Modelling and comparison of gas phase and liquid phase methanol synthesis reactors

(1)

•

• i!~~'(

'

T

U Delft

Technische Universiteit Delft

FVONr.

Preliminary Process Design

Section Chemical Process Technology

Subject

Modelling and Comparison of

Gas Phase and Liquid Phase

Methanol Synthesis Reactors

Authors

EH. Geertman

A.J.

Hoekstra

Date assignment

Date report

Telephone

(015) 123297

(015) 130813

February 1993

18 November 1994

Faculteit der Scheikundige Technologie en der Materiaalkunde

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•

• Modelling and Comparison of Gas Phase and Liquid Phase

Methanol Synthesis Reactors

Fabrieksvoorontwerp or 3047

F.H. Geertman

(253222)

Al.

Hoekstra (345241)

Coaching: DL

Ir.

A Cybulski

Delft Univ

.

ofTechn

.

Fac

.

Chem. Eng. and Mat. Sc.

November 1994

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•

Abstract

Abstract

Methanol is one of the basic chemicals which is manufactured at an annual rate of over ten million tons. Until now, methanol has been produced in two-phase systems: the reactants (CO, CO2 en H2) and produets (mainly CH30H and H20) forming the gas phase and the catalyst being the solid phase. An adiabatic reactor contains a multi-stage bed of a catalyst with quenching by cold gas injections (the ICI system). Recently, the use of a novel system has been reported: the Liquid Phase methanol synthesis process (LPMeOH), which is characterized by nearly isothermal reactor operation and a high conversion per pass of reactants through the reactor.

Kinetic studies extracted from literature have been used to model each reactor and calculate reactor characteristics. For both kind of reactors CO-rich gas has been used as the feed stream.

The results on the slurry reactor were in accordance to literature data. For the fixed bed

configuration the conversion per pass gave a poor result, due to a syngas composition that is not ideally suited for the catalyst. Conclusions were drawn on the basis of the simulations done in Mathematica and ChemCad for both reactor systems. For the bubble column slurry reactor the conversion per pass was 19.3 %. The conversion for the fixed bed reactor amounted to 1 %.

FOT the economic comparsion between the two reactor configurations, the methanol production has been integrated in a Coal Gasification Combined Cycle / Once Through Methanol

(IGCC/OTM) plant. In this sytem the unconverted fuel gas is combusted in a combustion turbine to generate power. Two economie models have been used to calculate for the fixed capital and operating costs. Power generation by the coal gasification plant has not been taken into account. Conclusions were drawn on the basis of the estimations done fOT the economie comparison. The fixed capital and operating costs for the liquid phase process amounted to 36.9 and 25.6 million US$ respectively. F or the fixed bed configuration these annual costs were 216 and 170 million US$ respectively.

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Con/ents I Il Il.I 11.2 Il.3 U.4 m m.l m.2 III.3 m.4

IV

IV. 1

IV.2

IV.3

IV4

IV.5

IV.6

App I . App Il . App III . App IV. AppV . App VI . AppVll Appvm App IX. App.X. App. XI App. XII

Cooteots

Introduction

Bubble Column Slurry Reactor .

Introduction

Bubble Column Reactor Kinetics

ModelDescription Bubble Column Reactor Results Bubble Column Reactor Simulation

Fixed Bed Reactor Introduction

Fixed Bed Reactor Kinetics

ModelDescription Fixed Bed Reactor Results Fixed Bed Reactor Simulation

Economics Introduction

Benefits of Coupling OlM to IGCC

Starting Points for the Economic Comparison Estimation ofthe Fixed Capital Costs

Estimation ofthe Operating Costs (Wessel) Economic Conclusions .

Schematic of the OlM / CGCC integration Calculation of the Heat Capacities

Mathematica Input Bubble Column Slurry Reactor Bubble Column Slurry Reactor Profiles

Flowsheet and Stream Composition Liquid Phase System Mathematica Input Fixed Bed Reactor

Fixed Bed Reactor Profiles

Flowsbeet and Stream Composition Gas Phase System Literature

Schematic ofthe Texaco Coal Gasification Process Economic Evaluation Bubble Column Slurry System Economic Evaluation Fixed Bed System

1 ,., -' ,., -' 5

7

12

14

15

17

21

24

25

26

27

29

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Introduction I

Introduction

The aim of this FabrieksVoorOntwerp is to compare two reactor types for methanol production. The comparison wiIl be made on the basis of a mathematical simulation of the performance of these reactors. In order for the comparison to make sense, the reactors wiIl both operate on the same feed and will be required to produce 100,000 tpa of methanol. The two reactor types will be part of an integrated once-through methanol (OTM) plant and coal gasification combined cycle (CGCC) plant. The focus wiIl be on the numeri cal solution of the equations describing the behaviour of the reactor systems. The ancillary equipment and the CGCC system will be treated briefly and receive attention to the extent where it is relevant for the calculation of the system economics.

1.1 Methanol

Methanol is one of the basic chemicals and it is manufactured at an annual rate of over 10 million tons [1]. Plant capacity for methanol production is still increasing. One ofthe major future uses of methanol will be the use as a peaking fuel in coal gasmcation combined cycle power stations. The strategy here is to produce methanol from the gas in periods of low power demand, and buring it during periods of peak power demand in combined cycle turbines. Traditionally, methanol has been produced by catalytic hydrogenation of carbon monoxide. Under synthesis conditions, an important side reaction takes place, namely the hydrogenation of carbon dioxide: the water-gas shift reaction.

CO+2H2~CH30H

CO2

+H2~CO+H20

The conversion in the fust reaction is limited by thermodynamic equilibrium. Since the reaction is exothermic and involves a contraction in volume, the yield can be increased by working at high pressures and low temperatures.

1.2 Methanol Reactors

Several types of catalyst have been developed to get areasonabie reaction rate. Zinc chromium oxide was one ofthe earliest used, later followed by mixed zinc/copper oxide catalysts

(CulZnO/Ah03 and CulZnO/Cr203). In the late sixties ICI introduced the low pressure methanol process, taking advantage ofthe development ofa very active catalyst. Typical operating conditions are 5 to 10 MPa, and 220 to 280

oe.

This process can be advantageously combined with production of syngas by partial oxidation, since the latter can be carried out at methanol synthesis pressure. In this way intermediate gas compression can be avoided. These reactors operate adiabatically, and the temperature rise in each bed is known to be extremely high. For this reason multibed quench reactors are necessary. Another possibility is carrying out the reaction in heat transfer tubes.

