The Selective Oxidation of n-Butane to Maleic Anhydride in a Catalyst Packed Tubular Reactor

GBH Enterprises, Ltd.

Process Engineering Guide:
GBHE-PEG-RXT-815

The Selective Oxidation of n-Butane to
Maleic Anhydride in a Catalyst Packed
Tubular Reactor
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Process Engineering Guide:
The Selective Oxidation of n-Butane to Maleic Anhydride in a
Catalyst Packed Tubular Reactors
CONTENTS

0

INTRODUCTION

1

n-BUTANE OXIDATION

2

REACTION KINETICS

3

HEAT AND MASS TRANSFER PARAMETERS

4

NON-ISOTHERMAL, NON-ADIABATIC REACTOR MODELLING

5

USE OF THE REACTOR MODEL IN OPERABILITY AND DESIGN
STUDIES

6

BIBLIOGRAPHY

7

NOMENCLATURE


TABLES

1 SCOPE OF EXPERIMENTS - BUTANE KINETICS AND NONISOTHERMAL RUNS
2 KINETIC PARAMETERS FROM THE DIFFERENTIAL REACTOR

3 KINETIC PARAMETERS FROM THE INTEGRAL REACTOR
4 STRUCTURAL PERMEABILITY DETERMINATIONS FOR THE nBUTANE CATALYST
5 HETEROGENEITY OF THE REACTOR MODEL
6 COMPARISON OF OBSERVED AND PREDICTED PARAMETERS

7 BASE CASE AND OPTIMISED RESULTS FOR YIELD IMPROVEMENT

DOCUMENTS REFERRED TO IN THIS PROCESS ENGINEERING GUIDE


0

INTRODUCTION
This case study divides into two parts:
(a)

The synthesis and testing of the reactor model, and

(b)

The exploitation of the model for yield improvements.

1 n-BUTANE OXIDATION
At present the majority of Maleic anhydride is produced by the oxidation of
benzene. However, process economics and environmental factors suggest that
n-butane is the feedstock of the future. In comparison with the historic, but
intrinsically less efficient route for benzene, butane catalysts are less selective.
One method of improving on existing selectivity is to employ reaction engineering
principles to optimize reaction yield.
The formation of Maleic anhydride from n-butane is accompanied by Maleic
anhydride decomposition and complete combustion of n-butane; the classic
series-parallel reaction scheme of 6.2.2 of GBHE-PEG-RXT-805.
(a)

Reaction I
C4H10 + 3.5 O2

(b)

C4H2O3 + 4H2O

Reaction II
C4H2O3 + m O2

(c)

(6 - 2m)CO + (2m - 2)CO2 + H2O 1 m 3

Reaction III
C4H10 + n O2 2* (6.5 - n)CO +

2* (n - 4.5)CO2 + 5H2O 4.5 n 6.5

The stoichiometric coefficients m and n for a particular catalyst are determined by
matching the above scheme to the observed product distribution.


2

REACTION KINETICS
Two types of laboratory reactor were used to determine the reaction
kinetics; a glass differential reactor and a steel integral reactor. The overall
experimental program is summarized in Table! 1. Also shown for
comparison are the conditions employed in the pilot plant experiments
running at commercial rates.

2.1

The Differential Reactor
Isothermality in the differential reactor was achieved by catalyst dilution
with glass beads in a 1:7 ratio and the conversions were usually limited to
a maximum of 8%.
Differential reactor experiments were designed to check the effect of pore
diffusion by using two particle sizes 0.7 mm and 7 mm. Their other
objective was to highlight the parallel Reactions I and III and quantify
reaction rates in terms of temperature and n-butane partial pressure (p1).
Because of the difficulty of feeding Maleic anhydride, no information could
be gained on the rate of the product degradation Reaction II or the
retarding effect of Maleic anhydride on the rates of Reactions I and III.


2.1.1 Pore diffusion effect
Figure 1 shows rates of Maleic anhydride (MSA) production as a function
of temperature for the two particle sizes at a relatively low level of nButane in the feed. Assuming that the crushed catalyst (0.7 mm) is
operating in the chemical rate-controlled regime, significant pore diffusion
is anticipated for 3 mm extrudates, employed commercially, for T > 370°C.

