Development of the Methanol-to-Gasoline Process - American

2 Development of the Methanol-to-Gasoline Process 1

J. E. PENICK, W. LEE, and J. MAZIUK

Downloaded by UNIV OF ARIZONA on December 2, 2012 | http://pubs.acs.org Publication Date: July 28, 1983 | doi: 10.1021/bk-1983-0226.ch002

Mobil Research and Development Corporation, New York, N Y

10017

Mobil has developed a new process for conversion of methanol to high quality gasoline using a unique class of shape-selective zeolite catalysts. This process provides a novel route to gasoline from either coal or natural gas via intermediate methanol. Two reactor systems have been devised by Mobil for this process—fixed bed and fluid bed. The f i r s t commercial application of the fixed-bed process w i l l be i n a 14,000 B/D (gasoline) plant based on natural gas which w i l l be located in New Zealand. The fluid-bed process i s being scaled up to a 100 B/D (methanol) demonstration plant located in W. Germany. In this paper, the chemistry of the process is reviewed and the engineering involved i n selection, scale-up, and design of an appropriate reactor system to produce an acceptable product i s discussed. M o b i l has developed a novel process f o r converting methanol i n t o high q u a l i t y g a s o l i n e . ( 1 ) Since methanol can be made from n a t u r a l gas o r c o a l by w e l l - e s t a b l i s h e d commercial technology, the Methanol-to-Gasoline (MTG) process provides a unique approach to the p r o d u c t i o n of g a s o l i n e from e i t h e r raw m a t e r i a l as shown i n Figure 1. In the MTG process, methanol i s q u a n t i t a t i v e l y converted to hydrocarbons and water. The hydrocarbons are p r i m a r i l y g a s o l i n e b o i l i n g range m a t e r i a l s s u i t a b l e f o r use as high q u a l i t y automotive f u e l . 1

Correspondence should be sent to W. Lee, Mobil Research and Development Corporation, Paulsboro, NJ 08066

0097-6156/83/0226-0019$08.50/0 © 1983 American Chemical Society

In Chemical Reaction Engineering—Plenary Lectures; Wei, J., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1983.


Ash

τ

Coal Gasifier

Stearn

MTG Process

GASOLINE Τ Water

Making gasoline from c o a l o r n a t u r a l gas v i a MTG.

Raw Methanol

Water

Τ

MTG r-*-GASOLINE Process

F i g u r e 1.

Methanol Process

Methanol Process

Raw Methanol

Synthesis Gas (CO + hh)

(CO + H2)

Synthesis Gas

Steam Reforming Process

NATURAL G A S

Steam

Oxygen

COAL



2.

PENICK E T A L .

Methanol-to-Gasoline Process

21

The MTG process w i l l soon see i t s f i r s t commercial a p p l i c a t i o n i n a p l a n t c u r r e n t l y being constructed f o r the New Zealand S y n t h e t i c Fuels Corporation L i m i t e d . This p l a n t , scheduled f o r completion i n 1985, w i l l produce g a s o l i n e from methanol derived from New Zealand n a t u r a l gas. I t w i l l produce some 570,000 tonnes per year (14,000 BPSD) of unleaded g a s o l i n e averaging 92 t o 94 research octane. This i s e q u i v a l e n t t o about 1/3 of New Zealand's g a s o l i n e consumption. Two 2,200 tonnes-per-day methanol p l a n t s w i l l provide the feed f o r the s i n g l e - t r a i n MTG p l a n t . In t h i s l e c t u r e , the development of the MTG process w i l l be reviewed. F i r s t , the unique aspects o f MTG — the c a t a l y s t , chemistry, and i t s s p e c i a l r e a c t o r design aspects — w i l l be d i s c u s s e d . Next, the choices f o r the conversion system w i l l be presented along w i t h the dual-pronged s t r a t e g y f o r development of both the f i x e d - and f l u i d - b e d processes. F i n a l l y , our f u t u r e development plans f o r t h i s general area o f technology w i l l be h i g h l i g h t e d . Catalyst The ZSM-5 c a t a l y s t , the key element i n the MTG process, i s a s h a p e - s e l e c t i v e s y n t h e t i c z e o l i t e w i t h a unique channel s t r u c t u r e as shown i n Figure 2.(2) This s t r u c t u r e i s d i s t i n c t i v e l y d i f f e r e n t from the f a m i l i a r wide-pore z e o l i t e , f a u j a s i t e , and the narrow-pore z e o l i t e s . The l i n e s i n Figure 2 represent oxygen atoms i n a s i l i c e o u s framework. I t contains a novel c o n f i g u r a t i o n of l i n k e d t e t r a h e d r a c o n s i s t i n g of e i g h t five-membered r i n g s . These u n i t s j o i n through edges t o form chains and sheets, l e a d i n g to the three dimensional s t r u c t u r e r e p r e s e n t a t i o n shown i n the f i g u r e . I t contains two i n t e r s e c t i n g channels: e l l i p t i c a l 10-membered r i n g s t r a i g h t channels and s i n u s o i d a l , tortuous channels. The pore-opening of these channels i s about 6 A and i s j u s t wide enough t o produce hydrocarbons b o i l i n g p r i m a r i l y i n the g a s o l i n e range. Since the mid-1970 s, more than 25 commercial p l a n t s have been commissioned using t h i s c a t a l y s t f a m i l y . There are c u r r e n t l y f i v e l i c e n s e d processes (other than MTG) based on the ZSM-5 f a m i l y : 1