Since 1975 Chem Systems has been developing another technology for methanol production, namely the liquid phase methanol (LPMeOH) process. The system uses an inert hydrocarbon to fluidize a heterogeneous catalyst bed which controls the exothermic reaction heat (the heat control is only one ofthe manY advantages the system offers). In 1976 Sherwin and Frank [2] evaluated the liquid phase methanol process against the conventional system, emphasising the reduced power requirements and superior energy recovery capabilities ofthe LPMeOH plant.

Air

Products and Chemicals started a program to prove the technical feasibility of the liquid phase system in 1981. Bubble column slurry reactors were used and several aspects of the new

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•

Introduction 2

technology were investigated [3]. Two systems were screened: the flIst being a liquid fluidized mode with relatively large catalyst pellets, the second being an entrained system with fine catalyst particles slurried in an inert liquid. The slurry mode was ultimately chosen as the operating mode due to its superior performance. The development progressed from bench scale reactors to an intermediate process development unit (PDU), into the LaPorte PDU in 1984. The last system has been used to provide operating experience for a large scale process,

and it produces some 2500 tpa of methanol [3].

Air Products feel that the LPMeOH process can he considered a mature technology, offering a number of advantages listed further on in this paper.

1.3 Feed gas

Basically three feed gas compositions can be considered when designing a methanol process. They are listed in table I. The fust is CO-rich gas, which is not stoichiometrically balanced.

This mixture is representative of synthesis gas coming directly from a modem coal gasifier. This gas is typical for once-through methanol synthesis in a CGCC plant, configured to coproduce electric power and methanol [3,4]. The second (balanced) type is CO-rich gas, which has undergone shift and O2 rejection. In this way, the hydrogen and carbon monoxide concentrations are approximately stoichiometrically balanced. The ratio (2: 1) allows for conversion into an "all methanol" product. The last type is H2-rich type, which is

representative for the effluent of a steam methane reformer. This, too, is not stoichiometrica1ly balanced.

Table t Feed gas compositions

gas H₂(mol %) CO (mol %) CO₂(mol %) Inerts (mol %) I type

CO-rich 35.0 51.0 13.0 1.0

balanced 55.0 19.0 5.0 21.0

Hz-rich 71.0 18.0 7.0 4.0

Since the system presented here is an integrated methanol and CG CC system, the calculation was made for CO-rich gas.

1.4

. /

Coal Gasification Combined Cycle

A CGCC power plant converts coal into a clean gaseous fuel which is consumed in a combustion turbine. Steam, produced as a by-product of synthesis gas formation and utilization, is consumed in a steam turbine to generate additional power.

By feeding clean synthesis gas to a methanol converter, part ofthe energy in the gas becomes a storable liquid fuel. This can be sold as a by-product OT used to provide power during

periods of peak demand. Since the methanol catalyst is sensitive to sulfur poisonmg, sulfur removal must be increased from 95 % (normally specified in CGCC designs in the US) to 99.9 %. This results in an environmental benefit associated with OTM. A schematic of an integrated methanol and CGCC system can be found in Appendix 1.

The nature of a LPMeOH reactor makes it very suitable for enhancing CGCC flexibility. The amount of active slurry can be readily adjusted on-line, using pumps and storage, to vary production according to the anticipated power plant load [4].

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Reactor

Gas

Outlet---Normal

Liquid

Level

---

...

Utility

-::::::===u

0

il

Inlet

•

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•

The Bubble Column Slurry Reactor

11 The Bubble Column Slurry Reactor

U.l

Introduction

The synthesis of methanol from CO and H2 bas conventionally been carried out in a fixed-bed reactor in the gas phase. More recently, liquid phase methanol synthesis in a slurry reactor has received considerabie attention, owing to the advantages that this system offers. Commercial processes employed today effect the reaction of synthesis gas (CO + H2) over a bed of

heterogeneous catalyst arranged in either sequential adiabatic beds or placed within heat tranfer tubes.

3

A so-called three-phase methanol reactor operates in a different manner: The working principle of the system is as follows: synthesis gas is blown upward through a dispersion of catalyst particles in a liquid of inert hydrocarbons (the slurry). It enters the reactor through a gas sparger, thus providing the driving force for mixing and suspension of the catalyst particles. The reactor is fitted with an upper freeboard zone wbere the slurry disengages from the unconverted synthesis gas and the methanol vapor.

The inert liquid has a two-fold function. The fust is to keep the fine catalyst particles dispersed, while the second is allowing for an easy temperature control. T 0 achieve this, the reactor is fitted

with an internal heat exchanger, where utility oil is circulated in order to keep the temperature at the preset level. This heat exchanger consists of parallel tubes placed inside the reactor. In the pilot set-up that bas served as a basis for the model in this report the tubes make up for as little as 3.S % ofthe reactor cross-sectional area [3]. Therefore, it does not significantly impact the reactor hydrodynarnics.

The liquid phase process modelled in this report pertains to the system as it is being developed by Air Products and Chemicals, in the Process Development Unit (PDU) at LaPorte (USA). Chem Systems has been the key subcontractor in the program [3]. Figure 1 shows the reactor used in the test program.

The liquid phase methanol process has, in principle, a number of advantages over the gas phase process. However, some caution should be observed in comparing the two systems, since experience with the former is limited to pilot-scale set-ups.

The advantages are listed below:

V

1. Due to excellent reaction temperature control, high conversions of syngas per pass can be economically realized: this means a great reduction of recycle gas flow and compression requirements. In the OTM (once through methanol) process discussed in this paper, this is

V

s

.

clearly less significant.

r

e

V\.1 () Vil

ol

~

The heat of reaction can be largel)(feco{erIDn a very simple manner. One can either use an internal heat exchanger [3], or one can pump the slurry through an externalloop with a beat exchanger.

Reactor design is simplified in that liquids and gases are readily distributed across the reactor cross-sectional area, without the necessity for redistribution and quench along the reactor length.

A small size of catalyst particles may be used, thereby achieving higher rates of reaction than with the larger catalyst particles used in the fixed bed design.

The catalyst can be added to the system and withdrawn without the necessity of system shutdown. In this way, the reactor is continuously operated at the economicallevel of catalyst activity.

The near isothermal temperature of operation throughout the reactor permits optimum conditions favouring the desired reaction kinetics.

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The advantages of a fluidized bed system are present, without the problem of attrition resulting in catalyst loss. The oil functions as a lubricant between the catalyst particles. Since the catalyst is at equilibrium 0 erating level of activity, it is not necessary to

overdesign the reactor size r end-of-life catalyst activity level.

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The Bubble Column Slurry Reactor 5

- -

0.2 Bubble Column Reactor Kinetics

Very little is known about the reaction mechanism of the liquid phase methanol synthesis.