2.1.2 Kinetic parameters from the differential reactor
The parallel reactions given in Reaction I and Reaction III were described
by power law rate equations of the form:


The six parameters in this model can, in principle, be estimated from
measurements of Maleic anhydride, CO and CO2 in the product gas, together
with measurements of n-butane at the reactor inlet and outlet. This model led to
poorly determined and highly correlated values of k3 and a3. A simplified four
parameter model with a1 = a3 and E1 = E3 fitted the data equally well and led to
improved parameter estimates, although the standard errors on the reaction
rates varied between 10 to 30%, and a high degree of parameter correlation still
was evident. The results of the differential reactor experiments are given in Table
2.

2.1.3 Testing the differential reactor kinetics
To emphasize the essential feedback in reactor development, the
differential reactor kinetics were evaluated by incorporation into a reactor
model including independently determined pore diffusion and heat transfer
coefficients. This model was integrated and compared with experimentally
observed temperature and concentration profiles from a 4 meter pilot plant
reactor in Figure 2. Severe discrepancies are apparent. It is clear that
reaction rates are under-predicted at the front end of the reactor and
overestimated in the tail. This suggests that product inhibition may be
important.


2.2

The Integral Reactor

The failure of the differential reactor kinetics to scale up to pilot plant results
necessitated a rethink of the kinetics study. It was thought that experiments with
an integral tubular reactor would emphasize kinetics in the presence of Maleic
anhydride and improve the predictions of the reactor model, providing isothermal
conditions could be obtained in the laboratory study. It was decided to employ a
full scale reactor tube (25 mm inside diameter x 4 m length) containing
commercial size 3 mm extrudates for this purpose. The tube had several
intermediate sampling points and an axial thermowell for temperature
measurements and was contained within a molten salt bath. In the front 40% of
the reactor, the catalyst was diluted with inert-pellets in a 1:1 ratio, while in the
tail end region the dilution factor was 1:0.5 catalyst to inert. Concentration
measurements were taken only over the "isothermal" tail-end region.


2.2.1 Kinetic parameters from the integral reactor
In keeping with the spirit of GBHE-PEG-RXT-805 regarding model
parsimony, the previous 4 parameter model given by Equations (1) and (2)
was extended to the following 7 parameter form to encompass the product
degradation Reaction II, together with inhibition by Maleic anhydride (p2).
Reaction between adsorbed C4H2O3 and gas phase O2 led to:

Surface reaction between adsorbed C4H2O3 and adsorbed oxygen, but
KMP2 >> KOpO2, led to:

The product and reactant concentrations were well fitted to within 2 to 7%
standard error for n-butane conversions 60%. Parameter values, together with
their approximate 95% confidence intervals are summarized in Table 3.


In comparison with the differential reactor parameters in Table 2, relatively small
changes in k1 and a1 are observed, although confidence intervals are
considerably reduced. On the other hand, larger changes to E1 and k3 are noted.
The high value of E2 in relation to E1 clearly shows the disadvantage with respect
to selectivity of operation with large hot-spots.
No physico-chemical interpretation of the form of Equations (3) to (5) can be
attempted, since other quite different forms were found to fit the data almost as
well.
Non-kinetic means are needed to cast light on the true nature of the surface
reactions.


3
3.1

HEAT AND MASS TRANSFER PARAMETERS
Pore Diffusion
Pore diffusion effects were encountered above 370°C in the differential
reactor studies for 7mm extrudates. Since temperatures in the "hot spot"
region of a commercial reactor are in the region of 450°C, significant
diffusional modifications to both catalyst activity and selectivity on 3 mm
extrudates are anticipated.
The catalyst is formed by extruding microporous particles of vanadium and
phosphorous oxides formed by co-precipitation from either aqueous or
organic media. It consists of overlapping regions of precipitate and
pelleting pores within the range 0.01 to 1 o m diameter. The specific
surface, as measured by the BET method, is approximately 11 m2/gm, and
the porosity, as determined from apparent density and pore volume
measurements, is 0.35.
The treatment of effective diffusion within two ideal types of pore system
(the simple unimodal distribution and the bimodal distribution) was
considered in GBHE-PEG-RXT-805. These analyses do not strictly apply
to the non-ideal pore structure found in the catalyst under study.
Nevertheless, for the determination of the structural specific permeability
φ, Equation (18) was employed with r in Equation (20) being interpreted
as the volume-averaged radius from mercury intrusion measurements. De
was measured by a pulse-broadening technique using helium pulses in
nitrogen carrier gas passing through a packed column of catalyst. The
effect of axial dispersion was "subtracted" by carrying out similar
experiments using near-identical non-porous glass particles (Ref. [1]). The
gases were then interchanged and the experiments repeated. Estimates
of φ are shown for both cases in Table 4.
It is difficult to say within the uncertainty limits on φ, imposed by the
particular technique employed, whether φ depends on the diffusing gas. It
appears to be independent of the carrier gas velocity, however. Given the
assumptions made in the determination of | from Equations (18) to (20) of
GBHE-PEG-RXT-805, the overall result is reasonably satisfactory. A value
of φ = 0.12 is chosen to determine the binary diffusitives of reactant nbutane and product Maleic anhydride in the pseudo-component "air".