D i s t i l l a t e dewaxing Xylene i s o m e r i z a t i o n Toluene d i s p r o p o r t i o n a t i o n Ethylbenzene

synthesis

L u b r i c a t i n g o i l dewaxing



22

C H E M I C A L REACTION

ENGINEERING

Figure 2. ZSM-5 s t r u c t u r e . (Reproduced "with p e r m i s s i o n from Ref. 2. Copyright 1978, Macmillan.)


2.

PENICK E T A L .

23


A l l o f these a p p l i c a t i o n s use fixed-bed r e a c t o r s , and, very i m p o r t a n t l y , a l l were scaled up from bench-scale p i l o t p l a n t data. This s u c c e s s f u l scale-up experience w i t h the ZSM-5 c a t a l y s t was an important c o n s i d e r a t i o n i n f o r m u l a t i n g the MTG development s t r a t e g y . Chemistry The conversion of methanol t o hydrocarbons over ZSM-5 i s well-represented by the f o l l o w i n g s i m p l i f i e d r e a c t i o n network(3): -H 0 ^ ^ +H 0

-H 0


2

2CH OH 3

2

2

CH3OCH3

> C -C Olefins 2

5

Paraffins > Aromatics Cycloparaffins

The i n i t i a l dehydration r e a c t i o n i s s u f f i c i e n t l y f a s t t o form an e q u i l i b r i u m mixture of methanol, dimethyl e t h e r , and water. These oxygenates dehydrate f u r t h e r t o g i v e l i g h t o l e f i n s . They i n t u r n polymerize and c y c l i z e t o form a v a r i e t y of p a r a f f i n s , aromatics, and c y c l o p a r a f f i n s . The above r e a c t i o n path i s i l l u s t r a t e d f u r t h e r by Figure 3 i n terms o f product s e l e c t i v i t y measured i n an isothermal l a b o r a t o r y r e a c t o r over a wide range of space v e l o c i t i e s . ( 3 ) The r a t e l i m i t i n g step i s the conversion of oxygenates t o o l e f i n s , a r e a c t i o n step that appears t o be a u t o c a t a l y t i c . I n the absence of o l e f i n s , t h i s rate i s slow; but i t i s a c c e l e r a t e d as the c o n c e n t r a t i o n of o l e f i n s i n c r e a s e s . Under MTG c o n d i t i o n s , almost no hydrocarbons are found higher than C ^ Q due to the shape s e l e c t i v i t y o f ZSM-5. While higher molecular weight compounds may be formed w i t h i n the c a t a l y s t channels, they do not escape due t o the s p a t i a l hindrance of the c a t a l y s t s t r u c t u r e . Reactor Design Although the design of a r e a c t o r system f o r the MTG process i n v o l v e s c l a s s i c a l chemical engineering p r i n c i p l e s , the unique c a t a l y s t and r e a c t i o n mechanisms impose important design c o n s t r a i n t s . These i n c l u d e the h i g h l y exothermic nature of the r e a c t i o n , the need f o r e s s e n t i a l l y complete methanol conversion, steam d e a c t i v a t i o n of the c a t a l y s t , the "band-aging" phenomena, and durene formation. •

Exotherm - The heat of r e a c t i o n i s 1740 kJ/kg of methanol which i f u n c o n t r o l l e d would g i v e an a d i a b a t i c temperature r i s e of about 600°C. Therefore, a p r i n c i p a l c o n s i d e r a t i o n i n designing a r e a c t o r system i s r e a c t i o n heat management.



