 hydrocarbons

550 575 600 625 650

Temperature (K)

560 570 580 590 600

Temperature (K)

1 1,5 2 2,5 3 3,5

Total Pressure (MPa)

1,5 2 2,5 3 3,5 4

Total Pressure (MPa)

outline

• conclusions12

hydroisomerization yields

isomerization

Consecutive reaction mechanism1. physisorption2. dehydrogenation3. protonation4. isomerization/cracking

max isomer. yield ≈ 60% in idealn-decane hydroconversion

0 20 40 60 80 100

n-decane conversion (%)

C10 isomers

Cracking products

0 20 40 60 80 100

C10 isomers

Cracking products

Unidirectional medium-pore zeolites:� diffusion of branched species in

micropores restricted� reaction nearly exclusively at pore

mouths USY

ZSM-22

synergy effect between ZSM22 and Y

physical mixtures of ZSM22 and Y with tailored activity → enhanced isomerization yield in gas phase model component conversion

reaction pathway1. monobranching on ZSM222. dibranching on Y

while avoiding excessive cracking on Y→ Y should only be mildly active

tridimononormal

ZSM22 + Y

tridimononormal

60 70 80 90 100

total nC10 conversion (%)

gas phase, feed mixture

component Molar composition (%)

n-decane 17

n-undecane 43

n-dodecane 28

n-tridecane 12

0 20 40 60 80 100

total parapur conversion (%)

commercial n-alkane mixture ‘parapur’

preferential physisorption and reactionof heaviest compounds

40 50 60 70 80 90 100

total parapur conversion (%)

ZSM-22

liquid phase, model compound

Attenuated synergy effect

micropore saturation at liquid phaseconditions→ ZSM22 micropores saturated with

linear species→ enhanced cracking at the expense of

isomerization50

0 20 40 60 80

gas phase

liquid phase

gas ↔ liquid: ZSM22 micropores

60 70 80 90 100

total nC10 conversion (%)

liquid phase, feed mixture

0 20 40 60 80 100

conversion (%)

synergy preserved for high ZSM-22 content

more pronounced physisorptioncompetition

catalyst mixture compositionoptimization viaSEMK model simulation

60 70 80 90 100

total papapur conversion (%)

outline

• conclusions18

multi-scale reactor model

REACTOR SCALEmass balance:

energy balance:

impulse balance:

compii niR

dF...1==

iiifBpfs TT

dTCu −−∆= ∑

dp 2ρ=−

0at0 == WTT

0at0 == Wpp

10, <≤→> ϕipi Vp

( ) li

gii RWERWER +−= 1

1...)(Re0 ≤=≤ lfWE

PELLET DIAMETER

liquid formation:CRYSTALLITE SCALEintra particle transfer:

∂∂+

∂∂

∂∂+

∂∂−=

∂∂

ξξθ

ξθ iiii

∂∂+

∂∂

∂∂+

∂∂−∆=

∂∂

ξξξλ

ξTTTs

==∂∂

ξξθ

ξθθ

for all t, except t = 0 for t = 0

≠===

ξθξθθ

==∂∂

TT sfor all t, except t = 0 for t = 0

NANO SCALE

−=rxnn

jjjii rR

( )...,,,, DTyxpfRi =

intrinsic kinetics:

industrial reactor design

0 5 10 15 20 25

Conversion (%)

Butene

Hexene

Ni+ Ni

C2H4C2H4

C2H4C4H8

termination

insertion

chemisorptionchemisorptioninitiation

SEMK simulation net production ratesdesign criteria

100 kT/y CH4 and 30% yield to C2H4→ 30 kT/y C2H4 = 37.2 mol s-1

design variables� dimensions (Lr, Lr/dr)� heating regime (isothermal, adiabatic, heat exchanging, …)� number of catalytic beds� inlet conditions (T0

max = 573 K, p0max = 3.5 MPa, …)

� catalyst properties (size …)

catalyst: Ni-Si/Al (no strong acid sites)absence of acid catalysis

mainly dimerizationreaction conditions independent selectivity

→ maximization of XC2 (and YC4)

0 1 2 3 4 5 6

Wcat (ton)

0 1 2 3 4 5 6

Wcat (ton)

reactor scale: heating regime

1. difference between single bed heat exchanging and multibed is negligible

2. economic analysis needed (cfr. cooling utilities)

30 kT/y C2H4 | p0C2H4 = 3.5 MPa | T0 = 573 K

single vs. multi fixed bed operation:

single bed, adiabaticsingle bed, heat exchangingmulti bed (3), adiabatic

0 1 2 3 4 5 6

Wcat (ton)

0.0001 0.001 0.01

dp (m)

30 kT/y C2H4 | p0C2H4 = 3.5 MPa | T0 = 573 K

single fixed bed operation:

heat exchanging (573 K)

Lr/dr = 5 – 10 – 15

Lr/dr↓ → heat exchanging surface area↓→ higher temperature and later occurrence of hot spot in the reactor

Lr/dr > 10: ∆T < 20 K

∆p is negligible(<<1%)

reactor scale: dimensions and pressure drop

Lr/dr↓

particle diameter

0 2 4 6 8 10

Wcat (kton)

0 2 4 6 8 10

Wcat (kton)

pellet scale: liquid formation

• liquid formation can enhance ethene conversion

• future prospective: including acid catalysis

→ liquids to branched olefines

30 kT/y C2H4 | p0C2H4 = 3.5 MPa | T0 = 303 K

single fixed bed operation, isothermal� C4+ (liquids) are inert on Ni(+)

� micropores are filled with liquids� ‘decelerating’ and ‘accelerating’ effect→ reduced surface slows the rate down→ product condensation keeps ethylene

partial pressure high � axial gas phase composition: ± constant�RC2: ± constant at available sites

no liquids cat. 1 cat.2 Amacro > Amicro Amacro < Amicro

better performance worse performance

crystallite scale: mass transfer limitations

under current conditions (D/Lc²>103 s-1):no mass transfer limitations

future prospect: heavy componentse.g.: D/Lc² = 0.1 ms-1

effect of shape factor of the catalyst particles = 2 (sphere) → η = 75%s = 1 (cylinder) → η = 62%s = 0 (slab) → η = 39%

-1.0 -0.8 -0.6 -0.4 -0.2 0.0 0.2 0.4 0.6 0.8 1.0

ξ (-)

Burghardt et al. Chem. Eng. Proc. 35 (1996) 65-74

0 0.5 1 1.5 2 2.5 3 3.5

Wcat (ton)

ultimate industrial reactor design

30 kT/y C2H4 (≈37.2 molC2H4 s-1)95% conversion of C2H4

p0C2H4 = 3.5 MPa

T0 = 573 Kmulti (3) fixed bed adiabatic operation

Lr = 8.1 mdr = 0.81 m mcat = 3.75 ton

no liquid formationnegligible pressure dropno intraparticle mass and

heat transfer limitations

conclusions

• starting from ‘intrinsic’ catalytic kinetics– phase effects where accounted for– adequate thermodynamics description

• multiphase reaction engineering– optimization of synergy in isomerization yield

• multi-scale industrial reactor modeling– ethylene oligomerization– reactor, pellet and crystallite scale

• variety of model reactions– HDN, hydroisomerization, ethylene

oligomerization

future work

fossil

resources

carbon

hydrogen

biomass

carbon

hydrogen

oxygen

innovative catalyst design for large-scale sustainable processes (i-CaD)

acknowledgements

prof. Guy B. MarinBart VandegehuchteKenneth TochChetan Raghuveer

Industrial Catalytic Reactor Simulation from Lab-Scale ... · Industrial Catalytic Reactor ......

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