Moreover, the kinetics ofthe gas ph ase process are not fuUy understood and the models presented by different research groups are often conflicting. Therefore, it would be unwise to adapt the same equations for the slurry bed calculation, as will be used for the flxed bed. Moreover, practice has shown that the selectivity ofthe slurry bed reaction is much higher, compared to the flxed bed.

Öztürk [5] takes into account both the water-gas shift reaction and the methanol synthesis reaction. The reactions occur according to the following stoichiometry:

c

o

+

2 H₂- - - - ) CH₃0H CO₂

+

H₂- - - - ) CO

+

H₂0

(methanol synthesis)

(shift reaction)

Practice at LaPorte has shown, though, that in liquid phase methanol synthesis the so-called shift reaction hardly occurs, making the kinetics ofthe system somewhat easier to describe [3]. Out ofthe various models describing the reaction kinetics the foUowing has been chosen. The choice was motivated by the fact that this equation was derived at the LaPorte LPMeOH reactor. The model presented in this report therefore uses the equations derived from a pilot system c10sely resembling the reactor described by the model.

K D1I3F,

[1

PMeOH ] ,",'vIeOH

=

f r co H2 - 2

KeqPCOPH2 where

the methanol productivity (mol / s / kg cat) the rate constant (mol / s / kg.,., / atm) the methanol equilibrium constant (atm-2)

the partial pressure of component i (atm)

(1)

The rate constant Kr is a function of temperature, as is deftned by the Arrhenius equation. In

addition, it is a function of time. The catalyst deactivates at a certain rate, which is described by a time-dependent expression. Air Products has found this expression to be an exponential decay function ofthe number of days from catalyst introduction. The reaction rate constant has been determined by the same company using a two CSTR in series model of the LPMeOH reactor [3].

-Eact Kf=Ko(t)·eRT K (t)

=

K . .. . (-0.0039480) o O.mmal

e

where Eact D K _O_,_IDlU. ··al (2,3)

the energy of activation (J / mol)

the time elapsed from catalyst introduction (days) the rate constant for a fresh catalyst (mol / s / kg.,.. / atm)

One ofthe advantages ofusing a slurry bed reactor is the fact that the reactor does not have to be shut down for catalyst replacement. The catalyst can be refreshed almost continuously, thereby keeping the activity and the catalyst use at an economical optimum. Air Products reports that a

(14)

•

• :

I

.

I

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catalyst residence time of 60 days is ideal in this respect [3]. The time dependence of the rate

constant can thus be taken out of the equations by assuming an average activity, equal to that Of0

-60 days old

ca~~

~J

l ,

) \

The rate

consti:U~ecomes

P)'

i

IA.

c

i

P ,.

'l 0(,.

~

v+

.

I

~

r

t

ttJv

\5

(h

v~

' j

we

'f~

.

Ko(t)

=

_KO

,

initial· e

(-OOO394S60)

=

KO

,

initi

a

l

· O

.

789

(4) The values supplied by Air Products [3] for Ko,initial and Eact are

Ko,initial Eact

657.8 mol / s / kgc.1! / atm 57520 J / mol

The equilibrium constant Keq is of course temperature-dependent and is given by

log\O(K

_e_q)

=

5139 -12.621

(Tin

Kelvin)

T

(5)

The reaction lànetics described here will be used in the differential equation, which will be derived in the following paragraph.

(j Do

J

<k>

1

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(J

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ou

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•

D.3 Model Description Bubble Column Reactor

The method used to compute the product flow coming out of the reactor is integrating a

differential mass balance over the reactor length. Since three phases are present, a model has to be developed comprising terms for production and consumption, as weIl as mass transport between the three phases. Build-up terms are superfluous since a steady state situation is assumed.

The reaction occurs within the catalyst particles (the solid phase). There is a net flow ofproducts out ofthe solid phase, while the reactants go in. Before the gas flow is able to carry the product to the reactor outlet, the liquid phase has to be passed. This means that both solid-liquid and gas-liquid mass transport is relevant.

A schematic ofthe model used is depicted in figure TI.I, showing the Oavidson-Harrison model.

Fig. Il.I The Oavidson-Harrison model

The gas moves upwards in plug flow through the liquid phase in the reactor. This assumption is widely accepted in modeling tall and narrow bubble column slurry reactors, where the Bodenstein number is usually high [5]. In this case Bodenstein number is given by the following equation

Where

the superficial gas velocity (m / s) the reactor height (m)

the gas holdup

the gas phase dispersion coefficient (m2

/ s)

(6)

Both the liquid and the solid phase are considered ideally mixed. The liquid phase is mixed by the bubble column, the solid phase consists of very small particles where diffusional resistances can be neglected [5].

The abovementioned assumptions are important in assessing the validity of the chosen model. A number of additional assumptions follow. They are listed here, in order to prevent scattered remarks throughout the paragraph. This will help to keep the reasoning clear when deriving the differential equation describing the system.

V

2.

Pressure: the total pressure within the reactor is considered constant. This means that the influence of the liquid head on the gas expansion is neglected. Oue to the relatively high pressure used, this assumption is reasonable.

Isothermicity: there is no heat transfer to the surroundings, except for the heat transfer through the heat exchanger. Furthermore, the reactor is isothermal. This approximation is

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•

• J

]he Bubble Column Slurry Reactor 8

7.

8.

9.

10.

not far from what has been achieved by Air Products [3]. There is no temperature gradient in the catalyst particles, and the solid to liquid heat transfer is considered instantaneous. Heat exchange: the heat exchanger only absorbs the heat of reaction and does not provoke temperature gradients in the reactor.

Mixing: the liquid is considered ideally mixed by the bubble column. This is also the case for the solid phase: the catalyst is even1y distributed over the reactor. Although the catalyst has a higher density than the liquid, sedimenta~on of the catalyst particIes due to gravity is neglected. - - i )

h

~

'-V CLvll €.. 'v ll.-

cA .

Gas flow: the variation of gas flow rates over the column is neglected. For the flow pattern in the gas phase, a plug flow madel is used. As was already stated, this corresponds to a general case in bubble column slurry reactors where the gas phase Bodenstein number is usually high.

Diffusion: the effectiveness factor for the pore diffusion inside the catalyst particIes is taken as unity. The diffusionallimitations can safely be considered negligible because of the small particIe size [5].

Mass transfer: All the mass transfer resistances from one phase to another are considered negligible.

Selectivity: The water-gas shift reaction is not taken into account. Like many ofthe other assumptions this is validated by the Air Products experience.