Wt. %

Products,

( ^ )

i l ι 111

ι

1

1—I

Olefins)

5

ι ι 1 1 ifr—Q-o-i

Olefins 1

Aromatics

P—D-D-

Paraffins (and C $ +

10

M i l l

o-o o- W a t e r

ι ι ι 111

C2-C5

1—τ—ι

Figure 3. E f f e c t o f space v e l o c i t y a t 101 kPa and 371 C. (Reproduced w i t h permission from Ref. 3. Copyright 1977 Academic Press.)

S p a c e Time

1


2

w w S ο

3

δ w ο

H

ο

m >

ο w

to -Ρ»


2.

PENiCK E T A L .


25

•

Complete Methanol Conversion - The major products of the MTG conversion are hydrocarbons and water. Consequently, any unconverted methanol w i l l d i s s o l v e i n t o the water phase and be l o s t unless a d i s t i l l a t i o n step to process the very d i l u t e water phase i s added to the process. Thus, e s s e n t i a l l y complete conversion of methanol i s h i g h l y p r e f e r r e d .

•

C a t a l y s t D e a c t i v a t i o n - Under MTG r e a c t i o n c o n d i t i o n s , ZSM-5 c a t a l y s t undergoes two types of aging which c o n t r i b u t e to a gradual l o s s of c a t a l y s t a c t i v i t y . A r e v e r s i b l e l o s s r e s u l t s from "coke" formed on the c a t a l y s t as a r e a c t i o n byproduct. This d e a c t i v a t i o n i s a t y p i c a l c a t a l y t i c process design problem. What i s unusual about the MTG process i s that a r e a c t i o n product, steam, i s a l s o r e s p o n s i b l e f o r a gradual l o s s o f a c t i v i t y . However, low r e a c t o r temperatures and low p a r t i a l pressure of water w i l l minimize t h i s aging and favor a long c a t a l y s t l i f e .

•

Band-Aging - E s p e c i a l l y w i t h f r e s h c a t a l y s t s , the r e a c t i o n occurs over a r e l a t i v e l y s m a l l zone i n a f i x e d bed. This r e a c t i o n f r o n t marches down the c a t a l y s t bed as the coke d e p o s i t s f i r s t d e a c t i v a t e the f r o n t part of the bed (Figure 4 ) . Use o f a s u f f i c i e n t c a t a l y s t volume permits a fixed-bed design i n which on-stream periods are long enough to avoid o v e r l y frequent regeneration c y c l e s .

•

Durene Formation - An aromatic compound, durene ( 1 , 2, 4, 5-tetramethylbenzene) i s produced i n the MTG process. Durene has e x c e l l e n t research octane blending q u a l i t y (110 RON c l e a r ) and b o i l s w i t h i n the g a s o l i n e d i s t i l l a t i o n range (197°C); however, i t s f r e e z i n g p o i n t i s r e l a t i v e l y high a t 79°C. Based on our current designs, the g a s o l i n e d i r e c t l y from the ZSM-5 r e a c t o r s could c o n t a i n 4 t o 7 wt% durene. These concentrations of durene could lead to problems w i t h s o l i d s build-up i n carburetors during c o l d s t a r t s . While durene i s present i n conventional g a s o l i n e , the t y p i c a l concentrations are low enough not t o cause problems.



26

CHEMICAL

II

0

n

«frft»^frofr;ii^

0.1

0.2

0.3

ENGINEERING

ι

ι

ι

ι

ι

ι

1

0.4

0.5

0.6

0.7

0.8

0.9

1

Fractional Bed

Figure k. bed.

REACTION

Length

Movement o f normalized temperature p r o f i l e through


2.

