Fischer-Tropsch Synthesis Revisited;
Efficiency and Selectivity Benefits from Imposing Temporal and/or
Spatial Structure in the Reactor
Fischer-Tropsch Synthesis Revisited;
Efficiency and Selectivity Benefits from Imposing Temporal and/or
Spatial Structure in the Reactor
Proefschrift
ter verkrijging van de graad van doctor aan de Technische Universiteit Delft,
op gezag van de Rector Magnificus prof. dr. ir. J.T. Fokkema, voorzitter van het College voor Promoties,
in het openbaar te verdedigen op dinsdag 27 januari 2004 om 15.30 uur
door
Ronald Martijn DE DEUGD
scheikundig ingenieur geboren te Dordrecht
Prof. Dr. F. Kapteijn
Samenstelling promotiecommissie:
Rector Magnificus, Technische Universiteit Delft, voorzitter Prof. Dr. J.A. Moulijn, Technische Universiteit Delft, promotor Prof. Dr. F. Kapteijn, Technische Universiteit Delft, promotor Prof. Ir. J. Grievink, Technische Universiteit Delft
Prof. Dr. Ir. H.E.A. van den Akker, Technische Universiteit Delft Prof. Dr. Ir. S.T. Sie, Technische Universiteit Delft
Prof. Dr. A. Holmen, Norwegian University of Science and Technology, Trondheim, Norway Dr. Ir. M.M.G. Senden, Shell Global Solutions International, Amsterdam
Reservelid:
Prof. Dr. R.F. Mudde, Technische Universiteit Delft
This research has been done as a part of Delft Interfacultary Research Center “Mastering the Molecules in Manufacturing”, abbreviated M3.
ISBN: 90-9017647-0
Copyright 2004 by R.M. de Deugd
All rights reserved. No part of the material protected by this copyright notice may be reproduced or utilized in any form or any means, electronic or mechanical, including photocopying, recording or by any information storage and retrieval system, without written permission of the author.
Contents
Chapter 1 Introduction 1
Chapter 2 FTS Experimental Set-Up 25
Chapter 3 Trends in Fischer-Tropsch Reactor Technology - Opportunities for Structured Reactors
37
Chapter 4 Activity and Selectivity of Monolithic Catalysts in Fischer-Tropsch Synthesis - Mass Transfer Limitations and Secondary Reactions
55
Chapter 5 Design of a Monolithic Loop Reactor for Fischer-Tropsch Synthesis 75
Chapter 6 The Potential of a New Generation of High-Activity Co/Al2O3 Catalysts for Fischer-Tropsch Synthesis
99
Chapter 7 Model-Based Optimization of the Periodic Operation of the Fischer-Tropsch Synthesis
113
Chapter 8 Evaluation of the Technical and Economic Feasibility of Fischer-Tropsch Synthesis for Large-Scale Production of Liquid Fuels
133
Summary 145
Samenvatting 149
Dankwoord 153
Publications and presentations 155
1
1.1 Fischer-Tropsch Synthesis (FTS) process description
The Fischer-Tropsch Synthesis (FTS) is a process converting synthesis gas into liquid fuels (Figure 1.1). Synthesis gas, a mixture of carbon monoxide and hydrogen, can be obtained from several sources like coal, natural gas or even biomass. After separation and upgrading FTS derived liquid can be used as transportation fuels. Non-fuel range products can be hydrocracked to the desired transportation fuels or be recycled to the synthesis gas generation process step.
The Fischer-Tropsch reaction forms hydrocarbons and water over a heterogeneous catalyst (Figure 1.2). The main hydrocarbon products are n-paraffins and α-olefins varying between methane and heavy waxes, but also other molecules such as branched paraffins, alcohols and aldehydes are found in the product mixture. Several metals such as ruthenium, cobalt and iron can catalyze the Fischer-Tropsch Synthesis. Cobalt and iron are commercially used, although ruthenium is a more active catalyst, but the price of ruthenium is too high to be compensated for by the higher activity. Hydrogen and carbon monoxide are consumed in a molar ratio of about 2:1. Over iron catalysts the overall consumption of hydrogen and carbon monoxide shifts to 1:1 as carbon monoxide is also consumed in the water gas shift reaction:
2 2 2
H O CO+ ←→H +CO (1.1)
The window of process conditions in which the Fischer-Tropsch Synthesis is performed is wide, especially in research studies. Typical conditions are pressures between 10-60 bar and temperatures in the range of 200-350 oC [1-3]. An important aspect of the Fisher-Tropsch Synthesis is the large heat of reaction, ∆Hr = -167 kJ/mol, which imposes strong requirements on the process design of
large-scale units.
Natural Gas
Coal
Fischer-Tropsch
Synthesis Product Cracking
Liquid Fuels Synthesis Gas
Biomass
Figure 1.1: Overall process scheme Fischer-Tropsch Synthesis
Figure 1.2: Overall reaction scheme Fischer-Tropsch Synthesis
n CO + 2n H
2Catalyst
(-CH
2-)
n+ n H
2O
(-CH
2-)
n+ n CO
22n CO + n H
2 Paraffins, Olefins Synthesis GasIntroduction
1.1.1 History and prospect
After its discovery by Franz Fischer and Hans Tropsch in the 1920s [4,5] FTS gained rapidly industrial importance and was a major source of transportation fuels in Germany during World War II. After World War II interest decayed, as a crude oil based fuel supply system was economically more attractive than the coal based Fischer-Tropsch Synthesis. Only South Africa continued using FTS based processes for fuel production as the boycotts against the Apartheid policies limited access to the world’s crude oil reserves. Since the 1970s and 1980s the interest in Fischer-Tropsch Synthesis is growing again. At the moment, most major oil companies are considering, developing or even operating Fischer-Tropsch Synthesis based processes to convert natural gas into liquid fuels. Shell [1] and South Africa’s SASOL [3,6] already operate one or more commercial gas-to-liquid (GTL) plants. Other oil companies like EXXON Mobil [7], BP and Conoco Phillips [8], but also small and independent companies such Rentech [9] and Syntroleum [10] have large pilot plants with capacities up to several hundreds of barrels per day. The plans for the near future of these companies include large-scale FTS plants up to 70000 barrels per day at several locations around the world. Most plans focus on locations near remote gas fields, as the value of stranded natural gas is lower than market prices in densely populated areas to compensate for the high transportation cost. Also locations where natural gas is associated with crude oil production are considered attractive for FTS as an alternative to avoid environmental damaging and therefore government regulated flaring or expensive reinjection.
1.1.2 Commercial FTS processes
At the moment, three different types of FTS processes are operated commercially. Shell and SASOL have a process focused on the production of middle distillates, typically C10-C20 products.
SASOL uses for this process a slurry bubble column reactor [6], while Shell has a multi tubular fixed bed reactor for their process [1]. The process operates at a relatively mild temperatures below 250 oC. The products of these processes are primarily unbranched n-paraffins, with excellent product quality. Especially the cetane number of the gasoil (>70) and the smoke point of the kerosene (> 100 mm) are much higher than the currently required specifications. The cold flow properties are worse as a result of the high concentration of unbranched molecules. As the FTS route ensures zero sulphur and aromatics levels in the products, Fischer-Tropsch fuels are very attractive from an environmental point of view.
Besides the low temperature slurry bubble column process, SASOL has a high temperature fluidized bed reactor typically operating at 340 oC producing lighter products, primarily in the gasoline range. The products contain large fractions of olefins and oxygenates, and even some aromatics. Therefore, the high temperature FTS process is not only used for transportation fuels production, but also as a source of base chemicals.
1.2 Reaction mechanism
The Fischer-Tropsch Synthesis mechanism has been subject of fierce debate for many decades. The complex nature of FTS makes it difficult to determine the exact chemistry of the route from carbon
monoxide and hydrogen to hydrocarbons. First of all, various catalytic systems are used for FTS such as iron, cobalt and ruthenium catalysts. Moreover, the complicated product mixture contains compounds ranging from methane to heavy waxes. Besides the main products n-paraffins and α-olefins, many other types of molecules like branched paraffins, alcohols, aldehydes and carbon acids are present, but normally in much smaller quantities.
