Manufacture of Dimethylamine Using Zeolite Catalyst

Manufacture of Dimethylamine Using Zeolite Catalyst Y. Ashina, T. Fujita, M. Fukatsu, K. Niwa, J. Yagi Central Research Laboratory, Nitto Chemical Industry Co. Ltd., Tsurumi-ku, Yokohama, 230, Japan Utilization of a zeolite catalyst for the manufacture of dimethylamine from ammonia and methanol was investigated. Among the zeolites tested, mordenite exhibited a high dimethylamine selectivity which exceeded the equilibrium level. Selevtivity and activity of mordenite were optimized by carefully controlling the alkali metal content of the catalyst. Treatment of the catalyst with high temperature steam further improved the selectivity. Thus modified mordenite gave a commercially feasible activity at lower range of temperature i.e., 290-340'C, while retaining the selectivity more than double the conventional level. The catalyst life could be approximately one year with no regeneration in this temperature range. The use of this catalyst has made it possible to reduce significantly the quantity of the process stream thereby presenting potentialities for various cost savings. INTRODUCTION Methylamines are commercially manufactured in a vapor phase heterogeneous reaction of methanol and ammonia using a solid acid catalyst such as silicaalumina at a temperature around 400'C and under a pressure around 20 atm. Major two-molecule reactions involved are as follows; K(400'C) 41.8 187 324 0.129 4.48 1. 73

MeOH + NH3 = MeNH2 + H20 MeOH + MeNH2 = Me2NH + H20 MeOH + Me2NH = Me3N + H20 NH3 + Me3N = MeNH2 + Me2NH 2MeNH2 = Me2NH + NH3 2Me2NH = MeNH2 + Me3N

(1)

Dehydration between methanol and ammonia or methylamines and transmethylation of methylamines and ammonia take place consecutively and simultaneously, ultimately reaching to the equilibrium state. In the coventional catalyst system, the selectivities of methylamines are determined thermodynamically in a moderate methanol conversion range, consequently a considerable amount of mono and tri substitutes are unavoidably formed along with dimethylamine. In Japan, however, the demands for monomethylamine and trimethylamine are very small in comparison with that for dimethylamine so that the most part of monomethylamine and trimethylamine formed are recycled back to the reactor to be reused for transmethylation. The basic process in general is shown in Fig.1.

779

780 (CA-5-4)

r-----,----,-------r--MMA

MeOH

NH3 Fig.1 Methylamine production process.

(2)

This process is characterized by a large-scale and cumbersome recovery system due to the use of excess amount of ammonia necessary to have a favorable product distribution and to suppress the formation of impurities and due to the complex distillation system because of the existence of trimethylamine which forms an azeotropic mixture with ammonia and other amines. The utility cost of this process therefore becomes very high; incidentally 22-198KWh of electlicity and 7.5-13 tons of steam for one ton of methylamine were reported for similar processes to Fig.1, though the figure actually is lower than these for most of the manufacturers today through their energy saving efforts such as computer optimization of distillation system. (3) ,(4) Impelled by such circumstances, methylamine producers have made various trials, including the use of zeolite catalyst, to realize a non-equilibrium product distribution in favor of dimethyla~ine thereby to reduce utility cost and/or to expand the plant capacity. (5)-(11) We have studied the utilization of zeolite catalyst for the selective manufacture of dimethylamine and have put it into commercial use with a successful outcome. This paper describes the outline of our study to present one of the instances where the unique characteristic of zeolite catalyst is commercially utilized. EXPERIMENTAL Catalysts tested include naturally ocurring mordenite (JAPAN), synthetic mordenite (Norton 200H), Y zeolite(Shokubai Kasei) and silica alumina (Shokubai Kasei High-alumina). Cation content of the zeolites were controlled either by immersing the catalyst in 0.5 to 2 equivalent of aqueous ion solutions at temperatures of 10 to 70'C for 10 hrs to 2 days, or by refluxing them in the same solution for 2 hrs or by ionexchanging the H-type zeolites which was prepared by the exchange with ammonium ion followed by calcination. The Catalysts were calcined at the temperature over 400'C for 2 hrs before use. Ammonia, methylamines and methanol of industrial grades are used without further purification. Reaction tests were conducted using a fixed bed flow reactor consisting of four half-inch diameter and 300mm long stainless tubes linked together. A salt bath was used to heat the reactor. A semi-adiabatic pilot-scale reactor was also used for some catalysts. The feed and product mixtures were occasionally collected for gas chromatograph analysis.