PENICK E T A L .


27


In the MTG process, durene i s mostly formed by a l k y l a t i o n of lower molecular weight aromatics w i t h methanol and/or ether. Low methanol p a r t i a l pressures and high r e a c t i o n temperatures tend t o reduce the durene l e v e l , presumably by reducing the c o n c e n t r a t i o n overlap of methanol/ether w i t h the aromatics formed. This overlap tends t o i n c r e a s e w i t h l a r g e r c a t a l y s t p a r t i c l e s . As a consequence the e a r l y fixed-bed development work was conducted using 1 mm diameter extrudates. A l l of the above f a c t o r s must be considered i n the v a r i o u s r e a c t o r systems f o r the MTG process: fixed-bed r e a c t o r s , f l u i d - b e d r e a c t o r s and t u b u l a r r e a c t o r s . Figure 5a shows three v a r i e t i e s of fixed-bed r e a c t o r s and Figure 5b i l l u s t r a t e s heat-exchanger and f l u i d - b e d r e a c t o r v a r i a n t s . Fixed-Bed

Systems

The major advantage of these fixed-bed systems i s that the fixed-bed technology can be simple, and w i l l r e q u i r e minimum scale-up s t u d i e s . As i n d i c a t e d e a r l i e r , t h i s type o f design has been s u c c e s s f u l l y scaled-up t o commercial s i z e d i r e c t l y from bench-scale p i l o t p l a n t data many times. One common method f o r managing the a d i a b a t i c temperature r i s e i n fixed-bed r e a c t o r s i s t o d i l u t e the reactant stream w i t h a gas that would provide the mass to absorb the heat of r e a c t i o n . L i g h t hydrocarbon r e a c t i o n products (methane, ethane, propane) can be r e c y c l e d f o r t h i s purpose. However, the c o s t s a s s o c i a t e d w i t h the r e c y c l e o p e r a t i o n could be s u b s t a n t i a l i f the r e a c t i o n i s s t r o n g l y exothermic. S p e c i a l a t t e n t i o n must be given t o reduce the r e c y c l e r a t i o . Use of m u l t i p l e beds/reactors i n s e r i e s w i t h i n t e r c o o l i n g or quenching i s a method which can be used t o reduce the amount of r e c y c l e and i t s a s s o c i a t e d c o s t s . M u l t i p l e c a t a l y s t beds reduce c o s t s by using the r e c y c l e m a t e r i a l s e v e r a l times before i t i s separated from the r e a c t i o n products. In Figure 6a, a multibed r e a c t i o n system i s shown f o r which only p a r t i a l conversion i s taken across each bed; the e f f l u e n t i s cooled before e n t e r i n g each succeeding bed such that the e f f l u e n t from each bed does not exceed a s p e c i f i e d maximum temperature. The band-aging phenomena p r e v i o u s l y discussed would complicate the development o f such an MTG process s i n c e the c a t a l y s t appears t o age from the f r o n t t o the back. Thus, i t might be d i f f i c u l t , e s p e c i a l l y f o r the f i r s t bed, t o m a i n t a i n p a r t i a l conversion a t a given l e v e l f o r a reasonable amount o f time.


28

C H E M I C A L REACTION ENGINEERING

Standard


Feed

_

Quench

Radial Flow

Recycle

Feed

^ ^Recycle

Feed

_

_

Recycle

Figure 5a. A d i a b a t i c r e a c t o r s .

Tubular Isothermal

Feed

Heat Exchange Medium

m

Fluid-Bed Systems

Recycle

Φ



Feed

Figure 5b. Nonadiabatic r e a c t o r s .


PENICK E T A L .


MeOH Feed ^ ^Recycle

i


X r Steam H 0' 2

χ

Gas Recycle %

-C

Gas Make

• Raw Gasoline

H0 2

Figure 6a. M u l t i b e d designs.

MeOH Feed

I n t e r s t a g e heat removal.

Recycle^

Gasoline

Figure 6b. M u l t i b e d design.

I n t e r s t a g e feed quenching.