Over the years several mechanisms have been proposed to describe the Fischer-Tropsch Synthesis. In the following paragraphs a short description is given for the major mechanisms found in literature.
1.2.1 Carbide mechanism
The most commonly used mechanism is the carbide mechanism (Figure 1.3). This mechanism assumes dissociative adsorption of both carbon monoxide and hydrogen. The surface carbide is hydrogenated to form a *CH2 building block. Propagation to longer hydrocarbon chains attached to
the catalyst surface takes place via coupling of these building blocks. Products are formed by termination reactions like hydrogenation, forming paraffins, or hydrogen abstraction, forming α-olefins. Other products like alcohols and aldehydes can also be formed via termination reactions with oxygen containing surface species [11].
1.2.2 Enolic mechanism
Many of the mechanisms proposed over the years assume that the FTS reaction proceeds via an oxygen containing intermediate building block rather than the carbidic monomer of the carbidic mechanism. The enolic mechanism and the CO insertion mechanism are examples.
CO (g) C* CH* CH2* CH3* CH4 (g) O* OH* H2O* H2O (g) H* H* H* H* H* H* CH2* C2H5* H* C2H4 (g) C2H6 (g) CH2* CH2* H* C3H7* C3H6 (g) etc. C3H8 (g) H* H* H2 (g) 2 H*
Introduction
In the enolic mechanism (Figure 1.4), the adsorption of carbon monoxide is non-dissociative. After reaction of adsorbed carbon monoxide with hydrogen a HC*OH intermediate is formed that acts as the building block. Coupling of two building blocks in combination with hydrogen to remove oxygen in the form of water leads to chain propagation. The growing chains can terminate to form the products via the same kind of reactions as in the carbidic mechanism [11].
1.2.3 CO insertion mechanism
Another mechanism frequently used in Fischer-Tropsch Synthesis research is the CO insertion mechanism. Like the enolic mechanism, the CO insertion mechanism (Figure 1.5) is based on oxygen containing intermediates. The basic intermediate is formed after reaction of carbon monoxide with a surface hydroxyl group and subsequent hydrogenation to an O*CH3 attached to
the catalyst surface. Subsequent reactions with carbon monoxide and two hydrogen molecules result in the growth of a hydrocarbon chain. Again termination to products can take place via the various reactions as also discussed above in the previous mechanisms [11].
1.2.4 Comparison of mechanisms
It is hard to proof which of the discussed mechanisms is the true mechanism of the Fischer-Tropsch Synthesis. The wide variety of catalysts and conditions (feed composition, pressure and temperature) under which FTS has been investigated in combination with the broad spectrum of products (n-paraffins, α-olefins, alcohols, aldehydes, carbonic acids, branched paraffins, internal olefins, etc.) make it probably impossible to speak about one mechanism as the true and unique mechanism of FTS. Dry [12] even suggests that a complex of parallel product-forming reactions, a combination of reactions from several mechanisms, is a more likely description. Furthermore, it should be realized that “disturbances” in the reaction mechanism may lead to different products like branched molecules and internal olefins besides the main products, n-paraffins and α-olefins. However, a high level of agreement is reached on the carbidic mechanism as the dominant one [13].
CO (g) CO H2 H2 H2O
+
etc. H3C C OH * H C OH * * * H C OH+
H C OH * H2 H2OFigure 1.4: Schematic representation enolic mechanism
O OH CO CH * * O O H2 C * HO H2 O CH3 * H2 H2O O CO CCH3 * O O H2 CCH3 * HO H O CH2CH3 * H2 H2O etc.
1.3 Reaction kinetics
Because of the complex chemistry in the mechanism of Fischer-Tropsch Synthesis, it is difficult to derive a relation describing the kinetics. Over the years, the FTS kinetics has been subject of many research studies resulting in a large number of relations describing the rate of consumption of either one of the two or both reactants, CO and H2. Van der Laan and Beenackers [13] published an
extensive review on the kinetics of the Fischer-Tropsch Synthesis over cobalt and iron catalysts. Due to the complexity of FTS almost all derived expressions are empirical power law equations or simple Langmuir-Hinshelwood-Hougen-Watson type relations. Table 1.1 provides a small overview of relations published in literature.
Intrinsic kinetic expression Catalyst Reference
2 2 2 H H CO CO a p R p + − = (1.2) Co/MgO/ThO2/kieselguhr Brötz [14]
(
2)
2 2 2 2 1 H CO H CO H CO a p p R bp p + − = +(1.3) Co/ThO2/kieselguhr Anderson [15]
2 2 0.5 H H CO CO a p R p + − = (1.4) Co/CuO/Al2O3 Yang et al [16] 2 2 0.55 0.33 H H CO CO a p R p +
− = (1.5) Co/La2O3/Al2O3 Pannell et al [17]
(
2)
0.5 3 0.5 1 CO H CO CO a p p R bp − = +(1.6) Co/Al2O3 Rautavuoma and Van
der Baan [18]
(
2 2)
0.5 0.5 2 0.5 0.5 1 H CO CO CO H CO a p p R bp cp dp − = + + +(1.7) Co/kieselguhr Sarup and
Wojciechowski [19]
(
2 2)
0.5 2 0.5 1 H CO CO CO H a p p R bp cp − = + + (1.8) Co/kieselguhr Wojciechowski [20](
2)
2 1 H CO CO CO a p p R bp − = +(1.9) Co/MgO/SiO2 Yates and Satterfield
[21] 2 2 2 H CO H CO CO H O a p p R p bp + − = +
(1.10) Fused Fe/K Atwood [22]
2 2 2 2 2 H CO H CO CO H H O a p p R p p bp + − = + (1.11) Precipitated Fe Deckwer [23]
Table 1.1: Overview of kinetic relations for CO or synthesis gas consumption rate. Constants a, b, c and d are temperature dependent constants.
Introduction
It should be realized that the conditions under which these relations were derived are quite different. Different reactor types have been used: fixed bed reactors, slurry reactors and Berty reactors. Also temperature, pressures, feed compositions and conversion level varied between the various research groups. Despite these experimental differences the resulting expressions show many similarities. All relations have a positive order in hydrogen, while carbon monoxide is reported both with a negative and a positive order. Carbon monoxide is present as an inhibitor in the denominator of almost all kinetic expressions.
In Chapter 5 of this thesis the kinetic expression as derived by Yates and Satterfield [21] is used for designing a Monolithic Loop Reactor for Fischer-Tropsch Synthesis. The temperature dependence of reaction constants a and b was derived by Maretto and Krishna [24] from the kinetic data by Yates and Satterfield
3 1 1 8.8533 10 exp 4494.41 493.15 a T − = × − (1.12) 1 1 2.226 exp 8236 493.15 b T = − − (1.13) 1.4 Product selectivity 1.4.1 ASF distribution
Although the reactor effluent contains a wide variety of components, the product distribution shows a strong regularity. As discussed above, the mechanism of the Fischer-Tropsch synthesis resembles a polymerization reaction as monomeric elements, -CH2-, are added to a growing hydrocarbon
chain (Figure 1.3). The chances for either chain propagation or termination are basically constant regardless the length of the hydrocarbon chain, but can be influenced by catalyst properties and reaction conditions.
Herington [25] defined a relation to characterize FTS product distributions in which βn is more or less constant for all carbon numbers.