Properties: densities and surface tension are considered temperature-independent, while viscosity is calculated for the reactor temperature. Gas hold up, interfacial area, and the heat transfer coeflicient are considered spatially independent. Although the interfacial area does not appear in the equations, a variation along the reactor axis would make it

necessary to consider C1>ls and C1>gl as functions ofthe reactor axial coordinate.

m '\

I

Densities are considered independent oftemperature. _____ 1 l . The equation describing the reaction kinetics uses partial pressures. The ideal gas law is

considered applicable in order to calculate the concentrations of the various components.

I

A mass balance needs to be solved for only one of the components, since the stoichiometry determines the concentrations of the other participants in the reaction. Any of the substances may therefore be chosen, in this case CO is arbitrarily taken as the key component.

At each point in the reactor, the concentrations of the other reacting components (hydrogen and methanol) can be calculated using the following equations:

_ ( \1eOH ( X )

=

Ceo

(0) -

Ceo ( x )

\>~{Ol.V:>

:

~

-o~O~~

CH2(x) = CH2(O)-

2·

[Ceo(O)- Ceo(x)]

'11\v1~

i v-.

.2..-'1';

vC(.~

· r~\"~'\~IA."'~

- S CJ \ The rate of consumptlOn of CO in the catalyst particIe (the solid phase) is given by

where

rMeOH

=

C_cat

I-Eg

veo

reaction rate according to the presented kinetics (mol / k8cat / s) catalyst concentration (k8cat / m\quid)

liquid holdup (m\quid / m3101a1) stoichiometric coefficient of CO (-)

(7,8)

(9)

Since a steady state situation exists, the fate of CO consumption within the catalyst particIes (the solid phase) must equal the mass transfer of CO from the liquid phase into the solid phase, denoted with the symbol <DIs'

(17)

•

Therefore,

(9) where

<DIs liquid-to-solid mass transfer of CO (mol / s / m3)

In the gas phase, a differential mass balance is used. In this mass balance over an infmitesimal control volume with a height dx, three terms have to be taken into account (again, assuming steady state conditions): the CO that comes in at a height x, the CO that goes out at x + dx, and the tranfer term from the gas phase into the liquid phase.

or

where

gas-to-liquid mass transfer of CO (mol / s / m3 ) superficial gas velocity (m / s)

(10)

(11)

concentration of component Y in the gas phase at height x (mol / m3 )

No hold-up terms need to be defined since a steady state situation is assumed, thus eliminating a time-dependency of the equations. Moreover, the reaction occurs solely in the solid phase. Therefore, the mass balance over the liquid phase is limited to the two terms <1>Is and <1>gl. The equation for the liquid phase is thus reduced to

<l>

Is

=

<l>

gl (12)

Combining equations 9, 11 and 12 yields the following differential equation, used to describe the system mathematically in the computer program.

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As is said, one of the advantages of the three-phase methanol process is the excellent temperature con trol. Heat is continuously generated by the reaction, and is taken out of the system by the reactor intemal heat exchanger. The mean temperature difference can be viewed as the driving force ofthis heat transfer. Since the reactor in this model is operated isothermaIly, the driving

force is given by the following equation: •

No

,

T

-T

~

T

= m.c ou/.c am

2

where

T-h

Tin,c Tout,c

coolant inIet temperature (K) coolant outlet temperature (K)

Jlw,

woulJ. {,~

(14)

i

L..kpe",cLe.J.

1 ~

L

~ r~c(.(lo

(18)

«-•

•

The heat balance, then, is represented by

(15) where

the heat exchanger's overall heat transfer coefficient (W / K / m2 ) the heat exchanger's surface (m2

) the enthalpy ofreaction (1 / mol)

The enthalpy of reaction is taken as temperature-dependent, using the following expression

where

the enthalpy ofreaction at 298

K

(J / mol) the heat capacity change ofreaction (J / mol / K)

the calculation of ~Cp is elucidated in appendix IT.

(16) 10

The catalyst concentration appears in many ofthe equations above, and some additional

computing is needed to obtain its value. The only parameter which is given by Air Products is the catalyst loading (expressed in wt. % ) which gives the best results in terms of conversion, aging, heat transfer etc. The catalyst concentration is then calculated using the following equations:

CCal

=

w"al . P avge

P avge

=

VcarPcal

+

(1 -

_{VCal )PI}

V

=

P avgew"al

cal

Pcat - w"al (Pcal - PI )

where W_cat mcat msl PI Pcat

IDcat / msl (as taken from Air Products [3])

the catalyst weight (kg) the sluny weight (kg) the liquid density (kg / m3

) the catalyst density (kg / m3

)

(17,18,19)

The only remaining parameter to be defined is the gas hold-up. It is calculated using the equation supplied by Öztürk [5]. _--=-g E --:-

=

0 2.

[2 ]0125 [ 3

gDcp,

. gDc

]0

.

833 J{f2

.

_ g_ (1-

Egt

.

0",

0 gD

_c

777.4

8.8778 .10-6.10

r-v,

=

p,

7 V

_J

-(20,21)

(19)

•

Where

the viscosity of the liquid phase (kg / m / s) the surf ace tension ofthe liquid phase (N / m) the column diameter (m)

The model described in the previous pages is typed in an ASCII file in the format Mathematica reads. First, the gas holdup (as defined in equation 21) is calculated. Subsequently, the equations describing the other parameters and the assumed values are fed into Mathematica using the GET «<) commando The differential equation is typed and the command NDSolve calculates the concentration profiles in the reactor. The actual input files and commands are listed in Appendix

III.

The heat flow and the mean temperature difference are calculated arithmetically, and need not be part ofthe simulation. Once the conversion per pass is known, the required heat flow through the internal heat exchanger is easily calculated.

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•

D.4 Results Bubble Column Reactor Simulation

The model described in the previous chapter was fed into the mathematical analysis program (Mathematica, by Wolfram Research). Appendix III contains the actual input needed to calculate the system. Mathematica uses a Runge Kutta procedure to solve the system of equations

numerically.

12

Mathematica generates tables of values for the dependent variables, which can subsequently be used to produce the graphs featuring in the figures of Appendix IV. The graphs show the rather high conversions that can be expected from the liquid phase methanol process, which have also been experienced at LaPorte [3]. In order to make a sensible comparison between the gas phase and the liquid phase process, the production bas been set at 100.000 tpa of methanol for both, in a single pass process. The feed gas composition is ofthe CO-rich type. Generally, the process parameters have been chosen at the value where they have provided satisfactory results at LaPorte. The catalyst loading for instance has been set at 0.35, which is the highest concentration at which the catalyst gave 100 % perfonnance in the POU, relative to autoc1ave tests [3]. The operating temperature was chosen at 480 K, which is within the range studied at LaPorte, but on the low end. This thennodynamically favours the methanol production.