30

In Figure 6b, a multibed r e a c t i o n system i s shown i n which only part of the reactant i s fed to each bed. The reactant fed to each bed i s e s s e n t i a l l y completely converted and the r e a c t i o n products from that conversion become part of the temperature-moderating m a t e r i a l f o r each succeeding bed. In such a r e a c t i o n system the methanol feed to each succeeding bed would be mixed w i t h aromatics made i n the preceding beds - - an arrangement that tends to increase durene formation. S t i l l another m u l t i - r e a c t o r approach i s to d i v i d e the MTG r e a c t i o n i n t o two steps as shown i n Figure 7. In the f i r s t step, methanol i s p a r t i a l l y dehydrated to form an e q u i l i b r i u m mixture of methanol, dimethyl ether and water over a dehydration c a t a l y s t . About 15% of the r e a c t i o n heat i s released i n t h i s f i r s t step. In the second step, t h i s e q u i l i b r i u m mixture i s converted to hydrocarbons and water over ZSM-5 c a t a l y s t w i t h the concomitant release of about 85% of the r e a c t i o n heat. Though t h i s two step approach does not have any of the inherent complications of the p r e v i o u s l y mentioned multibed r e a c t i o n systems, i t leaves one w i t h a s u b s t a n t i a l amount of the r e a c t i o n heat (85%) s t i l l to be taken over one c a t a l y s t bed. This r e q u i r e s a f a i r l y high r e c y c l e stream to moderate the temperature r i s e over the second r e a c t o r . Such a high r e c y c l e design would r e q u i r e c a r e f u l engineering i n order to t r a n s f e r heat e f f i c i e n t l y from the r e a c t o r e f f l u e n t to the r e c y c l e gas and r e a c t o r feed. However, t h i s two stage r e a c t o r system i s the simplest of the fixed-bed systems to develop. Fluid-Bed Reactor Systems A r e a c t o r system which has s e v e r a l a t t r a c t i v e f e a t u r e s f o r heat removal i s the f l u i d bed. With a f l u i d bed, the r e a c t i o n heat can be removed by using heat exchange c o i l s i n the bed, or c a t a l y s t can be c i r c u l a t e d to an e x t e r n a l c o o l e r f o r heat recovery. Heat t r a n s f e r i n a f l u i d bed i s so r a p i d that the danger of l o c a l overheating i s minimal. Furthermore, the heat of r e a c t i o n can be recovered e f f i c i e n t l y as high pressure steam. As expected, the f l u i d - b e d y i e l d p a t t e r n i s d i f f e r e n t from that of the f i x e d bed; C 5 + g a s o l i n e y i e l d i s lower, but the l i g h t o l e f i n s y i e l d s are much higher, p a r t l y because l i g h t product gas i s not r e c y c l e d . A l k y l a t i o n of these o l e f i n s w i t h product isobutane and blending gives an increased g a s o l i n e y i e l d compared to that of the f i x e d bed. Two f a c t o r s complicate the development of a f l u i d - b e d MTG process. One i s the need to ensure the complete conversion of methanol; the other i s to develop a f l u i d c a t a l y s t s u f f i c i e n t l y rugged to withstand the a b r a s i v e f o r c e s inherent i n f l u i d - b e d operation.



PENICK E T A L .


Figure T.

Two-stage fixed-bed MTG.


C H E M I C A L R E A C T I O N ENGINEERING

32

Since i n a l a r g e diameter f l u i d bed, bubbles may bypass the c a t a l y s t , elaborate hydrodynamic s t u d i e s are required to scale-up the r e a c t o r i n t e r n a l s to accomplish the complete conversion of methanol. As mentioned e a r l i e r , e s s e n t i a l l y complete conversion of methanol i s a process requirement; otherwise, a c o s t l y d i s t i l l a t i o n step must be added to recover unconverted methanol. Thus, the f l u i d - b e d MTG process was assessed as being a d e s i r a b l e but a long-term development. The scale-up would have to proceed through s e v e r a l stages before i t would be ready f o r commercialization.


