Chemical Reaction EngineeringâII

45 Estimation of Platinum Catalyst Requirement for Ammonia Oxidation Downloaded by UNIV OF MASSACHUSETTS AMHERST on March 1, 2016 | http://pubs.acs.org Publication Date: June 1, 1975 | doi: 10.1021/ba-1974-0133.ch045

DONALD ROBERTS Department of C h e m i c a l Engineering, University of Birmingham, Birmingham B15 2TT, E n g l a n d G. RICHARD GILLESPIE Engelhard Industries, 430 Mountain Ave., M u r r a y Hill, N. J. 07974 Despite the temperature gradients present in the Pt catalyst pad used in NH oxidation and the usual surface rearrangement of the platinum, measured mass transfer coefficients for stacked metal screens can be used to estimate the amount of catalyst required for satisfactory operation of a commercialNH burner. Burner design as a mass transfer operation using a transfer unit concept is recommended. Considerable reduction in platinum inventory, from previously accepted levels, seems possible. Calculations indicate that operation at a higher mass velocity or with closer mesh screens woven from finer wire would reduce catalyst requirement although other factors might be adversely affected. 3

3

t has long been argued that the catalytic oxidation of ammonia w i t h air is a mass transfer controlled reaction. U n d e r typical operating conditions of industrial ammonia burners, Oele ( I ) demonstrated that the kinetic stoichio metric 0 / N H ratio is exceeded, ensuring an oxygen-rich atmosphere at the catalyst surface w i t h ammonia diffusion as the limiting reaction step. Such a situation is probably necessary to conduct this reaction satisfactorily. The stoichiometry (not intended to represent mechanism) of formation of the three alternative products can be represented as follows:

I

2

3

4 N H + 3 0 -> 2 N 4- 6 H 0 4 N H + 4 0 -> 2 N 0 + 6 H 0 4 N H + 5 0 -> 4NO + 6 H 0 3

2

2

3

2

2

3

2

2

2

2

E q u i l i b r i u m conditions are too far to the right to present any problem. The desired product is usually nitric oxide, and it is significant that this represents the highest oxidation state of the three alternatives. Its preferential formation w i l l probably be encouraged by a high relative proportion of oxygen to ammonia at the catalyst surface. Irrespective of the chemical mechanisms involved and considering nitric oxide as the desired product, yield losses can occur because of: (a) insufficient catalyst, and (b) poor selectivity; it is constructive to distinguish the two. B y distinguishing the former it is possible to make more meaningful studies of 600

In Chemical Reaction Engineering—II; Hulburt, Hugh M.; Advances in Chemistry; American Chemical Society: Washington, DC, 1974.

45.

ROBERTS

AND

GILLESPIE

Platinum

Catalyst

601

the effect of operating variables on selectivity. This, we contend, we are now able to do. This paper demonstrates that recent correlations of mass transfer coefficients enable one to estimate the amount of catalyst required for any specific duty.

Downloaded by UNIV OF MASSACHUSETTS AMHERST on March 1, 2016 | http://pubs.acs.org Publication Date: June 1, 1975 | doi: 10.1021/ba-1974-0133.ch045

The Catalyst

System

T h e traditional catalyst for ammonia oxidation is a stack of woven wire screens, typically of 9 0 P t / 1 0 R h alloy or occasionally of some other platinum alloy. A common standard for the screens is 80 meshes per linear inch, con structed of 0.003-inch diameter wire although other standards are i n use. The number of screens constituting a catalyst pad varies from as few as three at near atmospheric pressure to as many as 20 or 30 at around 8 atm. A n increase in mass velocity usually accompanies an increase i n operating pressure. D u r i n g use, substantial catalyst is lost; despite this the leading wires i n the pad thicken considerably because of cauliflower-like growths; these growths progressively weaken the structure of the screen, necessitating eventual replace ment. Photographs published by Schmidt and Luss (2) represent the appear ance of used catalyst. Clearly the mechanical support of such flimsy catalyst structures requires careful consideration. A suitable support should have a fine structure able to spread the mechanical load evenly across the full w i d t h of the reactor. It should be deep enough to prevent bulkier elements of the reactor structure from interfering w i t h flow through the catalyst pad, and the support system and catalyst pad together should present a sufficient and uniform resistance to flow to smooth out minor flow variations i n the entering gas stream. If installed catalyst weight is to be critically designed, it w o u l d be a useful safety precau tion if the support system could catalytically react w i t h any residual ammonia, not necessarily to the desired product N O . A m m o n i a might otherwise escape to cooler parts of the plant and form potentially dangerous ammonium nitrite. Random Pack system has been recently introduced (3, 4) and seems to possess these properties. Claims (5) that its use enables satisfactory operation of an ammonia burner w i t h as little as 4 0 % of the normal catalyst charge appear to suggest that earlier designs relied upon excess catalyst as support. The high cost of installing an excess charge of a noble metal catalyst argues the need for a design procedure to estimate catalyst requirement in a system where the flow through the pad is uniformly distributed. Mass Transfer