1 n term n prop i n m r r m β ∞ + = =
∑
(1.14)Friedel and Anderson [26] introduced a slightly different relation to describe the molar product distribution of the Fischer-Tropsch Synthesis. A similar equation was earlier developed by Flory [27] to describe polymer kinetics. In Fischer-Tropsch Synthesis literature this product distribution is known as Anderson-Schulz-Flory (ASF) product distribution:
(
)
11 n
n
m = −α α − (1.15)
The ASF distribution can also be expressed in terms of a weight distribution (Equation 1.16):
(
)
2 11 n
n
w =n −α α − (1.16)
The product distribution is characterized by the chain growth probability αn, which physically represents the ratio between the rate of propagation and the sum of the rates of propagation and termination:
1 i prop i n n prop term i i n m r r r m α ∞ = + ∞ = = = +
∑
∑
(1.17)In FTS product distributions the chain growth probability α is again assumed to be more or less a constant for all carbon numbers.
The characteristic parameters αn and βn from the ASF and the Herington product distribution models are related to each other (Equation 1.18):
n n 1 1+ α β = (1.18)
Plotting the logarithm of the mol fraction mn against the carbon number n yields straight lines with
the logarithm of the chain growth probability α as the slope (Figure 1.6).
Due to the polymerization-like growth mechanism, the selectivity to a certain product or product range will always be limited. Figure 1.7 shows that the maximum primary selectivity towards products in the range of gasoline or diesel fuels is about 40%. For industrial application of the Fischer-Tropsch Synthesis, these selectivities are much too low. A solution to this problem is operating FTS with a high chain growth probability (α > 0.9) yielding a reasonable primary selectivity to fuel range products and a large amount of waxes which are subsequently cracked to fuel range products using a hydrocracker unit. Shell uses this route in its Shell Middle Distillate Synthesis (SMDS) process [28]
It is important to realize that the chain growth probability α and other selectivity aspects such as the olefin to paraffin ratio and the methane selectivity can be influenced both by the catalyst and the process conditions. Also carbon deposition on the catalyst is dependent on the process conditions applied. Schulz [29] provides a qualitative overview of the influence of several process conditions on the selectivity of the Fischer-Tropsch Synthesis (Table 1.2).
Figure 1.6: Schematic view of the ASF distribution and observed deviations
0 5 10 15 20 25 30 Carbon number Mol fraction (-) slope = ln α ASF non-ASF
Introduction Chain growth probability Olefin/paraffin ratio Carbon deposition Methane selectivity Temperature ↑ ↓ ↓ ↑ ↑ Pressure ↑ ↑ * * ↓ H2/CO ratio ↑ ↓ ↓ ↓ ↑ Conversion ↑ * ↓ ↑ ↑ Space velocity ↑ * ↑ * ↓
Table 1.2: Influence of certain process conditions on product selectivity characteristics; ↑: increase; ↓ decrease; * complex relation (adapted from Schulz [29])
1.4.2 Non-ASF behavior
Already early in the history of the Fischer-Tropsch Synthesis deviations from the Anderson-Schulz-Flory distribution are reported. Widely known are the higher than expected methane selectivities and low C2 yields [13]. Besides these deviations in the C1-C2 yields, also products with higher
carbon numbers do not always precisely follow the ASF distribution (Figure 1.6). Already in the 1940’s, higher than expected yields of high molecular products were reported [25]. Many other studies also confirm these kinds of observations, e.g. [30-34]. The explanations for these phenomena are very diverse. Satterfield et al. [31] suggest experimental artifacts causing the
deviations. Another explanation was found in two different types of sites with a different chain growth probability [35] or different termination mechanisms [20]. The bended overall product distribution can be built from the superposition of the two product distributions resulting in a bend in the product distribution around C10.
Figure 1.7: Product selectivity as a function of ASF chain growth probability α
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1 0 0.2 0.4 0.6 0.8 1
Chain growth probability α(-)
Mass fraction (-) Methane Ethane Propane Gasoline Diesel Wax Butane
Others [33,36] doubt this conclusion and suggest that secondary reactions cause the deviations for products with higher carbon numbers. This explanation is now widely regarded as the most plausible one [13]. Olefinic products can readsorb on the catalyst and undergo secondary reactions such as hydrogenation, reinsertion and isomerization (Figure 1.8). These reactions can either take place on the FTS growth sites or on different sites only catalyzing secondary hydrogenation and/or isomerisation [37]. The rate of readsorption of the α-olefins is the crucial process in explaining the non-ASF behavior. Increased readsorption reduces the net rate of termination (Equation 1.19), the sum of termination to both paraffins, Rt,n, and α-olefins, Ro,n, less the readsorption of α-olefins, Rr,n,
and causes therefore an increase in chain growth probability. The probability for a produced α -olefin to readsorb on the catalyst increases with carbon number for different possible reasons, as will be discussed below.
, , , ,
net term n t n o n r n
R =R +R −R (1.19)
Increased secondary hydrogenation and reinsertion of α-olefins into the growth process therefore leads to deviations from the ordinary ASF product distribution. The olefin to paraffin ratio decreases with increasing carbon number as a larger fraction of the olefins are hydrogenated to alkanes and the apparent chain growth probability increases as result of reinsertion in the growth process of the readsorbed olefins, shifting the product distribution towards product with longer chain lengths.
Although opinions diverge about the causes of the chain length dependent residence times and associated α-olefin readsorption, the common theme in all explanations is an increased and chain length dependent local concentration of α-olefins resulting in a larger driving force for readsorption. In literature, three different chain length dependent processes are posed:
• Transport limitations [32,33] • Physisorption [34,36]
• Product solubility in FTS wax [34,37-40]
1.4.2.1 Transport limitations
Diffusion is an important phenomenon in Fischer-Tropsch Synthesis. Both reactants and products need to be transported to and from the catalytic sites. The stoichiometry of the reactants is very
*C
n-1*C
n*C
n+1C
n paraffins rp rpC
n olefins paraffins rhs rr ro#C
nC
n ros rrs*C
1CO
H
2 rt catalyst catalystIntroduction
important for the selectivity of the Fischer-Tropsch Synthesis. More than stoichiometric amounts of hydrogen lead to a lower chain growth probability and therefore shift the product distribution to shorter products. Diffusivity plays an important role, as both hydrogen and carbon monoxide should be delivered to the catalytic sites in stoichiometric amounts. However, hydrogen diffusivity is faster than carbon monoxide diffusivity leading to excess amounts of hydrogen towards the inside of the catalyst pellet [41]. Larger particles therefore not only suffer from reduced effectiveness, but also from reduced selectivity to heavy products.
Diffusive removal of the reactive products also influences the selectivity of the process. The residence time inside the catalyst pore affects the chances of secondary reactions such as hydrogenation and reinsertion.
Little experimental data are available on the diffusivity of Fischer-Tropsch products in FTS wax. Erkey et al. [42] report experimental diffusion coefficients for octane, dodecane and n-hexadecane in Fischer-Tropsch wax. They roughly find a linear dependency between diffusion coefficient and carbon number. Wilke and Chang [43] developed a correlation to calculate diffusion coefficients which predicts that diffusion coefficients scale with n-0.6.
Iglesia and co-workers [32,33,41] formulated a model for the Fischer-Tropsch process based on reaction and diffusion in the catalyst pores in order to explain deviation from the ASF distribution. They derived two models for the reactants and the products. In the olefin readsorption model the behavior of the products, paraffins and olefins, are considered only, while the CO hydrogenation model deals with the diffusion and consumption of carbon monoxide and hydrogen. Iglesia et al.
justify solving the two models separately by assuming that the reactant diffusivity is much faster than the product diffusivity. The interesting phenomena described in the models with respect to the products occur therefore at much shorter diffusion lengths than reactant depletion. The result of this assumption is that it is allowed to use constant reactant concentrations in the olefin readsorption model.
In the olefin readsorption model differential equations are derived for the paraffins and olefins. Moreover, steady state balances for the growing chains on the catalyst surface are given. The reaction mechanism includes chain propagation, termination to paraffins and olefins by hydrogen addition and elimination, respectively, and olefin readsorption and reinsertion. Moreover, secondary hydrogenation on other than chain growth sites is considered. Iglesia et al. assume all reaction rates
to be independent of chain length.