Tubular, Heat-Exchange

Reactor Systems

Another r e a c t o r system which has s e v e r a l a t t r a c t i v e f e a t u r e s f o r heat removal i s the t u b u l a r , heat-exchange r e a c t o r . Good temperature c o n t r o l can be achieved i n the t u b u l a r r e a c t o r i f the coolant approximates an i s o t h e r m a l heat s i n k . L i g h t gas r e c y c l e can be reduced s i g n i f i c a n t l y compared to fixed-bed systems. Tubular reactors have been used f o r Fischer-Tropsch r e a c t i o n s and f o r s y n t h e s i s of methanol and p h t h a l i c anhydride, f o r example. The MTG r e a c t i o n temperature of approximately 400°C e s s e n t i a l l y precludes the use of b o i l i n g water as a coolant due to the h i g h steam pressures i n v o l v e d . While o i l c o o l i n g might conceivably be used, the temperatures are above the range u s u a l l y considered p r a c t i c a l f o r such a design and l a r g e f l o w r a t e s would be required to achieve uniform c o o l i n g . Heat t r a n s f e r s a l t and b o i l i n g mercury are other p o s s i b i l i t i e s ; each presents p a r t i c u l a r design c h a l l e n g e s . Of these, heat t r a n s f e r s a l t probably represents the best compromise f o r an MTG a p p l i c a t i o n p r i m a r i l y because the s a l t could a l s o be used to c o n t r o l the p e r i o d i c c a t a l y s t regeneration o p e r a t i o n which e n t a i l s temperatures of about 500°C. From our viewpoint the heat exchange r e a c t o r has the f o l l o w i n g disadvantages: (1) m o l t e n - s a l t c o o l i n g presents s i g n i f i c a n t o p e r a t i o n a l problems, (2) t u b u l a r r e a c t o r s g e n e r a l l y c o n t a i n l e s s c a t a l y s t per u n i t of v e s s e l volume than other r e a c t o r types, (3) the maximum p r a c t i c a l c a p a c i t y f o r a s i n g l e r e a c t o r i s l e s s than f o r other types and (4) very c a r e f u l mechanical design w i l l be required f o r a t u b u l a r r e a c t o r which can r e l i a b l y withstand both the conversion and c a t a l y s t regeneration c y c l e s . Strategy f o r Commercialization Since there are many p o s s i b l e r e a c t i o n systems f o r an MTG process, i t was necessary to focus the development e f f o r t s . In the mid-1970 s, M o b i l decided to pursue a double-pronged approach i n order to ensure both s h o r t - and long-term o b j e c t i v e s . The plan was: 1


2.

PENICK E T A L .


33


NEAR-TERM — P r o c e e d w i t h commercialization of a simple, d i r e c t , and uncomplicated fixed-bed technology. Such a development would p a r a l l e l extensive experience i n s u c c e s s f u l commercialization of many fixed-bed processes using s i m i l a r c a t a l y s t s and operating c o n d i t i o n s . The f i v e l i c e n s e d processes using ZSM-5 c a t a l y s t f a l l i n that category. The simplest fixed-bed MTG system was the one which employed dehydration and ZSM-5 r e a c t o r s . This system was s t u d i e d e x t e n s i v e l y i n bench-scale u n i t s . These s t u d i e s i n a 3 cm diameter by 30-50 cm l e n g t h r e a c t o r s were considered t o be s u f f i c i e n t f o r scale-up. However, i n 1978, the New Zealand government expressed strong i n t e r e s t i n applying the MTG technology f o r conversion of indigenous n a t u r a l gas to h i g h - q u a l i t y g a s o l i n e . I n support of the New Zealand e f f o r t , i t was decided t o demonstrate the fixed-bed process i n a 4 B/D p i l o t p l a n t . Such a demonstration t e s t could be f i t i n t o the o v e r a l l schedule f o r the New Zealand venture and would provide l a r g e q u a n t i t i e s of product g a s o l i n e f o r road t e s t i n g i n cars t y p i c a l l y s o l d i n New Zealand. LONG-TERM — P r o c e e d w i t h an o r d e r l y development of the f l u i d - b e d process, which may o f f e r an a t t r a c t i v e a l t e r n a t i v e to the fixed-bed process f o r the f u t u r e w i t h i t s p o s s i b l y higher g a s o l i n e y i e l d , more e f f i c i e n t heat recovery, and economy of s c a l e . C u r r e n t l y , a 100 B/D f l u i d - b e d p i l o t p l a n t i n Germany i s nearing s t a r t up. This study i s a j o i n t p r o j e c t w i t h URBK, Uhde, and M o b i l supported i n part by the U.S. DOE and the German government. S u c c e s s f u l completion of the 100 B/D program w i l l be a p r e r e q u i s i t e f o r proceeding w i t h c o m m e r c i a l i z a t i o n of the f l u i d - b e d process. Fixed-Bed Development and Commercialization Fixed-Bed Demonstration P l a n t . The major o b j e c t i v e of the demonstration t e s t was t o v e r i f y the bench u n i t r e s u l t s i n a l a r g e r p l a n t operating a t c o n d i t i o n s s i m i l a r to a commerc i a l - s i z e r e a c t o r . The l i n e a r v e l o c i t y of the reactant i s the only v a r i a b l e which i s s i g n i f i c a n t l y d i f f e r e n t between a bench-scale u n i t and a commercial-size r e a c t o r . The other operating c o n d i t i o n s are normally the same f o r both r e a c t o r s . By making the l e n g t h o f the c a t a l y s t bed i n the demonstration