Coefficients

Apelbaum and T e m k i n (6) and Oele ( I ) tried to compare the rate of ammonia oxidation on stacked screens w i t h the rate of ammonia mass transfer, although they used a heat transfer correlation for flow perpendicular to a single cylinder as the source of their mass transfer coefficient. Nowak ( 7 ) , i n a similar comparison, had the advantage of heat transfer coefficients for stacked screens reported by L o n d o n et al. (8) and the mass transfer coefficients for single screens reported by G a y and Maughan ( 9 ) . A l l three studies suggested that the limiting mass transfer rate and the reaction rate were similar. Since then, more relevant measurements of mass transfer coefficients to stacked metal screens were made by Satterfield and Cortez (10) and by Shah ( I I ) ; independently both suggested C h i l t o n - C o l b u r n type correlations based upon wire diameter as the critical dimension. They both had to base


602

CHEMICAL REACTION ENGINEERING

II


their correlations on mass velocities corrected for the reduction to flow area presented b y the screen; Satterfield a n d Cortez chose G/t the average mass velocity through a single screen, and Shah chose G/y the maximum mass velocity. T h e correlations otherwise agree generally, although the higher Reynold's number range of Shah's work a n d the greater number of screens i n the stack makes his measurements more relevant to h i g h pressure ammonia oxidation conditions a n d possibly less affected b y axial diffusion. T h e correla tions can be summarized as follows i n terms of Reynold's number based upon superficial velocity: Satterfield and Cortez: one to three screens; Shah:

0.4 < N < 9 e j = 0.865 (ATRe/e) * Re

-0

D

one to five screens;

648

107 TJD = 0.644 (iVRe/γ)- · These two correlations cover the range of interest i n ammonia oxidation. 3
•


ο

CO

ΓΓ 9 5

ΗΝΟ3

PRODUCTION

INSTALLED

LU > ζ ο ο

FINAL

90

178 T O N S / D A Y

CATALYST

CATALYST

WEIGHT

WEIGHT

144-8 OZ 55-2 OZ

20

60

AO

80

DAYS

Figure 2.

Nitric acid plant production run with reduced weight catalyst pad (with Random Pack and Degussa Getter)

Screen Characteristics. T h e only screen characteristics i n Equation 5 are d, γ, and a; to reduce catalyst requirement, it w o u l d seem necessary to decrease d and γ while increasing a, although these parameters are interdependent. E x tracting superficial mass velocity and gas properties from the equation, (HTU)

oc

γ

0.43/

α

If a is f t / t r o y oz of catalyst, a oc l/d, and since γ ~ (1 ~ nd) , the catalyst requirement i n troy oz oc d - (1 — nd) . This function is plotted i n Figure 3 for different mesh sizes. Although the curves pass through a maximum, it is at too h i g h a wire diameter to be of practical interest. Obviously, the weight of catalyst required can be reduced significantly b y moving towards a closer mesh screen woven from finer wire. A similar analysis based on the Satterfield and Cortez correlation ( E q u a tion 1 ) , leads to the approximate prediction: 2

2

1

57

0

86

/4

\ 0.704

/

catalyst requirement in troy oz α d l - — nd v l -f- n d J 16iS

2

2

a function w h i c h has a form similar to that i n Figure 3, thus leading to the same qualitative conclusions. The characteristics of screens used b y Satterfield and Cortez (10) and b y Shah (11) are plotted i n Figure 3 as a guide to practical combinations. Process Variables. P R E S S U R E . F r o m Equation 5 w e w o u l d expect changes i n operating pressure to have little direct effect on catalyst requirement. H o w ever, higher pressures allow a higher mass velocity to be used w i t h a consequent reduction i n catalyst requirement. EXIT GAS TEMPERATURE. T h e only property i n E q u a t i o n 5 w h i c h is sig nificantly affected b y temperature is the viscosity of the gas mixture. F o r minor variations around 900°C, therefore, ( H T U ) oc Τ" · ; that is, the effect is minor. 0