Steady-state balance for surface chains with carbon number n at radial position r:
(
)
1 1 ,
0=kpθ θn− −θn N saT v−koθnN aT v+krθHColefin nN aT v−ktθ θH nN saT v (1.20) Balance for α-olefin with chain length n:
2 ,
, 2 , ,
0 De n Colefin n ko nN aT v kr HColefin nN aT v ks HsColefin nN aTs v
r θ θ θ
∂
= + − −
∂ (1.21)
Balance for n-alkane with chain length n:
2 , , 2 , 0 De n Cparaffin n kt H nN saT v ks HsColefin nN aTs v r θ θ θ ∂ = + + ∂ (1.22)
Boundary conditions for both α-olefins and n-alkanes: , 0 i i bulk C =C at r=r (1.23) 0 0 i C at r r ∂ = = ∂ (1.24)
The original model as derived by Iglesia and co-workers differs from the above equations in several ways. Iglesia et al. assumed the surface concentrations of the monomer θ1 and hydrogen θH to be constant, but this approximation is only valid for fixed pressure, temperature and composition (H2/CO ratio). Therefore, in this study this simplification is not made as carbon monoxide and
hydrogen concentrations are changing significantly over both the diffusion length of the catalyst pellet or layer and over the length of the reactor.
Furthermore, Iglesia and co-workers give concentrations in the pores as liquid concentrations and they do not define surface concentrations as fractions of the total number of active sites NT, but as
concentrations. The axial direction of the Iglesia model has been left out to simplify the model.
For small pellets, Iglesia et al. assume the H2/CO ratio to be constant, but for larger particles the
intraparticle gradients of the reactants need to be included. The CO hydrogenation model describes the diffusivity and consumption of synthesis gas inside the catalyst pellet.
The rate of synthesis gas consumption can typically be described by a relation as shown in Equation 1.25:
(
1 2)
CO a b H syngas syngas c CO CO C C R k K C = + (1.25)Hydrogen and carbon monoxide consumption follow similar relations differing from each other by stoichiometric constants. 2 2 CO H 1 1 ν ν syngas CO H R = R = R (1.26)
Similar to the olefin readsorption model molar balances can be derived for carbon monoxide and hydrogen (Equations 1.27 and 1.28).
Molar balance carbon monoxide :
2 , 2 0 CO ν e CO v CO syngas C D a R r ∂ = − ∂ (1.27)
Molar balance hydrogen
2 2 2 2 , 2 0 De H CH avνH Rsyngas r ∂ = − ∂ (1.28)
Boundary conditions for both carbon monoxide and hydrogen
, 0 i i bulk C =C at r=r (1.29) 0 0 i C at r r ∂ = = ∂ (1.30)
Introduction
Rearranging and substitution of variables (Equations 1.31-1.39) into mass balances 1.20-1.22, 1.27 and 1.28 results in dimensionless expressions for the reaction and diffusion in the catalyst layer or pellet (Equations 1.40-1.44). Also the synthesis gas consumption rate has been rewritten in dimensionless form to obtain analogy between the mass balances for reactants and products.
, * , 0 olefin n olefin n C C C = (1.31) * , , 0 paraffin n paraffin n C C C = (1.32) * 0 CO CO C C C = (1.33) 2 2 * 0 H H C C C = (1.34) 0 r Y r = (1.35) κ t t p k k = (1.36) κ o o p k k s = (1.37) κ r 0 r p k C k s = (1.38) 0 κ s Ts s p T k C N k sN = (1.39)
Table 1.3: Dimensionless groups
Dimensionless balance for surface chains:
(
)
*1 1 ,
0=θ θn− −θn −κtθ θn H −κoθn+κrθHColefin n (1.40)
Dimensionless balance for α-olefins:
(
)
2 *
, 2 * *
, ,
2
0 olefin n n κo n κr H olefin n κs Hs olefin n C C C Y θ θ θ ∂ = + Φ − − ∂ (1.41)
Dimensionless balance for n-alkanes:
(
)
2 * , 2 * , 2 0 Cparaffin n n n κt n H κs HsColefin n Y θ θ θ ∂ = + Φ + ∂ (1.42)Dimensionless balance for carbon monoxide:
(
2)
2 * 2 * * * 2 0 CO , CO syngas CO H C R C C Y ∂ = − Φ ∂ (1.43)Dimensionless balance for hydrogen
2 2 2 2 * 2 * * * 2 0 H H syngas( CO, H ) C R C C Y ∂ = − Φ ∂ (1.44)
The dimensionless group Φi2, the square of the Thiele modulus, which is found in the mass balances for both the reactants and the products plays a central role in the model of Iglesia and co-workers. It represents a ratio between the rates of reaction and diffusion for each of the components present in the system (Equation 1.45). Except for the rate constant ki, which is defined differently for reactants
and products (Equations 1.46 and 1.47, respectively), the expression is identical for reactants and products.
2 2 0 , 0 i T v i e i r k N a D C Φ = (1.45) i 0 ν a b i syngas k = k C + (1.46) i p k =k s (1.47)
The several parameters in the dimensionless group Φi2 can be rearranged to separate the physical properties of reactants and products and process conditions from the catalyst properties. The final form of Φi2 in the studies by Iglesia et al. is expressed in Equation 1.45. Φi2 is split in two other groups ψi and χ. ψi contains the physical properties and the effects of the process conditions, while
χ represents the catalyst properties.
Iglesia and co-workers relate the specific area to porosity and pore radius by assuming cylindrical pores (Equation 1.48). 2 v p a r ε = (1.48)
As Iglesia et al. used a different definition for surface concentrations Avogadro’s constant is introduced. θM expresses the loading of the catalyst.
2 2 0 0 2 r θ = i M i i i A p k D N C r ε ψ χ Φ = ∗ ∗ (1.49)
Iglesia and co-workers simply assume diffusion to be independent of catalyst properties. However, this is not realistic as diffusion inside catalyst pores is seriously restricted [44]. Restricted diffusivity inside catalyst pores is reduced by a factor 2 to 100 compared to free diffusivity depending on porosity and tortuosity. Tortuosity is in practice a function of porosity and can be approximated by 1/ε [45] (Equation 1.50): 2 , τ e i i i D =ε D =ε D (1.50)
Substituting Equation 1.50 into the relation for Φi2 moves porosity from the numerator to the denominator of χ (Equation 1.51). 2 2 0 0 2 R = i T i i i p k N D C r ψ χ ε Φ = ∗ ∗ (1.51)
A striking difference between model originally derived by Iglesia and co-workers and the model presented here is the position of porosity in catalyst properties group χ. Iglesia et al. arrive at an
expression in which porosity is in the numerator, while in this thesis an inverse relation is derived. This latter dependency is intuitively easier to understand as χ expresses the accessibility of the catalyst like the well-known Thiele modulus.
Iglesia and co-workers used their model to fit experimental data and found a good correlation between experimental and modeling data. Figure 1.9 shows the selectivity to C5+ products and
methane as a function of the catalyst parameter χ. The excellent fit of the model to the experimental data suggest that the catalyst parameter χ is an interesting design variable for catalyst design.
Introduction
However, besides the kinetic constants the diffusivity is used as a fitting parameter. Iglesia and co-workers use an exponential relation to describe diffusivity as a function of carbon number (Equation 1.52) 0.3 0 n i D =D e− (1.52)
Literature suggests, however, another relation between diffusivity and carbon number. Experimental data by Erkey et al. [42] show approximately a linear decrease of diffusivity with
carbon number, while the correlation by Wilke and Chang arrives at dependency of Dn∝n-0.6. The
relations by Erkey et al. and Wilke and Chang more or less coincide, while the fitted relation by Iglesia and co-workers assumes a much stronger decrease in diffusivity for longer hydrocarbon chains (Figure 1.10). The large difference between experimental diffusivity data and the fitted relation suggest that Iglesia and co-workers attribute their explanation for non-ASF product distributions, at least partially, to the wrong phenomenon.