34


p l a n t i d e n t i c a l to that of the proposed commercial-size r e a c t o r , the gas v e l o c i t i e s a l s o become i d e n t i c a l , and thus a l l the process v a r i a b l e s become i d e n t i c a l . For a simple gas r e c y c l e system, there i s a strong i n c e n t i v e to have the bed l e n g t h be as short as p o s s i b l e to hold down the pressure drop and thus compression c o s t s . But as c a t a l y s t beds get s h o r t , one has to be concerned w i t h m a l d i s t r i b u t i o n s i n the c a t a l y s t packing and consequently i n the flow through the bed. I f there were regions of high flow i n an MTG c a t a l y s t bed, unconverted methanol could breakthrough e a r l y , r e s u l t i n g i n e i t h e r short c y c l e s or i n methanol l o s s to the water phase. M o b i l came to the c o n c l u s i o n that w i t h a r e a c t o r bed l e n g t h of 2.5 meters, i t could e f f e c t i v e l y load the c a t a l y s t and ensure even d i s t r i b u t i o n of the gas i n r e a c t o r s having l a r g e diameters. The c a t a l y s t bed l e n g t h f o r the demonstration t e s t was set at the l i k e l y l e n g t h f o r a commercial p l a n t . For the demonstration u n i t , the i n t e r n a l diameter was s e l e c t e d to be about 10 cm, which i s s i g n i f i c a n t l y l a r g e r than c a t a l y s t p a r t i c l e s i z e to minimize w a l l e f f e c t s . In a d d i t i o n , heat t r a n s f e r along the r e a c t o r w a l l would a l s o be n e g l i g i b l e . The r e s u l t s of the demonstration plant are compared w i t h those of the bench u n i t i n Table I . The two u n i t s were operated at i d e n t i c a l c o n d i t i o n s except f o r gas v e l o c i t y . The product y i e l d s and s e l e c t i v i t i e s over the f i r s t c y c l e were remarkably s i m i l a r . The a d i a b a t i c temperature r i s e and the band aging behavior were a l s o the same. However, the demonstration u n i t performed considerably b e t t e r w i t h respect to c y c l e l e n g t h . T y p i c a l c y c l e s were about 50% longer. S i m i l a r behavior of longer c y c l e lengths have been observed i n commercial operation w i t h other ZSM-5 c a t a l y z e d processes. In the product q u a l i t y t e s t s , p a r t i c u l a r a t t e n t i o n was given to the e f f e c t of durene on automobile performance and g a s o l i n e handling since i t s content i n the MTG g a s o l i n e i s the only s i g n i f i c a n t d i f f e r e n c e from conventional g a s o l i n e s . A t e s t f l e e t designed to simulate the New Zealand automobile p o p u l a t i o n was used f o r the d r i v e a b i l i t y t e s t s . Figure 8 i l l u s t r a t e s the r e s u l t s of one t e s t s e r i e s conducted at -18°C ambient temperature. (4_) The data show the weight percent durene that could be t o l e r a t e d before encountering any abnormal d r i v e a b i l i t y c h a r a c t e r i s t i c s . A l l cars performed normally on g a s o l i n e c o n t a i n i n g 3 wt% or l e s s durene. With 4 wt% durene, the most c r i t i c a l car encountered a s l i g h t l y rougher-thannormal engine i d l e c o n d i t i o n . Of the 12 cars t e s t e d at -18°C, eleven had no decrease i n d r i v e a b i l i t y at 4 wt% durene, 7 were good at 5 wt%, and 3 were good at 6 wt% durene. Much higher durene t o l e r a n c e s were observed at higher temperatures. To ensure that there would be no d r i v e a b i l i t y or handling problems w i t h MTG g a s o l i n e , a t a r g e t of 2 wt% durene i n g a s o l i n e was set as the company s p e c i f i c a t i o n , though the MTG


2.