28


608

CHEMICAL

REACTION

II

MASS VELOCITY O F GAS. Although Equation 5 indicates that ( H T U ) oc G - , a reduction i n gauze requirement results from an increase i n mass velocity; this is because for a specific reactor duty i n terms of ammonia to be burned, the reactor cross sectional area can be reduced i n inverse proportion to the mass velocity. Thus, the gauze requirement oc G ~ - . A N O T E O F C A U T I O N . W h i l e it is tempting to suggest immediate m o d i fications to ammonia oxidation catalysts and ammonia burner operating con ditions i n line w i t h these observations, it would be wrong to do so. M a n y factors are involved, and the only one examined here is our estimate of the weight of catalyst required. For example, one should not assume that catalyst requirement can be reduced by increasing mass transfer rates without possible detriment to selec tivity. If at some point within the bed, bulk gas phase partial pressures of ammonia and oxygen are PNH3 ^ P02 respectively and are diffusing toward the catalyst surface where the corresponding partial pressures are P * N H and ρ * , then the mass transfer rate of oxygen/unit area of catalyst surface is 0

5 7

0


ENGINEERING

4 3

a n (

3

θ 2

kgo(po - p*o ) = 1.25 k (PNH, - P*NH ) — 1.25 2

2

g

3

k psn g

t

since operating conditions are such that ammonia mass transport is limiting and assuming the yield is 1 0 0 % (for simplicity).

WIRE

Figure 3.

DIAMETER

χ 10$,

INS

Effect of screen characteristics on mass transfer capacity of screens


45.

ROBERTS

Platinum

A N D GILLESPIE

Thus p*o ^ Po - 1.25 2

609

Catalyst

pm

2

h

KgO

If the mass transfer coefficients are doubled (e.g., b y increasing mass velocity), their ratio k /k w o u l d be unchanged, and p * w o u l d effectively remain the same. T h e mass flux of both reactants to the surface however, is doubled, so that the reaction must proceed at twice the previous rate. This probably could not happen without a marked increase i n Ρ*ΝΗ > even though it may remain small relative to p Thus, an increase i n mass transfer coefficient results i n a reduction i n P*O /P*NH > representing probable deterioration i n the reaction conditions at the catalyst surface and resulting from a modest move away from mass transfer control. F u l l implications of improvements i n mass flux to the surface w i l l become apparent only w i t h better understanding of the complete chemical mechanism. g

g0

0 2

3


N H 3

3

2

Catalyst

Loss

N o study of ammonia oxidation catalysts w o u l d be complete without some discussion of catalyst loss. Nowak (7) suggested that direct vaporization of P t 0 is significant, but Schmidt and Luss (2) pointed out that Nowak's cor relation differs from the observed loss rate b y a factor of about four. Edwards et al. (12) explain the higher loss rate as a result of temperature fluctuations. Since precious metal loss rate is a convex function of temperature, higher losses occur w i t h a fluctuating temperature than w i t h the same average steady tem perature. Gillespie a n d Kenson (5) claim that the Random Pack catalyst sys tem, containing as much as 6 0 % less P t / R h than a traditional catalyst p a d , results i n a platinum loss rate w h i c h is 25 to 3 0 % less than normal. These reduced loss rates are accompanied b y reports of reduced catalyst temperature fluctuations. A move towards thinner wire catalyst structures might accentuate these temperature fluctuations. Diffusion of P t 0 into the gas stream is one step i n such a process, but it may be a limiting one. If P t 0 diffuses more slowly than ammonia, then a catalyst pad correctly designed for ammonia consumption might contain insuffi cient surface for the platinum loss rate to reach its full potential. Additional catalyst w o u l d improve ammonia consumption very little but might a d d ap preciably to P t 0 vaporization. This argument is offered only tentatively. T h e catalyst loss mechanism is clearly a complex problem w h i c h deserves more thorough investigation. 2

2

2

2

Conclusions The financial penalties incurred from excess p l a t i n u m / r h o d i u m catalyst i n an ammonia burner are severe, comprising the servicing cost of the a d d i tional capital and apparently a higher platinum loss rate than necessary. T h e installed catalyst weight should be the m i n i m u m required for satisfactory operation. The equations presented here can be used to estimate catalyst require ments for any high pressure installation; if reduced weight catalyst pads are used, uniform flow distribution and support must be ensured b y installing a combined catalyst-support system such as Random Pack. Former methods of estimating catalyst requirement took little note of process variation between different installations.