Y. Ashina et al.

781

RESULTS AND DISCUSSIONS 1. Performance of catalyst as received form under the conventional reaction conditions. The catalysts were tested under the conventional reaction conditions. Conversion was calculated on methanol basis. Selectivity of respective methylamines is defined as; Carbon equivalent of the respective methylamines formed I Carbon equivalent of the total methylamines formed. The rate constants are calculated based on pseudo first order with methanol, though it is known the reaction rate is considerably affected by the existence of water and monomethylamine particularly at lower temperatures.

50

~~l.S:

__

"='i"

I:r-

--ti

_

TMA

2 MeOH Fig.2 Activity and Selectivity of the catalysts as received form. ------: equilibrium, (400'C, 19atm, Feed; MeOH/NH3=1/1(wt), .:silica alumina, o:natural mordenite, A:200H, x:NaY) 200H(synthetic H-mordenite) exhibited very high activity for methanol turnover while natural mordenite was slightly less active than the conventional catalyst. Di-selectivity of mordenites exceeded the equilibrium level obviously showing the shape selective effect of zeolite catalyst. In a moderate conversion range, these catalysts were more di-selective at lower conversions than at higher ones. The natural mordenite was higher than 200H in di-selectivity while Y type was more like a conventional catalyst with no shape selective effect. A long-run test, however, showed that the mordenite was deactivated very rapidly at this temperature by the significant coke formation. The activity actually dropped by half only in 150 hrs while di-and tri-selectivity declined and mono-selectivity rose sharply indicating the pore entrance being narrowed by coke deposition. 2. Improvement of activity and selectivity of mordenite catalyst. The effect of metal cation in mordenite was studied to improve both activity and selectivity particularly to enhance activity in the lower temperature range where the catalyst life would be prolonged to a commercially satisfactory level. It was found that the activity and selectivity were significantly affected by the quantity of alkali metal cations. The higher alkali metal content resulted in the higher mono-, di-selectivity, lower tri-selectivity and the lower activity. It was shown in Fig.3 that mordenites with the alkali metal content cntrolled to a middle range, i.e., 0.04 to 0.05 mol/100g, exhibited high activity while maintaining high di-selectivity.

782 (CA-S-4)

r----------.,xlO- t 60 4'U

~

C

3':': __________

~ conventional level (380"G)

2

4

•

2

•

u >-

~u

:~

QJ

3120

~

CIl

70

6 x10- 2

alkali metal content (molll00g)

Fig.3 Effect of Alkali Metal Cation on activity and selectivity of mordenite (325'C, 18 atm, Feed: NH3/MeOH= 1/1 (wt )

1.Oc-----------.

80

90

100

MeOH conversion ("to)

Fig.4 Effect of Steaming (reacted at 320'C, acid-treated natural mordenite) .: Nontreated 0: Steam-treated at 400'C, 15 atm, 20hr

.....u CIl

2

III ......

....,

z:

:~

1: 40 3l

«

::E

Cl

1.5

1.6

1.7 lIT (11K)

1.8xlO-3

Fig.5 Arrhenius Plot .: Nontreated O:Steam-treated at 400'C, 15 atm, 20hr

20 300

350

temp. ('t)

400

Fig.6 Temperature dependence of selectivity (acid-treated natural mordenite, conversion 95%)

It was found that treating mordenite with high temperature steam further increased di-selectivity and suppressed tri-formation at a minor expense of activity. As shown in Fig.4, acid-treated mordenite which had been brought contact with superheated steam at 400'C, 15atm for 20hrs exhibited the dimethylamine selectivity several points higher than non steam-treated mordenite. The zeoli tic structure was not affected by steaming at this temperature as far as X-ray diffraction pattern was concerned. Presumably active sites on the external surface which accounts for approximately 10% of total surface area and is responsible for equilibrium-controlled reaction was damaged by steaming. Acid property and effective pore dimension could also be affected by steaming. Contrary to the thermodynamic equilibrium, tri-formation was lower and diselectivity higher at a lower temperature as a result of the diffusion of trimethylamine being further disturbed at the pore entrance at a lower temperature.

