the selective oxidation of n-butane to maleic anhydride in a catalyst packed tubular reactor

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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 NON-ISOTHERMAL RUNS

2 KINETIC PARAMETERS FROM THE DIFFERENTIAL REACTOR

3 KINETIC PARAMETERS FROM THE INTEGRAL REACTOR

4 STRUCTURAL PERMEABILITY DETERMINATIONS FOR THE n-BUTANE 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 C4H2O3 + 4H2O (b) Reaction II

C4H2O3 + m O2 (6 - 2m)CO + (2m - 2)CO2 + H2O 1 m 3 (c) 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 n-Butane 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 HEAT AND MASS TRANSFER PARAMETERS 3.1 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 n-butane 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 full-scale reactor tube operating under commercially relevant i.e. non-isothermal, 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 n-butane 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 "hot-spot" 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

Technology

process economics

particular catalyst

nbutane catalyst

butane catalysts

situ exsitu sulfiding

selective oxidation

butane kinetics

complete combustion