Figure 1.9: Effect of catalyst parameter χ on the selectivity to methane (Panel A) and C5+ products (Panel B) according
to Iglesia et al. [32] (Conditions: Co catalysts on Al2O3, SiO2 and TiO2, T = 473 K, H2/CO feed ratio = 2.1, p = 20 bar,
50-60% conversion) 0 5 10 15 10 100 1000 10000 100000 Catalyst parameter χ (* 10-16 m-1) CH 4 s e le c tiv ity (-) 75 80 85 90 95 100 10 100 1000 10000 100000 Catalyst parameter χ (* 10-16 m-1) C5+ selectivity (-)
Moreover, the structure of the catalyst properties group χ suggest that the various parameters can be tuned independently to obtain the optimal catalyst, but some hidden dependencies are present between site density, porosity and pore radius. For instance, an increase of the pore radius leads to a decrease of the specific surface area given a fixed porosity. As site density, NT, is defined as moles
active sites per square meter, NT will also increase given the catalyst loading and active phase
dispersion.
1.4.2.2 Physisorption
A second phenomenon influencing the chain-length dependence of secondary reactions is physisorption. The physisorbed state is a transition state between the chemisorbed state and the gas phase which is governed by Van der Waals interaction and repulsion forces [13]. The physisorption energy is determined by the molecular structure of both the adsorbent and the absorbate. Van der Laan and Beenackers [13] provide an overview of literature data on the carbon number dependency and shows a linear relation between physisorption enthalpy and carbon number. Kuipers et al. [34] reach the same conclusion for Gibbs free energy for physisorption.
1.4.2.3 Solubility
Several research groups [34,37-40] attribute non-ASF deviations for products with higher carbon number to the increasing solubility of hydrocarbons with longer chains. Higher concentrations lead to increased contact time and chances of readsorption. This mechanism raises the chances of secondary reactions such as hydrogenation and reinsertion of the higher olefins leading to decreasing olefin to paraffin ratios with increasing chain length.
The influence of solubility is subject of fierce debate. Iglesia et al. [33] argue that in absence of transport limitations gas and liquid phase are in thermodynamic equilibrium in a reacting system. The presence of a liquid phase does therefore not affect residence time distribution or the kinetic driving force. Kuipers et al. [34] state that chemisorption is a non-activated process. The Figure 1.10: Comparison of diffusivity data and correlations; solid line: correlation according to Iglesia et al. [32],
dashed line: correlation according to Wilke and Chang [43]; ▲ experimental data Erkey et al. [42]
0.01 0.1 1 1 6 11 16 Carbon number Normalised diffusivity
Introduction
readsorption rate is controlled by the effective “collision rate” which is governed by the concentration in the physisorbed layer. This concentration is again determined by the concentration in the wax.
Kuipers et al. [34] calculated the solubility of FT products using Raoult’s law and the vapour pressures of the pure components from the Clausius-Clapeyron equation. They find an exponential behavior of the solubility as a function of the carbon number which can be fitted with e(0.55±0.1)n. Schulz and Claeys [37] present a model in which they explain non-ASF behavior solely on the chain length dependent solubility. Using Henry’s law instead of Raoult’s law results in about a similar dependency for the chain length dependency of solubility (e0.43nat 190 oC and e0.38n at 250
o
C) as obtained by Kuipers et al. [34].
1.4.2.4 Combinations
All three factors (transport limitations, solubility, physisorption) discussed above influence the contact time of FTS reaction products with the catalyst and therefore the extent of secondary reactions on the FTS product distribution. Simply assuming one phenomenon to be dominant would be an oversimplification. A combination including diffusivity, solubility and physisorption is needed for an accurate description.
Kuipers et al. [34] performed gas phase model experiments using Co-foils. They determined the influence of surface area, the presence of a wax coating, the H2/CO ratio, and flow velocity on
concentration and selectivity and evaluated their experiments using a model accounting for effects of diffusion, solubility and physisorption on the paraffin to olefin ratio. They choose for the ratio between paraffins and olefins and not for the deviations from ASF distribution as selectivity criterion as secondary hydrogenation was stronger than reinsertion under the reaction conditions applied in their experiments, at least up to C14. The role of secondary reactions is therefore not to be
deduced from deviation from the ASF distribution. As both secondary hydrogenation and reinsertion are governed by the same factors, they use the chain length dependence of the paraffin to olefin ratio to study the effects of secondary reactions.
The key in describing non ASF-behavior is the chemisorption of reactive olefins on the surface of the catalyst. The model distinguishes four states for FTS products: chemisorbed on the catalyst, physisorbed in an interface layer, dissolved in wax and in the vapor phase (Figure 1.11). Raoult’s law determines the solubility in the wax at x = d. Diffusivity defines the concentration gradient over the wax layer. The concentration in the interface is a result of competitive physisorption. All of these three processes are chain length dependent. The final equation describing the paraffin to olefin ratio is given in Equation 1.53:
1 ~ max φ n G phys gas R T n s n n n nvap P A d e O C V D ∆ ∝ + (1.53)
In this equation several chain length dependent variables are present. Solubility, Cnvaps , is
proportional to e(-0.55±0.1)n, while molar volume,
~ max
n
V , scales linearly with carbon number. Diffusivity, Dn, has a chain length dependency of n-0.6. The change of the Gibbs free energy of
physisorption is stated to be n∆G1,phys, in which ∆G1,phys is found to be (0.2 ± 0.1)*Rgas*T.
In absence of diffusion limitations the paraffin to olefin ratio is proportional to Equation 1.54, while diffusion limitations reduce the effect of solubility in favor of the diffusion term and Equation 1.55 applies. ( ) 1 0.75 0.1 ~ max 1 phys gas n G R T n n s n n nvap P e e O C V n ∆ ± ∝ ∝ (1.54) 1 (0.2 0.1) 0.6 / phys gas n G R T n n n n P e D e n O ∆ ± ∝ ∝ (1.55)
Kuipers et al. [34] performed experiments with bare Cobalt foils and Cobalt foils covered with a wax coating to test their model. The slope of the Pn/On ratio as a function of carbon number should
decrease strongly upon the introduction of diffusion limitations. This dependence is indeed found by Kuipers and co-workers.
It is important to realize that in the study by Kuipers et al. a non porous catalyst is used and all products leave the reactor in gas phase. In cases where the reactor effluent also consists of liquids the influence of solubility disappears. Effects of intraparticle diffusion may introduce the need to modify the presented model to account for those.
Schulz and Claeys [37] oppose the introduction of physisorption into the model to explain non-ASF-behavior. They argue that Kuipers et al. [34] found their physisorption term only on fitting of experimental data and that no literature evidence is supporting that view. However, literature does provide data on chain length physisorption of α-olefins [13]. Moreover, Kuipers et al. show
catal y st interfac e wa x gas phase solubility diffusion phy siso rption x d 0 -δ c
Introduction
experimental results for the olefin to paraffin ratio as a function of carbon number supporting their theory. Others [46] also report experimental results in agreement with the theory by Kuipers et al. (kr ∝ e(0.29±0.07)n). Though opposing the modeling approach followed by Kuipers et al., Schulz and
Claeys do not show experimental evidence supporting their views.
1.5 Research objectives
The main challenge in Fischer-Tropsch Synthesis is maximizing the selectivity of the process to desired products or product ranges. However, the polymerization-like nature of FTS poses strong limitations on efforts to, for instance, selective production of valuable and high quality middle distillate range products. The Anderson-Schulz-Flory distribution of the FTS products limits the maximum primary selectivity to middle distillates, C10-C20, to only 39%. Moreover, all kind of
process shortcomings such as undesired temperature and concentration gradients prevent operation at optimal conditions. For a feasible commercial process this selectivity is insufficient. At the moment, commercial processes aiming at middle distillate production operate with as high as possible chain growth probability leading to reasonable primary selectivity to fuel range products and a high selectivity to waxes that can subsequently be converted to transportation fuels in a hydrocracker unit. However, new methods that can raise FTS selectivity to desired product ranges, preferably beyond the ASF limitations, are therefore very interesting.