3.2

Radial Heat Transfer
Heat transfer parameters λr , eff, hw and U were estimated from equations
similar to those in GBHE-PEG-RXT-810 for the following base-case
conditions:


4

NON-ISOTHERMAL, NON-ADIABATIC REACTOR MODELLING
The stage is now set for an assault on predicting the performance of a fullscale reactor tube operating under commercially relevant i.e. nonisothermal, non-adiabatic, conditions.
In order to provide a compromise between the degree of physico-chemical
mechanistic detail needed and the requirement for mathematical
tractability, a one-dimensional heterogeneous model was chosen for
evaluation. This model accounts for the following gradients:
(a)

Interparticle axial temperature, total pressure and partial pressure
gradients.

(b)

Interfacial temperature and partial pressure gradients.

(c)

Intraparticle temperature and partial pressure gradients.

A full description of the model and a summary of the method of numerical
solution is presented in Ref.[2].

4.1

The Base Case
In the base case, represented as Case No. 1 in Table 6, detailed
temperature and concentration measurements were made at points along
the reactor tube, thereby providing data for a stringent test of the model.
Figure 3 shows axial temperature measurements and observed gas
compositions alongside model predictions. In spite of the simplified
treatment of interparticle heat transfer, the overall agreement is good.
More detailed observations of the model, shown in Table 5, highlight the
balance of macroscopic and microscopic gradients.


It would appear that neither intraparticle nor inter-phase temperature gradients
develop to any significant extent, even for this set of highly exothermic reactions.
If it is assumed that a parabolic radial temperature profile exists, then the
average temperature in the cross-section can be written in terms of the axial
temperature Tax and the salt bath temperature Ts by the equation:

At the “hot spot”, Tax = 403°C, it follows from Equation (6) that T av = 389°C.
Thus, a significant radial temperature gradient across the bed would appear to
exist in this particular case, which corresponds to a low throughput.
Nevertheless, the one-dimensional model is still able to describe the observed
product distribution.


Very noticeable intraparticle concentration gradients develop, which lead to a
lowering of catalyst effectiveness factors for reactant degradation Reaction I and
III, but a raising of effectiveness for the product destructive Reaction II in Clause
1. The effectiveness factors also vary significantly along the bed, which
necessitates a detailed treatment of the intraparticle reaction-diffusion problem,
since this bears not only on activity but also selectivity.
Yield to Maleic anhydride inevitably falls along the bed, but its rate of decline is
exacerbated by both pore diffusion and radial heat transfer. Clearly, there is
considerable scope for improving catalyst selectivity.


4.2

A Wide Range Comparison
The reactor model was also tested against other data spanning a wide
range of operating conditions, such as different mass flow rates, coolant
temperatures, and butane concentrations in the feed, tube length and
diameter. Agreement between model predictions and experimental data is
satisfactory over the entire range of operating conditions of interest for
commercial scale operation (see Table 6).