Öztürk has modelled the system assuming a superficial gas velocity of 0.1 m / s [5]. This is reasonable for a real-life situation, since the boundaries supplied by Air Products are 0.1 ft / s (minimum) and 1 ft I s (maximum) [3]. This uantity was considered non-variable, and the required production has been obtained by solely c angmg ereactor unensIOns.

Table 1 gives an overview ofthe values used for process parameters in solving the model.

Table 1

Process parameters for the slurry bed design

parameter Symbol (in derivation) Unit Value in calculation

temperature T K 480

pressure p MPa 0.5

liquid density _Dl kg/m3 620

catalyst density _O,."t kg/m3 1980

superficial gas velocity Uo mis 0.1

catalyst loading W

-

0.35

energy of activation _E"rt kJ / mol 57520

column diameter Or. m 3.5

inlet CO concentration Cr()(O) mol/m3 639

inlet H2 concentration CJ..n(O) mol/m3 439

inlet inerts concentration Cin"'rt,,(O) mol/m3 12.5

liquid holdup _1-1::" m\quid / m3

totaJ (-) 0.815 (calc.)

overall heat transf coeff Uh .. W Im2/K 576.3

tube area int. heat exch. Ahp. ml 199 (calc.)

slurry height L m 14

liquid viscosity VI Pas 0.000370 (calc.)

liquid surface tension ()I Nim 0.016

stand. reaction enthalpy ~Hl98 kJ Imo! -94.084

(21)

•

Solving the system for the values presented above yields the results coUected in Table 2. The values for the outgoing compositions pertain to a sluny column ofvarying height. The conversion is calculated relative to the CO concentration using the following expression

(22)

where

ç(x) the conversion relative to CO at height x in the column

Table 2 Conversions at various heights in the sluny bed

height of bed COconc. H₂conc. MeOH conc. converslOn

(m) (mol / m3) (mol / m3) (mol / m3) (%)

0 639 439 0 0

12 526.520 214.040 112.480 17.60

13 520.529 202.058 118.471 18.47

14 515.955 192.910 123.045 19.26

15 511.454 183.908 127.546 19.96

Clearly, the amount ofmethanolleaving the reactor is dependent on the outlet concentration, but also on the superficial gas velocity and on the reactor cross-sectional area. A reactor diameter of 3.5 m combined with a sluny height of 14 m and a superncial gas velocity of 0.1 m / s gives a methanol production of 98.192 tpa (assuming that the reactor is operational 24 hours / day, 300 days / year).

The amount ofheat generated is calculated to be 11586 kW, by simply multiplying the methanol production by the reaction enthalpy. The overall heat transfer coefficient presented in table 1 was measured at LaPorte [3]. Proportionally increasing the heat exchanger tube area ofthe pilot set-up with the methanol production, yields a tube area of 199 m2

•

An effective tube length of 14 meter is assurned. If one uses the same material and tube diameter (2.54 cm) which gave excellent results at LaPorte [3], one can use the measured heat transfer coefficient. This ensures that the model is as realistic as possible, and leads to a mean temperature difference of 101 Kin this simulation (eq. 15).

This number of pipes will not affect the reactor hydrodynamics, since the pipes' total cross sectional area equals less than 1 % ofthe reactor's cross sectional area.

More complete data on the reactor input and output can be found in Appendix IV, which also comprises the concentration and conversion profiles along the reactor's axial coordinate. The most important conclusion that can be drawn from the data presented here and in the Appendices is that the system shows the high conversion per pass, which is experienced at LaPorte [3]. The advantages ofthe bubble column sluny reactor as far as conversion is concemed can be clearly observed in the results ofthis simulation. A conversion ofnearly 20 % is

significantly higher than what is obtained in practice using the fixed bed design.

The streams in and out of the ancillary equipment are calculated in ChemCad. The complexity of the system is kept to a minimum, since the aim is to compare two different reactor systems. The results ofthe ChemCad calculation are listed in Appendix V, as weIl as the process flowsheet.

(22)

•

The Fixed Bed Reactor 14

m

Tbe Fixed Bed Reactor

ID. I

Introduction

The fixed-bed reactor is the traditional type used to carry out methanol synthesis. It involves a so-called two-phase system. The reactants (CO, CO2 and H2) and the products (mainly H2

0

and MeOH) are in the gas phase, while the catalyst is the solid phase. Two designs

predominate in the plants that have been built since the 1970's: the fust is an adiabatic reactor containing a number ofbeds, featuring quenches with cold product injections in between the catalyst beds. The second type is an arragement of a multitubular reactor with an internal heat exchanger. These systems are known as the ICI design and the Lurgi design, respectively. A more recent development is the gaslsolid/solid trickle flow reactor, where a fine adsorbent powder flows downward through the reactor. The powder selectively adsorbs methanol as it is formed, thereby shilling the equilibrium towards the product side. A similar system is using interstage product adsorption units. The beds are now separated by the adsorbent, and higher conversion per pass can be attained [1].

The modern low pressure methanol synthesis catalysts are usually based on Cu/ZnOI Ah03 or

CU/ZnO/Cr2ÛJ. These mixed oxide catalysts have been found to be considerably more active than their individual components [1]. In the early sixties, the presence of carbon dioxide up to 6 % (by volume) was found to increase the activity significantly, when using a Cu/Zn/ Al catalyst. This has resulted in the full implementation of this catalyst by ICI in a methanol synthesis process at 50 bar. This has been the operating pressure used in the calculation presented in this paper.

A large number of publications can be found on methanol synthesis, particularlyon reaction kinetics and mechanism. Unfortunately, kinetics studies are often conflicting. Skrzypek [6] claims that methanol synthesis over copper containing catalysts occurs not from CO, but rather from the hydrogenation of CO2 • Simultaneously to this reaction, the water-gas shift reaction

takes place (or its reverse, depending on reactor conditions). Klier [7] and Öztürk [5], on the other hand, state that methanol is produced by the catalytic hydrogenation of CO. The side reaction is again the water-gas shift reaction. The conversion of CO to methanol is greatly lirnited by thermodynamic equilibrium.

In this paper the theory ofSkrzypek [4] is foilowed, as weil as the set of equations he provides, accurately describing the kinetics observed.

The aim in this paper is to compare the performance and economics ofthe conventional fixed-bed reactor (the multifixed-bed configuration with intermediate quenching) and the bubble column slurry reactor (a liquid phase methanol sysnthesis developed by Air Products). The feed must therefore be the same for both systems. In view ofthe situation under study (integration of OTM and CGCC) the most realistic feed is CO-rich gas. Clearly, this is not the ideal

composition for the fixed bed system as far as conversion per pass is concerned. Nonetheless it has been used in this calculation in order for the comparison to make sense.