Moreover, the strong exothermal nature of the Fischer-Tropsch Synthesis (∆Hr = -167 kJ/mol)
poses important limitations on the reactor design and operation of the FTS process. The adiabatic temperature rise of FTS, 1600 K, makes heat removal a crucial issue, especially considering that, as already mentioned earlier in this paragraph, temperature is an important parameter in controlling process selectivity. FTS process design has always been a struggle to cope with this, resulting in difficult compromises with respect to reactor productivity.
In this thesis opportunities are explored to improve the Fischer-Tropsch Synthesis process with respect to selectivity and productivity. Proposed improvements are found in introducing spatial or temporal structure in the FTS reactor. The aim of introducing these structures is either to order the complexity of the Fischer-Tropsch Synthesis or to benefit from it by making use of it for increasing selectivity.
Chapter 2 discusses the experimental facilities built to experimentally investigate the Fischer-Tropsch Synthesis. The flexible reactor set-up consists of six fixed bed reactors and a monolith reactor of which six can be used parallel at the same time. The configuration of the experimental unit allows on-line analysis of the reactants and the most important products.
Chapter 3 provides a analysis of process requirement of FTS using the systematic hierarchal method developed by Sie and Krishna [47]. After analyzing the needs of the Fischer-Tropsch Synthesis, the present commercial reactor types are compared to the reactor needs from the systematic analysis. As none of the existing reactor types appears to be an ideal solution for Fischer-Tropsch Synthesis alternative reactors are introduced; a Monolithic Loop Reactor and a Gas Lift Reactor.
Especially the Monolithic Loop Reactor might be a promising alternative to the present reactor configurations.
Therefore, Chapter 4 discusses experimental efforts to use monolithic catalysts for Fischer-Tropsch Synthesis. Aim of the investigations is to give a proof of principle of the monolithic concept for FTS and to explore the effects of the catalyst layer thickness on the monolith on the selectivity and activity.
Chapter 5 continues exploring the monolithic concept for FTS in a Monolithic Loop Reactor design study for a 5000 ton middle distillate per day plant with monolithic catalysts. The objective of the study is to build a mathematical model to describe the flow of gas and liquid through the monolith channels, the mass transfer from the channel to the catalytic walls and the reaction and diffusive transport inside these catalytic walls. Moreover, a window of possible conditions is explored as well as a number of alternative configurations.
Chapter 6 deals with the key questions posed in the modeling study of Chapter 5. As the mass transfer from the monolith channel to the catalytic walls is much faster than the reaction rate of the catalyst used in the modeling study, more active catalysts are needed to fully utilize the productivity potential of monolithic catalysts for FTS. In this chapter, the activity of a new type of cobalt catalysts, HDC catalysts, with both a high dispersion and a high cobalt loading is reported.
Chapter 7 describes a modeling study to improve the selectivity of the Fischer-Tropsch Synthesis in a different way, namely by introducing a temporal rather than a spatial structure. The modeling study describes Fischer-Tropsch Synthesis operated with a periodically changing feed composition. The aim of the study is to investigate possible beneficial effects on the product selectivity by manipulating chain initiation, propagation and termination steps of the reaction mechanism.
Finally, Chapter 8 provides an evaluation of FTS process from both an economical and a technological point of view. By analyzing the general patterns of technology adoption, the shifting interest in Fischer-Tropsch Synthesis since its discovery is discussed. Derived from this discussion, technical needs for FTS are formulated and translated to the objectives and findings of thesis.
Nomenclature
a kinetic constant mol/(kgcat*bar2*s)
av specificinternal surface area m2/m3
A surface area catalyst m2
b kinetic constant 1/bar
C concentration mol/m3
s nvap
C saturated vapor phase concentration mol/m3
d diffusion length m
D diffusion coefficient m2/s
∆G1,phys Gibbs free energy change physisorption C1 J/mol
KCO adsorption constant m3/mol
k rate constant 1/s,m3/(mol*s), m3(a+b)/(mol(a+b)*s)
m mol fraction -
n carbon number -
Introduction
NT total concentration of sites mol/m2
R reaction rate mol/(s*m2),-
Rgas gas constant J/(mol*K)
r particle radius m
rp pore radius m
s number of surrounding sites -
T temperature K
~ max
n
V molar volume m3/mol
w weight fraction -
Y dimensionless length -
Greek symbols
α chain growth probability -
β Characteristic parameter -
θ surface fraction -
θM metal loading sites/m2
ϕ volumetric flow rate m3/s
Φ Thiele modulus -
ψ compound specific part Thiele modulus m/mol, m
ε porosity -
κ dimensionless rate constant -
ν stoichiometric coefficient -
τ tortuosity -
χ catalyst specific part Thiele modulus mol/m, 1/m Subscripts
0 reference 1 monomer
bulk bulk properties
CO carbon monoxide e effective i compound i H2 hydrogen H surface hydrogen n carbon number o termination to olefins p propagation prop propagation olefin olefin
paraffin paraffin
r readsorption of olefins
s secondary hydrogenation
syngas synthesis gas
t termination to paraffins
term termination
Superscripts
* dimensionless a power law constant b power law constant
References
1. S.T. Sie, Process Development and Scale-Up: IV. Case History of the Development of a Fischer-Tropsch Synthesis Process, Reviews in Chemical Engineering 14 (1998) 109.
2. B. Jager and R.L. Espinoza, Advances in Low Temperature Fischer-Tropsch Synthesis, Catal. Today 23 (1995) 17.
3. A.P. Steynberg, R.L. Espinoza, B. Jager, and A.C. Vosloo, High temperature Fischer-Tropsch Synthesis in Commercial Practice, Appl. Catal. A: General 186 (1999) 41.
4. F. Fischer and H. Tropsch, DE484337 (1925).
5. F. Fischer and H. Tropsch, Uber die Herstellung synthetischer ölgemische (Synthol) durch Aufbau aus Kohlenoxyd und Wasserstoff, Brennstoff-chemie 4 (1923) 276.
6. R.L. Espinoza, A.P. Steynberg, B. Jager, and A.C. Vosloo, Low Temperature Fischer-Tropsch
Synthesis from a Sasol Perspective, Appl. Catal. A: General 186 (1999) 13.
7. B. Eisenberg, R.A. Fiato, C.H. Mauldin, G.R. Say, and S. Soled, EXXON's Advanced Gas-to-Liquids Technology, Studies in Surface Science and Catalysis 119 (1998) 943.
8. T. Knott, Alchemy in Alaska, BP frontiers (2002) 14. 9. http://www.rentechinc.com
10. http://www.syntroleum.com
11. R.B. Anderson, The Fischer-Tropsch Synthesis, Academic Press, Orlando (1984).
12. M.E. Dry, In: J.R. Anderson and M. Boudart (eds) Catalysis, Science and Technology, Vol 1. Springer-Verslag, Berlin, (1981).
13. G.P. Van der Laan and A.A.C.M. Beenackers, Kinetics and Selectivity of the Fischer-Tropsch Synthesis: A Literature Review, Catal. Rev. -Sci. Eng. 41 (1999) 255.
14. W. Brötz, Zur Systematik der Fischer-Tropsch Katalyse, Zeitschrift für Elektrochemie 53 (1949) 301.
15. R.B. Anderson, In: P.H. Emmett (ed) Catalysis Vol IV: Hydrocarbon Synthesis, Hydrogenation and Cyclization, Reinhold Publishing Corporation, New York, (1956).