5

USE OF THE REACTOR MODEL IN OPERABILITY AND
DESIGN STUDIES

5.1

Defining Operability Limits
It is well known in hydrocarbon partial oxidations that reactors can become
unstable. In particular, these reactors are prone to temperature runaways,
a condition for which a small change In either salt bath temperature, feed
throughput or hydrocarbon concentration in the feed causes the "hotspot"
temperature to tend to rise uncontrollably. Temperature runaways can
lead to immediate and costly plant shutdown, catalyst replacement or
even, in extreme cases, mechanical failure and tube replacement.
Nevertheless, economic considerations dictate the necessity for high
product yields and this invariably means operating the reactor as close to
runaway as is considered practicable. To achieve the desired economic
targets it is necessary to determine the limits of operability and this is
preferably done through a judicious combination of modeling and
experimental work.
Figure 4 displays operability (or runaway) limits for the oxidation of nbutane as a function of the tube flow rate, the inlet concentration of butane
and the salt bath temperature, as calculated by the reactor model. The
region of 'safe' operation lies below the surface. In planning production
rate changes, the diagram is valuable in determining the admissible
combinations of the operating variables.
In more detailed studies, the influence of salt flow hydrodynamics and
heat transfer may be considered, leading to improved baffling and
recirculation through the reactor shell.


5.2.

Increasing Reactor Yield through Optimization
For a 30,000 t/year Maleic anhydride plant with a total capital investment
of 37.5 x 106 $, running with a base case reactor yield of 91 wt%, it has
been estimated (Ref. [3]) that a 4 wt% yield improvement would increase
profits by 1.35 x 106 $ per year, raising the return on investment (ROI)
from 29.6 to 33.2%.
In relation to the maximum theoretical yield of 168.9 wt%, current yields of
91 to 92 wt% reflect the poor selectivity of n-butane oxidation catalysts.
While the search for new, more selective catalysts continues, the chemical
engineer faces the challenge of improving present yields through
innovative design.


5.2.1 Dual catalyst systems
The maximum achievable yield in a 5 m reactor isothermally, is about 98
wt% and occurs at a salt bath temperature of 392 to 393°C. Thus, a
valuable 6 to 7 wt% yield increase could be realized by removing the "hotspot" limitation. In practice, because this is not entirely possible, a target 4
to 5 wt% might be realistically set.
The simplest and certainly the cheapest way of progressing towards such
a target is to employ a dual catalyst system, that is, a dual support system
offering high heat transfer to pressure drop at the front of the reactor
where the "hot spot" is, and high activity in the tail part of the reactor. The
questions for the designer should be addressed at optimizing the size and
shapes and lengths of the respective packed zones. A detailed reactor
model can answer these questions fairly quickly.
Table 7 compares the current base case with an optimized solution
specifying the optimal shapes and sizes of packing and their packed
lengths, with no incurred pressure drop penalty. The axial temperature
profiles are compared in Figure 5.
In relation to the conventional fixed bed packed with uniform sized pellets
along its entire length, the dual system displays a considerably flattened
temperature distribution, while at the same time providing a 4%
enhancement in yield. It is estimated that even a modest 10°C reduction in
the "hot-spot" temperature may significantly increase the lifetime of the
catalyst and its time "on stream" yield.
Other optimization measures have been reported elsewhere (Ref. [3]).


6

BIBLIOGRAPHY

[1]

Cresswell, D.L. and Orr, N.H., "Measurement of Binary Gaseous Diffusion
Coefficients within Porous Catalysts" from Residence Time Distribution
Theory in Chemical Engineering, ed A. Petho and R.D. Noble, Verlag
Chemie, Weinheim, p 41 (1982)
IC 07039/C Cresswell, D.L. Simultaneous sorption and diffusion within
adsorbent granules using pulse chromatography (Nov 1986).

[2]

Sharma, R.K., Cresswell, D.L. and Newson, E.J. "Selective Oxidation of
Benzene to Maleic Anhydride at Commercially Relevant Conditions"
ISCRE 8 p 353 Edinburgh, (Sept 1984) I.Ch.E. Symp Series No 87.

[3]

Wellauer, T.P., Cresswell, D.L., and Newson, E.J., "Optimal Policies in
Maleic Anhydride Production through Detailed Reaction Modelling" to he
presented at ISCRE 9 Philadelphia (May 1986).


7 NOMENCLATURE


DOCUMENTS REFERRED TO IN THIS PROCESS ENGINEERING GUIDE
This Process Engineering Guide makes reference to the following documents:

ENGINEERING GUIDES

GBHE-PEG-RXT-808

Solid Catalyzed Reactions
(referred to in Clause 1, 2.2.1 and 3.1)

GBHE-PEG-RXT-810

Heterogeneous Reactions, Gas Solid
Systems (referred to in 3.2)


The Selective Oxidation of n-Butane to Maleic Anhydride in a Catalyst Packed Tubular Reactor

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