@

AHer

v..<L-h" e

:

"flti

"",i~c: U~ ~.t

'1

flI)

C~F'.n;

h:"",

r

v...

Y

~cL~

r '-

'7

P<-

Ct.A-

J

Lo

'-...I

J

ct

Ve...

(23)

•

• m.2

Fixed Bed Reactor Kineties

Many research groups have reported on kinetics and mechanism of gas phase methanol synthesis on copper-containing catalysts. As was already mentioned, there is still no agreement in the literature on the kinetics of this process. Moreover, until now there has been no agreement as to the basic reactions in the systern, and it is feIt that the role of CO2 is insufficiently understood. It is generally accpted that some quantity of CO2 is necessary for the process to start and to proceed.

According to Klier [7] the promotion-effect of CO2 in the feed can be attributed to its ability to keep the catalyst in its intermediate oxidation state, rather than a direct hydrogenation of CO2 to

methanol. On a copper-based catalyst methanol would be formed from CO. This argument is also followed by Öztürk [5]. Skrzypek, however, indicates that methanol formation on

copper-containing catalysts occurs not from CO, but rather via CO2 hydrogenation [6]. Simultaneously to the methanol synthesis, the water-gas shift reaction (or the reverse water-gas shift reaction, depending on the process conditions) takes place.

It was felt that the best comparison between gas phase and liquid phase methanol synthesis could be made by using data pertaining to real-life situations. Therefore, the kinetics derived at the process developmënt unit at LaPorte were used m sunulating the bubble column slurry reactor [3].

Similarly, kinetics derived from an industrial catalyst, taken from a running methanol plant, will be used for calculating the fixed bed system, provided by Skrzypek [6]. The catalyst was prepared on an industrial scale and had been used in the industrial reactor for one year. The numerical values of the parameters are supplied in the same paper [6]. It is therefore assumed that methanol production occurs from hydrogenation of CO2 •

Thus, the two major reactions are defined as:

CO₂+3H₂ Kpl ) CH

30H

+

H2

0

CO+ H₂0 Kp, ) CO

2

+

H2

F or the desired reaction (formation of methanol from carbon monoxide) the achievable conversion is limited by thermodynamic equilibrium; as was already stated in the introduction, the reaction is exothermic and involves a contraction in volume. The highe st yields of methanol are therefore obtained at high pressure and low temperature. The exothermicity of the reaction thus makes it necessary to cool the mixture after a certain reaction time (after a certain length of bed). This is possible by injecting cold feed in between a number of catalyst beds, using the so-called multi-bed quench system.

Skrzypek performed a large series of experiments to determine the final form of the equations used here [6]. They are ofthe Langmuir-Hinshelwood type, with surface reactions between CO2 and H2 as rate-determining steps.

(24)

•

here,

K

=

Ko

.e

(~)

I I

(-E,)

kj

=

kiO. e

Kr (3,4)

the equilibrium constants used were taken from Beenackers [8]. The expressions are as follows:

(3~ -10592) ₆

_H

_:-2b

kJ

)1.-

Q )

7

Kpl

=

10 kj

j)._\-.A

7

(5,6) ( -2~73 +2.029) .A }\ ~ - I

t

K p2

=

10

The values for the activation constants, pre-exponential factors and reaction rate constants are given in Table lIl. 1.

Table

m.l

Constants used in the kinetics equations, as determined by Skrzypek [6]

Constants Value Unit

klO 3.0 E+9 kmol / k~/ hr k20 2.5 E+9 kmol I k~t I hr

EI 104.7 kJ I mol

E2 104.7 kJ I mol

KH2 KIO 0.14 E-8 atm-l

LllI l 75.4 kJ I mol

Kc02 K20 0.44 E-8

atm-l

LllI2 75.4 kJ I mol

KMeOH

K

30 0.11 E-9

atm-l

MI3 29.3 kJ I mol

KH20 ~o 0.35 E-8 atm-l

~ 75.4 kJ I mol Kco

K

50 0.50 E-1O atm-1

LllI5 75.4 kJ I mol

The kinetic model described in these pages will be part of the differential equations describing the system. These equations will be derived in the next paragraph.

(

-6

&-41.

)

R1 -;

0

10 \/

(25)

•

The Fixed Bed Reactor

ID.3

Model DescriptioD Fixed Bed Reactor

The fixed-bed reactor is simpler than the bubble column reactor in the sense that two phases are considered, while the latter contains three. The gas phase is made up by the reactants and the products, and the solid phase is made up by the catalyst. The reaction occurs within the catalyst pellets, and mass transport is govemed by diffusional limitations.

To describe these two-phase systems one generally uses a heterogeneous reactor model, where the gas phase can be considered in plugflow. The solid (catalyst) phase is considered ideally mixed.

G

Fig. ID.I Two-phase model for heterogeneous reactor

17

A number of assumptions need to be made in order to obtain a workable model describing the fixed bed system. In order to keep the derivation as clear as possible, th~ assumptions are not mentioned . J asthe reasoning goes along, but they are listed here at the beginn,t'ng of the paragraph.

V

1. Plug flow of the gasphase: the significance ofaxial dispersio ends on reactor length,

vi

2.

V

5. ~.

~.

effective diffusivity and gas velocity. For high velocities ( e> I) or a reactor w ose length to diameter ratio is weIl above unity, axial dispersion IS negligible. A plug flow

model can be applied.

No radial pro:files: since radial gradients in concentration and temperature are caused solely by radial variation in axial velocity, and since plug flow is assumed, radial profiles can be neglected in this adiabatic plug flow reactor.

Ideal gas law: fixed bed methanol synthesis operates at relatively moderate pressures and

temperatures (at 480 K and 50 atm), where the nonideaIity ofthe gas mixture can be , . C~

C"

1)

I

neglected. The ideal gas law may be applied. -

ol\

J..

'1 uv-

cl..

~L" 1 '- .... •

Reaction rates: both reactions, methanol formation and the shift reaction, proceed with finite reaction rates and neither of them is considered to be at equilibrium.

Pore diffusion limitation: For each reaction, pore diffusion limitation is taken into account by using an overall effectiveness factor according to Berty [9].

Film diffusion limitation: a semi-homogeneous reactor model was used, and differences in concentrations and temperature between the bulk gas stream and the particles were not taken into account.

Catalyst deactivation: aging of the catalyst is not taken into account since the numerical values for the kinetic parameters were derived from experiments with a very stabie catalyst which had been used in a commercial reactor for a year [6]. Differences in catalyst activity along the reactor' s axial coordinate are considered negligible.