16. C. Yang, F.E. Massoth, and A.G. Oblad, In: E.L. Kugler and F.W. Steffgen (eds) Hydrocarbon Synthesis from Carbon Monoxide and Hydrogen. ACS, Washington, D.C., (1979).
Introduction
17. R.B. Pannell, C.L. Kibby, and T.P. Kobylinski, A Steady-State study of Fischer-Tropsch Product Distributions over Cobalt, Iron and Ruthenium, Studies in Surface Science and Catalysis 7A (1980) 447.
18. A.O.I. Rautavuoma and H.S. van der Baan, Kinetics and Mechanism of the Fischer-Tropsch Hydrocarbon Synthesis on a Cobalt on Alumina Catalyst, Appl. Catal. 1 (1981) 247.
19. B. Sarup and B.W. Wojciechowski, Studies of the Fischer-Tropsch Synthesis on a Cobalt Catalyst II. Kinetics of Carbon Monoxide Conversion to Methane and to Higher Hydrocarbons, Can. J. Chem. Eng. 67 (1989) 62.
20. B.W. Wojciechowski, The Kinetics of Fischer-Tropsch Synthesis, Catal. Rev. -Sci. Eng. 30 (1988) 629.
21. I.C. Yates and C.N. Satterfield, Intrinsic Kinetics of the Fischer-Tropsch Synthesis on a Cobalt Catalyst, Energy & Fuels 5 (1991) 168.
22. H.E. Atwood and C.O. Bennett, Kinetics of the Fischer-Tropsch Reaction over Iron, Ind. Eng. Chem. Prod. Res. Dev 18 (1979) 163.
23. W.D. Deckwer, R. Kokuun, E. Sanders, and S. Ledakowicz, Kinetic Studies of Fischer-Tropsch Synthesis on Suspended Fe/K Catalyst: Rate Inhibition by CO2 and H2O,
Ind.Eng.Chem.Proc.Des.Dev. 25 (1986) 643.
24. C. Maretto and R. Krishna, Modelling of a Bubble Column Slurry Reactor for Fischer-Tropsch Synthesis, Catal. Today 52 (1999) 279.
25. E.F.G. Herington, The Fischer-Tropsch Synthesis considered as a Polymerization Reaction, Chemistry and Industry (1946)
26. R.A. Friedel and R.B. Anderson, Composition of Synthetic Liquid Fuels. I. Product Distribution and Analysis of C5-C8 Paraffin Isomers from Cobalt Catalyst, J. Am. Chem. Soc.
72 (1950) 1212.
27. P.J. Flory, Molecular Size Distribution in Linear Condensation Polymers, J. Am. Chem. Soc. 58 (1936) 1877.
28. S.T. Sie, M.M.G. Senden, and H.M.H. van Wechem, Conversion of Natural Gas to Transportation Fuels via the Shell Middle Distillate Synthesis Process, Catal. Today 8 (1991) 371.
29. H. Schulz, Trends in Research and Development of Coal Conversion to Liquid Fuels and Basic Chemicals in Europe, Pure & Appl. Chem. 51 (1979) 2225.
30. C.N. Satterfield and G.A. Huff, Carbon Number Distribution of Fischer-Tropsch Products Formed on an Iron catalyst in a Slurry Reactor, J. Catal. 73 (1982) 187.
31. C.N. Satterfield, G.A. Huff, and J.P. Longwell, Product Distribution from Iron Catalysts in Fischer-Tropsch Slurry Reactors, Ind. Eng. Chem. Prod. Res. Dev 21 (1982) 465.
32. E. Iglesia, S.C. Reyes, and S. Soled, In: E.R. Becker and C.J. Pereira (eds) Computer-Aided Design of Catalysts. Marcel Dekker, Inc., New York, (1993).
33. E. Iglesia, S.C. Reyes, and R.J. Madon, Transport-Enhanced α-Olefins Readsorption Pathways in Ru-Catalyzed Hydrocarbon Synthesis, J. Catal. 129 (1991) 238.
34. E.W. Kuipers, I.H. Vinkenburg, and H. Oosterbeek, Chain Length Dependence of Alpha-Olefin Readsorption in Fischer-Tropsch Synthesis, J. Catal. 152 (1995) 137.
35. G.A. Huff and C.N. Satterfield, Evidence for Two Chain Growth Probabilities on Iron Catalysts in the Fischer-Tropsch Synthesis, J. Catal. 85 (1984) 370.
36. E.W. Kuipers, C. Scheper, J.H. Wilson, I.H. Vinkenburg, and H. Oosterbeek, Non-ASF Distributions Due to Secondary Reactions during Fischer-Tropsch Synthesis, J. Catal. 158 (1996) 288.
37. H. Schulz and M. Claeys, Kinetic Modelling of Fischer-Tropsch Product Distributions, Appl. Catal. A: General 186 (1999) 91.
38. H. Schulz, K. Beck, and E. Erich, Kinetics of Fischer-Tropsch Selectivity, Fuel Proc. Techn. 18 (1988) 293.
39. L.M. Tau, H.A. Dabbagh, and B.H. Davis, Fischer-Tropsch Synthesis: Carbon-14 Tracer Study of Alkene Incorporation, Energy & Fuels 4 (1990) 94.
40. W.H. Zimmerman, D.B. Bukur, and S. Ledakowicz, Kinetic Model of Fischer-Trospch Synthesis Selectivity in the Slurry Phase, Chem. Eng. Sci. 47 (1992) 2707.
41. E. Iglesia, S.C. Reyes, R.J. Madon, and S.L. Soled, Selectivity Control and Catalyst Design in the Fischer-Tropsch Synthesis: Sites, Pellets, and Reactors, Advances in Catalysis (1993) 221. 42. C. Erkey, J.B. Rodden, and A. Akgerman, Diffusivities of Fischer-Tropsch Gas and n-Alkanes
in Fischer-Tropsch Wax, Energy & Fuels 4 (1990) 275.
43. C.R. Wilke and P. Chang, Correlation of Diffusion Coefficients in Dilute Solutions, AIChE J. 1 (1955) 264.
44. O. Levenspiel, Chemical Reaction Engineering, 3rd edn., Wiley, New York (1999).
45. T. Elias-Kohav, M. Sheintuch, and D. Avnir, Steady-State Diffusion and Reactions in Catalytic Fractal Porous Media, Chem. Eng. Sci. 46 (1991) 2787.
46. G.P. Van der Laan and A. Beenackers, alpha-Olefin Readsorption Product Distribution Model for the Gas-Solid Fischer-Tropsch Synthesis, Studies in Surface Science and Catalysis 119 (1998) 179.
47. R. Krishna and S.T. Sie, Strategies for Multiphase Reactor Selection, Chem. Eng. Sci. 49 (1994) 4029.
2
FTS Experimental Set-Up
Fischer-Tropsch Synthesis is a demanding process, as it requires dedicated equipment for gas-liquid separation, sampling and analysis. For this research project an experimental set-up has been developed dealing with these requirements in a flexible way.
The set-up is designed not only for gas phase feed operation, but also allows addition of liquids to work in either trickle-flow or supercritical mode. Fast and on-line GC equipment enables direct analysis of reactants and products from C5 to C20.
Moreover, to speed up the slow and laborious catalyst screening and testing the set-up has been built with six parallel fixed-bed reactors. One of these reactors can be replaced with a reactor suited for monolithic catalyst research. Altogether this makes the designed and built set-up a versatile tool for Fischer-Tropsch Synthesis research studies.