Y. Ashina et al.

783

The conversion for the optimum selectivity favorably shifted to a higher side by reducing reaction temperature. Temperature dependence of activity and selectivity of this catalyst are shown in Fig.5 and Fig.6. Activation energy of the reaction (Er) was calculated as 20.5 kcal/mol. 3. Effectiveness factor To examine the effect of intraparticle diffusion in the catalyst in this reaction, the intraparticle model as shown in Fig.7 was assumed based on various physical measurements such as pore size distribution by Porosimeter (:>35A) and nitrogen gas adsorption (:>10A), crystal size, surface area and pore volume.

DIAMETER particle pore ~~--

300-1000 A

crystallite

)+,-1- secondary

particle

catalyst granule

0.3 )J

200A

4 mm

2000 A

Fig.7 Intraparticle model of mordenite catalyst Effective diffusion coefficients of methanol/ammonia at 320'C were determined based on gas diffusion equation for large and medium pores assuming tortuosity factor as 3 and crystaline density as 1.7 g/cm3.(12) It was assummed the coefficient in the crystallite was equivalent to that of propane in Na-A crystal.(13) These values are listed in Table 1 as well as the relative diffusion resistance for respective pores. Table 1

Intraparticle diffusion of mordenite

Pore Pore Void Particle radius vol. density (g/m!) r (A) (mllg)

Particle radius R (m)

De (m2/sec)

R2/Defor respective particle

R2/De for catalyst granule

large 1000

0.33

0.333

1.01

2 X 10-3

3.2 X 10- 7

13

13

med.

100

0.07

0.106

1.52

1.5 X 10-7

6.2 X 10-8

3.6 X 10-7

65

small

(3)

0.21

0.357

(1.7)

1.5 X 10-8

(8 X 10- 12)

2.8 X 10-5

5 X 105

In spite of the fact that the diffusion in a small pore is extremely low, the overall diffusion rate is not determined by the intracrystaline diffusion because the crystal size is very small. Effectiveness factor of this catalyst was calculated as approximately 90% at 320'C using the effective diffusion coefficient of De = 3.2 X 10-7 + 6.2 X 10- 8 = 3.8 X 10-7 m2/sec, indicating the influence of diffusion would begin to appear around this temperature. 4. Commercialization As shown in the above, an optimized mordenite catalyst still maintained a commercially feasible activity at a temperature as low as 300·C. Di-selectivity around this temperature was 64%, 2.3 times as high as the conventional level.

784 (CA-5-4) The catalyst life of this mordenite was checked by a pilot-scale reactor and the result shown in Table 2 was obtained. Table 2

Long-run test result (acid-treated natural mordentte)

On-strea m React. Temp.(·C) tim e: t(hr)

ko(fresh)1

k(used)

Degeneration * 1 Const: b(1/hr)

310

1400

1.128

8.58

X

10-5

330

522

1.244

4.19

X

10-4

*1) b = lit In (ko/k) *2) b = bo exp (-Ed/RT)