Heat of reaction: the heats of reaction are considered to be functions only of temperature and not of pressure.

Heat capacities: these are known to be a function of composition and temperature. They were calculated for the feed inlet conditions. Since the conversion in this process is rather low, the change in composition results in only a small change ofthe gas heat capacity.

Accordingly, a mean heat capacity was used.

Pressure drop: the pressure drop over the reactor is taken into account insofar as a result of the decreasing total number of moles. At 50 bars pressure the pressure drop caused by the catalyst bed is assumed negligible. Increasing the mole flow with the quench streams does not affect the pressure. Only the gas velocities increase.

(26)

•

\

11. Parameters: The superficial gas velocity, the average heat capacity ofthe gas and the density of the gas are considered independent of the catalyst bed axial coordinate,

18

The key components in this simulation were chosen to be methanol (MeOH) and carbon monoxide (CO). Their concentrations will directly appear in the differential equation that will be derived. The other concentrations are related to those ofthe key components through the following equations

C H20 (x) =

C H20 (0)

+ [

C MeOH (x) - C MeOH (0)] - [ CCO (0) - CCO (x) ]

CH2

(x)

=

CH2

(O) -

3·

_{[CMeOH (x)- CMeOH(O)]+ [CCo(O)-}

CCo(x)]

CC02 (x)

=

CC02

(0) - [

C

MeOH (x) - C MeOH

(0)]

+ [

CCO (0) -

CCO (x) ]

A steady state exists, therefore no build-up terms need to be specified. The transfer of mass from the catalyst to the gas phase must be considered equal to the change in mass (for each component) in the reaction. Therefore,

sg,MeOH

=

Peat TMi

<l>

sg.CO

= -

_{Peat '}

h

r2

where pcat Tli ri sg, McOH sg,co

the catalyst particle density (k~, / m3 ) the effectiveness factor for reaction i (-) the reaction rate ofreaction i (mol / kgc.r / s) the solid to gass mass tranfer of methanol (mol / m3

/ s) the solid to gass mass tranfer of CO (mol / m3

/ s)

(10,11)

Since no build-up terms are defined, the mass flow out ofthe solid phase must equal the convective transport by the gas in an infmitesimal control volume. We can write

<l>

sg

,

MeOH .

d"C

=

U

g

C

MeOH

( X

+

dx) -

U

g

C

MeOH

(x)

sg

.

co

·dx

=

UgCCoe x )-

UgCCOeX

+

dx)

where

x the catalyst bed axial coordinate (m)

C(x) the concentration of component i at coordinate x (mol / m3 )

(12,13)

Although convective transport occurs solely in the gas phase, and the reaction occurs solely in the solid phase, these terms may appear in the same equation. The same argumentation was followed in

chapter II. The mass balances, then, for methanol and carbon monoxide are obtained by combining equations number 10 up to and including 13.

!

[UgCMeOH ]

=

Peat1Jl'i

; [UgCco ]

=

-Peat1J2r

2

(14,15)

An important difference with the bubble column slurry reactor is the size ofthe catalyst particles. The small catalyst particles in the slurry enabled us to consider diffusionallimitations negligible. This is not the case in the fixed bed configuration. Öztürk [5J has taken diffusional resistances into account by introducing an effectiveness factor. He has found that the effectiveness factor varies with

(27)

•

the p\ace in the reactor. In this simu\ation the effectlveness factors have been a'\sumed constant as computed by Berty [9].

Since tbe reactions occur within tbe catalyst particles, film resistances may play a roie in tbe process. However, extemal mass and energy transfer resistances were found to be negligible by Öztürk [5), and are tberefore not taken into account in this model.

19

The reactor is operated adiabatically, and high temperature diftèrences over tbe catalyst bed's axial coordinate can be expected due to tbe exotbermicity of tbe reactions. The temperature has an effect not only on tbe reaction rates (through tbe kinetic equations), but also on tbe heats ofreaction. The heats of reaction of tbe two reactions are tberefore entered into tbe simulation as functions of

temperature.

The heat balance is entirely different from tbe case of a bubble column sluny reactor, which was operated isotbennaIly. Since tbe bed is considered adiabatic, three tenns have to be taken into account in tbe infinitesimal control volume: tbe heat brought in and out by tbe gas flow, and tbe heat generated by tbe reactions.

lbis leads to tbe fol1owing heat balance

or, in differential notation,

witb where L1rH L1rHj298

pg

Cp..vgc L1Cp

tbe heat ofreaction i (l / mol)

tbe heat ofreaction i at 298 K (J I mol) tbe gas density (kg I m3

)

tbe average heat capacity (l I kgg..1 K)

tbe heat capacity change ofreaction (l / mol I K)

tbe calculation of L1Cp and Cp,avgc is elucidated in appendix I

(17)

These equations have been numerically solved in Matbematica for each catalyst bed. Since tbe ideal gas law may be applied, converting concentrations (in tbe model equations) into partial pressures (in tbe kinetics equations) is easy. The actual input files and commands used to calculate tbe profiles are listed in Appendix VI.

Quenching: After each bed calculation, tbe stream is (ideally) mixed witb tbe co\d quench stream. The result of this mixing process is tbe feed for tbe numeri cal simulation of tbe next bed, and so on, until tbe four stages have been passed. The simulation of tbe mixing in between tbe beds is

calculated using an (ideal) mixing model in ChemCad. It uses tbe UNIF AC K-value model, and its more useful output are tbe temperature and composition of tbe stream as it enters tbe next stage.

In addition to tbe reactor tbe system comprises a compressor and heat exchangers to bring tbe feed to quench temperature and pressure. Subsequently, a divider is used to split tbe stream into tbe feed and tbe three quenches. The reactor etlluent passes a heat exchanger where tbe feed stream is heated to tbe reaction temperature.

As was already mentioned, tbe primary goal is to compare tbe two reactor systems running on a particular feed gas. Therefore, tbe ancillary equipment will not be treated much in detail.

(28)

•

The equipment with the exception of the reactor itself was again simulated using ChemCad. A flow sheet of the system and the composition and temperatures of the feed and product streams as calculated can be found in Appendix VI.

(29)

•

• ID.4

Results Fixed Bed Reactor SimulatioD

The model described in the previous paragraphs was simulated combining the programs Mathematica and ChemCad. The strategy is calculating the concentration and temperature profiles in Mathematica. The height of a particular catalyst bed is determined by the maximum allowable temperature rise in the stage. The limit in view of catalyst deactivation and reaction thennodynarnics is 525 K. The point at which that temperature is reached using this feed type

and model, Is tM pomt at which the bed effluent is mixed with the quench. Accordingly, the four stages vary in height. The properties of the mixture are calculated in ChemCad, the result is subsequently fed into the next stage.