2.1 Introduction
In FTS, the feed carbon monoxide and hydrogen are converted into a mixture of products. The reactor effluent consists besides of unreacted feed gases of both gaseous and liquid products, which have to be separated in a gas and a liquid stream before analysis. Moreover, the gaseous reactor effluent can easily form condensates and solid deposits upon cooling making constant heating of the products necessary. Analysis of the unreacted feed gases and the products ranging from methane to waxes is complex and asks for specially configured analysis equipment. All of these aspects were dealt with in the design of the FTS unit used for this research
Besides the tailor-made specifications for Fischer-Tropsch Synthesis, the set-up is built to speed up the slow and laborious process of catalyst testing for screening and kinetics studies. Both screening of catalysts and kinetic studies require numerous experiments. Moreover, FTS catalysts and set-up require a long stabilization period and the GC analysis of the long chain products takes typically 45-50 minutes. The solution chosen in this set-up to overcome these problems is parallelization of the reactor section (Figure 2.1) to increase the throughput. An important side-effect of parallelization of reactors, especially for kinetic studies, is the opportunity to design and perform series of experiments fixing a set of conditions, while varying another one.
Feed & Mixing
Gas supply Liquid
Vent
Flow division
Six-flow reactor
Wax separation
Analysis
Vent Vent Monolith reactor H2 CO ArFTS Experimental Set-Up
2.2 Set-up
The experimental set-up designed and built for this research project has six parallel fixed bed reactors with a continuous flow configuration. Parallel to one of the six fixed-bed reactors a reactor suited for testing monolithic catalysts is placed. An important aspect of the parallel concept is the possibility to share a large part of the equipment such as mass flow controllers for reactant mixing and analysis equipment. Other parts such as mass flow controllers dosing the feed mixture to the reactors and gas-liquid separators are dedicated to a specific reactor.
Figure 2.3 provides a detailed scheme experimental set-up. The feed synthesis gas mixture, including the internal standard argon, is composed from pure gases using Bronkhorst Hi-Tec mass flow controllers (MFC1-3). Additionally, a fourth mass flow controller (MFC 4) is available to supply another extra gas, for instance ethylene, to the synthesis gas mixture. The mixed synthesis gas is divided over the six reactors using six individual mass controllers (MFC 7-10). The excess of the mixed synthesis gas mixture is sent through a backpressure controller (BPC3). Besides the feed of gas mixtures, the option to supply liquids to the reactors is available using 6 Bronkhorst liquiflow liquid mass flow controllers (LMFC 1-6). The liquid feed system offers the opportunity to operate either in trickle-flow regime by adding a component like for instance tetradecane or under super-critical conditions with pentane or hexane as solvents.
The mixed and dosed flows are led to six-flow reactor and the monolith reactor. The monolith reactor is placed in parallel to the sixth fixed bed reactor and cannot be operated at the same time as this sixth reactor. The six-flow reactor and the monolith reactor, together with the high temperature and pressure gas-liquid separators (HT1-6) hot traps, the back pressure controllers (BPC 4-9) and a Valco flow through 8-way selection valve (EWSV 1) are located in a large oven to ensure isothermicity and prevent plugging of the lines by deposited wax (Fig 2.3). The large surrounding oven is operated at a constant temperature of 175 oC.
Figure 2.2: Detailed flow scheme of the FTS six-flow reactor set-up including a list of symbols FS 1 PI VE NT 62 .5 80 TWV2 62. 5 H2 (net ) TWV3 TWV4 62. 5 TI TI TI TI TI TI TI TI TC P TI A PI 80 TWV1 62 .5 50 50 50 50 50 80 80 OT 3 OT 4 OT2 OT 1 SR V 5 CV 3 CV 4 CV 5 CV 6 CV 7 CV 8 EH TI WCU CO (remote)200 200 Ar PI PI PI PI PI P MV8 Li qu id VEN T MV7 MV6 MV5 MV4 MV3 SV 1 MV1 4 MV9 MV1 3 MV1 2 MV1 1 MV1 0 CV 9 CV 1 0 CV 1 1 CV 1 2 BP C 3 P BP C1 P N2 MV 1 55 PI CV 1 CV 2 SR V6 SR V7 SR V8 SR V9 SR V1 0 SR V1 1 PR 1 P SV 2 PR 2 P SR V2 SR V3 SV 3 PR 3 P SR V4 SV 4 PR 4 P Liquid container MV2 SV 5 PI PI PI PI PI VEN T CV 13 VEN T SR V1 BP C2 HT 1 H T 2 HT 3 H T 4 H T 5 H T 6 BP C4 P BP C5 P BP C 6 P BP C 7 P BP C8 P BP C 9 P CT1 CT2 CT3 CT4 CT5 CT6 Re fr ig er a tor Ov en EW SV 1 EW SV 2 TI GC /FID /TC D FM 1 CT7 fw v 6 CV 1 4 CV 1 5 CV 1 6 CV 1 8 C V 17 C V 19 CV 20 fw v 5 fw v 4 fwv 3 fw v 2 fw v 7 VEN T AOV6A-B AOV1A-B AOV2A-B AOV3A-B AOV4A-B AOV5A-B 50 TI TI TC P TI A TI TW V 5 TW V 6 TI A TC P N2 (net ) TI TCP TI TI A TI TC P fw v 1 62. 5 50 50 50 50 50 50 50 1 1 1 1 1 1 filt e r 1 fi lt e r 4 fi lt e r 3 filt e r 2 MF C 2 MF C 4 MF C 3 MF C 1 LMF C 6 LMF C 5 LMF C 4 LMF C 3 LMF C 2 MF C 5 MF C 6 MF C 10 MF C 9 MF C 8 MFC 7 N2 LMF C 1 H P LC pu m p
FTS Experimental Set-Up
The six-flow reactor oven (Figure 2.4) consists of six Hastelloy reactor casings inside an aluminium mantle. The casings have an internal diameter of 6 mm and a length of 18 cm in which quartz or ceramic inserts can be placed containing a catalyst bed (Figure 2.5). The inserts have an internal diameter of 4 mm. Viton O-rings seal the reactors preventing leakage to the outside or bypassing around the inserted tube. Thermocouples (1mm thickness) are inserted from the top of the reactor sticking into the catalyst bed. The reactor is heated by six heating elements inside the aluminum block that ensures an even heat distribution.
In the monolith reactor monolithic catalysts with a diameter of 1 cm (Figure 2.6) can be tested under gas phase feed conditions. Gas-liquid operation under Taylor flow conditions is not possible as much too high liquid flow rates are needed to reach the right liquid velocity window. The monolithic catalysts are placed in spacers to prevent bypassing and to ensure easy exit of the gas and liquid products from the bottom of the monolith (Figure 2.7). Thermocouples can be inserted in the reactor from the top and from the bottom of the reactor for temperature measurement. The reactor is heated by a heating element inside an aluminum mantle surrounding the reactor. The large heat capacity and the high thermal conductivity of the aluminium mantle ensure an even heat distribution.
The heavy waxes are separated from the rest of reactor effluent in a hot trap at reaction pressure. The temperature in the hot trap is kept constantly at 175 oC by the large surrounding oven. The collected waxes are drained regularly through two consecutive Valco on/off valves (AOV1-6 A-B) into a flask (Figure 2.8) and held at atmospheric pressure and 175 oC. The atmosphere in the flask is inertized with nitrogen to prevent oxidation of the collected waxes and build-up of unreacted synthesis gas.
pressure reducer
manual operated on/off valve
mass flow controller
electrical heater fixed-bed microreactor
electrical heated aluminum block
three-way valve
gas-liquid separator air-operated valve
liquid collection vessel back pressure controller
heated line
pressure indicator
temperature indicator
temperature indicator alarm
oxygen trap PI
TI
programmed temperature controller solenoid valve
filter
P P
FS
water cooling unit
check valve
flow sensor
TIA
TCP
50 maximum pressure/bar
safety relief valve gas cylinder
liquid storage container
eight-way selection valve
flow meter FM
high pressure liquid pump four-way valve
Figure 2.4: Photo of the six-flow reactor oven Figure 2.5: Quartz and ceramic reactor
insert with Viton seal O-ring inside two of the reactor casings
Figure 2.6: Photo of monolithic catalysts with and without catalyst washcoat (1 cm diameter)