610

CHEMICAL REACTION ENGINEERING

II

W h i l e major doubts remain over the errors introduced by using isothermal mass transfer coefficients i n the presence of large temperature gradients, such doubts are important only for the leading screens of the catalyst p a d , where surface rearrangement of the metal should ensure that our equations otherwise underestimate the mass transfer capability of these screens. Our calculations suggest that operation of an ammonia burner at higher mass velocity or use of closer mesh screens woven from finer wire w o u l d reduce catalyst requirement although other factors more important to nitric acid plant operations might be adversely affected. A p e l b a u m and T e m k i n (6) suggested thinner wires and higher gas velocities i n 1948, but little progress has been made i n this direction. Such developments have been restricted by high manufacturing costs and fears that such screens w o u l d have insufficient strength and poor aging properties. Nomenclature a

surface area of unit quantity of 1 f t of screen (units according to choice ) diffusivity, f t / h r wire diameter, inches aged wire diameter, inches clean wire diameter, inches superficial gas mass velocity, l b s / h r - f t height of transfer unit, = N$ /j a (units according to choice) value of ( H T U ) at mean conditions (units according to choice) / factor for mass transfer, = (k MP/G)N mass transfer coefficient for N H , lb-moles/hr-ft -atm mass transfer coefficient for 0 , lb-moles/hr-ft -atm depth of catalyst pad (units according to choice) mean molecular weight of gas Reynolds number, = dfi/μ Schmidt number, = ν/Ό number of transfer units mesh size, inches" total pressure, atm partial pressure of N H , atm partial pressure of N H , atm partial pressure of N H at catalyst surface, atm partial pressure of 0 , atm partial pressure of 0 at catalyst surface, atm partial pressure of N H exiting catalyst p a d , atm partial pressure of N H entering catalyst pad, atm N H loading of 1 f t of reactor cross section, short tons/day-ft temperature, °C mean film temperature, °C wire temperature, °C gas temperature, °C mole fraction of N H i n feed 2

D d d d G (HTU) (HTU) ; k k L M N N (NTU) η Ρ ρ PNH & c

D

g

g0

R e

Sc

3

P*NH3

Po ρ* p Pi R Ύ_ Τ T T χ

2 θ 2

E

N H 3

s

B

Greek ε γ μ ν

2

2

2/3

D

g

2/s

Sc

2

3

2

2

1

3

3

3

2

2

3

3

3

2

2

3

porosity of single screen m i n i m u m fractional opening of single screen gas viscosity, l b s / f t - h r kinematic gas viscosity, f t / h r 2


45. ROBERTS AND GILLESPIE Platinum Catalyst

611


Literature Cited

1. Oele, A. P., Chem. Reaction Eng., Meeting Europ. Fed. Chem. Eng., 12th, Amsterdam (1957) 146. 2. Schmidt, L. D., Luss, D.,J.Catalysis (1971) 22, 269. 3. U.S. Patent 3,660,024 (1972). 4. Canadian Patent 916,396 (1972). 5. Gillespie, G. R., Kenson, R. E., Chem. Tech. (1971) 1, 627. 6. Apelbaum, L., Temkin, M.,J.Phys. Chem. (USSR) (1948 ) 22, 179. 7. Nowak, E. J., Chem. Eng. Sci. (1966) 21, 19. 8. London, A. L., Mitchell, J. W., Sutherland, W. Α., J. Heat Transfer (1960) 199. 9. Gay, B., Maughan, R., Int. J. Heat Mass Transfer (1963) 6, 277. 10. Satterfield, C. N., Cortez, D. H., Ind. Eng. Chem., Fundamentals (1970) 9, 11. Shah, Μ. Α., Ph.D. thesis, University of Birmingham, England (1970). 12. Edwards, W. M., Worley, F. L., Luss, D., Chem. Eng. Sci.(1973) 28, 1479 RECEIVED

January 2, 1974.


Chemical Reaction EngineeringâII

Recommend Documents

Chemical Reaction EngineeringâII