Few parameters change from one stage to another. Clearly, the initial temperature and concentrations do, but the superficial gas velocity changes, too. The amount of gas increases with every quenching step, and since the pressure is assumed to remain that of the previous bed's effluent, the velocity must increase. It varies from 0.5

mis

in the fust bed, to 0.8

mis

in the fourth. This remains within the operational values given by Öztürk for fixed bed methanol systems, namely 0.4 to 0.8

mis

[5].

The feed gas is of the CO-rich type, in order to make a sensible comparison with the other system, and the production has been set to 100.000 !pa. Since the number of stages bas been set to four, and the maximum height of a stage is deterrnined by the tempeature rise, the easiest way to regulate the production ofthis system is to change the reactor diameter. Table 1 gives an overview of the parameters used in the calculations

Table 1 Process parameters for tbe fixed bed design

parameter Syrnbol Unit Value in calculation

reactor axial coord x m variabie

initial temperature T K 480

initial pressure p atm 50

gas superf velocity Ug,l- Ug,4

mis

0.5 - 0.6 - 0.7 - 0.8

height of stage LI - L4 m 3.1 - 2.9 - 3.5 - 4.0

total reactor height L m 16

feed mole flowrate _<Dm kmol I h 43920

quench mole flowrate <DQ kmol I h 42000

reactor diameter D m 5

catalyst density pp kg/m3

1980

porosity of bed E

-

0.3

effect. factor in rl _Tli

-

0.5763

effect. factor in r2 ₁₁₂

-

0.9090

gas heat capacity

₄

_.""!!

₀

J/kg/K 5965.6

heat cap change of rl ~Cp.l J I mol I K -35.7

heat cap change of rz L14.2 J I mol / K 9.34

standard heat of rl L1rH I 298 J I mol -52814

standard heat ofr2 L1rHl98 J I mol -41270

The amount of gas in a quench is deterrnined by the desired decrease in temperature. Since the amounts of gas flowing through the reactor are greater in each of the following stages, the volume of quench gas must increase proportionally. For each stage the ratio of effluent versus quench equals roughly 3.75. For the total system, the ratio of reactor feed and quench is roughly unity.

(30)

•

]he Fixed Bed Reactor 22

Table 2 Overall results of the fIxed bed system simulation

parameter stage 1 stage 2 stage 3 stage 4

L; (m) 3.1 2.9 3.5 4.0

Ug (m / s) 0.5 0.6 0.7 0.8

Tin (K) 480.0 479.8 479.4 479.3

Tout (K) 524.3 524.5 525.0 524.1

Pin (atm) 50 49.8 49.6 49.4

out, lnbI (kmo I / h) 40040 47840 55520 72030 <l>out, co (kmol / h) 20450 24420 28340 32310

<l>out,MeOH (lanol / h) 108.0 213.0 326.1 438.4

f,J%) 0.48 0.40 0.36 0.31 rt(O) (kmol / k~ / s) 2.6 E-07 2.5 E-07 2.5 E-07 2.5 E-07 r2(0) (lanol / kg.,.. / s) -5.2 E-07 6.8 E-07 5.6 E-07 5.5 E-07

Considering that two reactions occur simultaneously, it is not possible to define the conversion unambiguously as was done for the bubble column slurry reactor, using CO as the key

component. In the kinetics describing the fIxed bed catalyst, CO partakes in both reactions. The conversion is therefore defined as follows

(19)

From the results presented here and in Appendix VII, the most obvious conc1usions that can be drawn are:

- the conversion decreases from 0.48 to 0.31 % from the fust to the fourth stage - the height ofthe stages increases from 3.1 to 4.0 m from the fust to the fourth stage - the pressure decreases from 50 to 49.4 atm from the fust to the fourth stage

The defInition of conversion in eq. 14 leads to low values for the methanol conversion, since we are using CO-rich gas. However, the rather low values for

ç

presented in table 2 are not only the consequence of a definition. It follows c1early from the reaction rate profIles that the side reaction occurring has a certain importance, which is not surprising considering the composition of the feed gas.

Moreover, the decreasing conversion can be explained through the nature ofthe reactions involved. The kinetic equations described in paragraph m.2 show a strong dependency on temperature and partial pressure of the products. The temperature range remains more or less equal while proceeding through the beds, but larger amounts of product are in the system and exert their influence.

The simulation Öztürk presents shows a higher methanol production [5], but here the

differenees are understandabie. Öztürk uses kinetics supplied by Rosenman [10], which gave totally unrealistic results in a simulation for the system presented here. The rate equations that Skrzypek [3] proposes come from an industrial catalyst, and show realistic reaction rate profIles. Therefore, these equations have been used here. The assumption of methanol production fiom CO2 rather than fiom CO was necessarily made in using these equations.

When running a plant on CO-rich gas, the methanol conversion is undestandably low, while the shift reaction is not affected. Moreover, Öztürk allows for a temperature rise of over 180 Kin one stage, using a superfIcial gas velocity of 0.1 m / s [5]. This is rather unrealistic in

(31)

•

However, some interesting observations are made in the same artic1e [5]. Diffusional limitations inside catalyst pellets are shown to have a considerabie influence on syngas conversion. The effectiveness factors used to describe this, may vary drarnatically with the axial reator coordinate. The superficial gas velocity is also thought to have a strong influence.

The height of each stage is deterroined by the temperature rise. The stages become

progressively higher, except for stage 1, which is significantly greater than would be expected. The principal difference between stage 1 and the other stages is the fact that at the inlet no products are present. This has a profound influence on the reaction rates, as can be seen in Appendix VIT. The rate of reaction 1 remains very low for the fITst meter of catalyst bed. Reaction 2 proceeds with a fairly high rate, but in the negative direction: that is to say, the reverse water-gas shift reaction occurs. This is an endothermic reaction which initially takes up heat, and thus allows the bed to become 3.1 m high before the temperature limit is exceeded. In the following stages the rate of reaction 2 shows a different behaviour. Although a negative value is not observed, rz goes through a minimw:n, which lies around 1 m from the beginning of the catalyst bed. The explanation for this lies in the changing concentrations of water and methanol due to reaction 1. Both rl and rz are seen to flatten out after some 4 meters of bed. This is also the effect of the increasing temperature on the exothermic reaetions.

The total conversion based on equation 19 for the four beds equals 1.0 %.

A flowsheet ofthe (minimum) ancillary equipment and the strean calculations as perfonned by ChemCad can be found in Appendix VITl.