Activation Energy*2 of degeneration: Ed(kcall mol) 55.3

Feed: NH3' MMA, TMA, MeOH 53%, 11 %, 16%I 20 %

Based on Ed = 55.3 kcal/mol for exponential degeneration and Er = 20.5 kcal/mol, the catalyst life was calculated as 375 days at the temperature shifting from 300'C to 328'C to maintain the conversion of 98.0%. Though the degeneration of mordenite catalyst is very rapid at a higher temperature, one-year continuous operation could thus become possible at a lower temperature as Ed is much larger than Er. Various elements must be considered for the commercial use of this catalyst. The reaction temperature significantly affects catalyst life thereby determining the necessity of catalyst regeneration step. Selection of conversion is also important. It may be necessary to operate at the conversion which gives the highest di-selectivity, or high conversion operation at some sacrifice of selectivity may be more attractive for it eliminates the need for methanol recovery. The drastic change in quantity and composition of process stream makes the reoptimization of the recovery system essential. For example, the reduction in the amount of ammonia in a process flow may create an opportunity for another heat recovery based on the raised dew point of the stream. Problems may also arise in such a low flow rate operation, particularly in the adiabatic reaction system. All these factors must be taken into consideration in optimization of the process to realize the highest return from the employment of this catalyst. Another important problem would be encountered when the tri-formation still exceeds the market demand and the surplus must be recycled for reuse. Trimethylamine is inactive over this mordenite catalyst because of its reactant shape selective effect and would be accumulated in the recycle system. This is experimentally shown as below.

~lr~~react~3?

feed

5}o

product

Fig.8 One-step reaction Process

Fig.9 Two-step reaction Process

Y. Ashina et al.

Table 3

785

Mass Balance for Recycle Process (one-step reaction) Relative now rate (wt) silica alu mina (400' C)

acid treated natural mordenite (350'C)

Cat. No. Compo.

(1)

(2)

(3)

(4)

NH3

5.2

55.9

50.7

50.7

11.7

12.4

11.7

MMA DMA

11.9

TMA MeOH

(5)

20.0

17.8

19.0

20.0

0.3

H2 O

(1)

(2)

(3)

(4)

5.2

52.7

47.5

47.5

11.0

11.7

11.0

0.7 11.9

17.8

20.1

16.2

17.4

20.1

0.4

11.1

Total

105.4

105.4

0.7 11.9

11.9

1.2

(5)

16.2

1.2

11.1 100.0

80.2

100.0

74.7

Under the conditions as in Table 3, the product restriction effect and the reactant restriction effect offset each other giving no substantial merit. It was found this problem could be overcome by the use of a small reactor with an equilibrium type catalyst in combination with a main reactor charged with a mordenite catalyst. Table 4 shows that the shape selective ability of zeolite catalyst can be fully displayed in this two-step process as a result of the surplus trimethylamine being converted according to equilibrium. Table 4

Mass Balance for Recycle Process (two-step reaction) Relative now rate (wt)

Cat. now No.

A: acid-treated natural mordenite (350'C)

B: silica alumina (400'C)

(2)

(3)

(4)

(5)

43.3

37.7

37.7

38.0

MMA

8.2

11.6

11.2

8.2

0.4

DMA

6.3

13.5

6.3

13.5

TMA

8.0

11.7

11.6

8.0

0.1

20.1

0.4

60.5

60.5

Compo. NH3

MeOH

(1)

5.3

20.1

H2 O Total

11.0 85.9

85.9

(6)

786 (CA-S-4) As described in the foregoing, the merit of the use of this catalyst either in terms of form or of magnitude may vary greatly depending on the choice of process conditions, business environment and other factors peculiar to respective producers. Roughly speaking, by employing this catalyst, 40% cut in utility cost is theoretically possible for average methylamine makers with reasonably optimized process provided the market balance of methylamines is the same as in Japan. However, since the maximum reduction in utility cost may not be justified because of the high investiment necessary for the plant modification, careful cost/performance optimization is required. The use of this catalyst also makes possible approximately 40% capacity increase for an existing plant or 30% reduction in construction cost for a new plant. BIBLIOGRAPHY 1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13.

J. ISSOIRE, et co Van Long, "Etude de la thermodynamique chimique de la reaction de formation des methylamines". J. RAMIOULLE, et al., Hydrocarbon Processing, July (1981), 113 Hydrocarbon Processing, Nov. (1973) Hydrocarbon Processing, Nov. (1981) US. 3384667 US. 4082805 BE. 873851 BE. 875674 DT. 3010791 US. 4313003 US. 4254061 A. EDWARD et al., Ind.Eng.Chem.Fundam., 11, No.4 (1972), 540 C. SATTERFIELD et al., AIChE J., July (1967), 731