PENICK E T A L .


Table I .

35

Scale-Up R e s u l t s of Fixed-Bed MTG Bench Unit

4 B/D Unit


Conditions Methanol/Water Chg. (W/W) Conversion Reactor I n l e t Temperature (°C) Conversion Reactor O u t l e t Temperature (°C) Methanol Space V e l o c i t y (WHSV) Separator Temperature (°C) Recycle R a t i o (mol/mol chg) Pressure (kPa)

83/17

83/17

358

360

404 1. 49 9 2,163

407 1.6 52 9.2 2,156

Average F i r s t Cycle Y i e l d (Wt%) Hydrocarbon Products Methane Ethane Ethylene Propane Propylene Isobutane n-Butane Butènes C 5 + Hydrocarbons

Gasoline (Wt%) Octane, R+0

1.33 0.82 0.02 8.54 0.15 8.45 4.06 0.71 75.92 100.00

1.25 0.86 0.03 8.60 0.15 8.39 4.20 0.74 75.78 100.00

80.2 95

80.2 95


CHEMICAL REACTION

ENGINEERING


36



2.

PENICK E T A L .


37

process was producing about 5 wt% at that time. To meet t h i s s p e c i f i c a t i o n w i t h conventional processing would have required undercutting the g a s o l i n e product. Consequently, an e x p l o r a t o r y program was undertaken t o see i f t h i s s p e c i f i c a t i o n could be met by a d d i t i o n of a simple g a s o l i n e t r e a t i n g step. I n a short time, a heavy g a s o l i n e t r e a t i n g (HGT) technology evolved. With the HGT, durene l e v e l s even higher than 5 wt% could be t r e a t e d down t o the 2% l e v e l . The HGT e f f i c i e n t l y f i t s i n t o the processing scheme. The hydrocarbon l i q u i d from the separator (shown i n Figure 9) i s normally s t a b i l i z e d by f r a c t i o n a t i n g a t around 140 p s i g ; t h i s pressure produces a r e l a t i v e l y hot g a s o l i n e bottoms stream. Sending t h i s high heat content stream d i r e c t l y t o a low-pressure g a s o l i n e s p l i t t e r a l l o w s us t o concentrate the durene i n a small bottoms stream. Only small amounts of energy need to be expended since the feed i s so hot and the s e p a r a t i o n can be made w i t h l i t t l e r e f l u x . The small amount of hydrogen required f o r the HGT can be r e a d i l y obtained from a p o r t i o n of the methanol s y n t h e s i s purge gas. This purge stream i s normally used as f u e l ; the f u e l value of the hydrogen can be replaced by the small amount of gas produced from the HGT step. The bottom p o r t i o n of the g a s o l i n e (approximately 15% i n the NZ design) i s processed i n a c a t a l y t i c r e a c t o r designed to convert the durene by both i s o m e r i z a t i o n and d e a l k y l a t i o n . The t r e a t e d heavy g a s o l i n e i s then s t a b i l i z e d and reblended w i t h the l i g h t g a s o l i n e . Neither the f i n i s h e d g a s o l i n e volume nor i t s octane numbers are a f f e c t e d . The minor weight l o s s from the t r e a t i n g step i s o f f s e t by a decrease i n the s p e c i f i c g r a v i t y of the product. I n c o r p o r a t i o n of the HGT process has g r e a t l y eased the durene c o n s t r a i n t o r i g i n a l l y imposed on the MTG r e a c t o r design. I n f a c t , as a consequence o f i n c o r p o r a t i n g HGT, we are able t o switch t o a l a r g e r c a t a l y s t p a r t i c l e s i z e , which makes more durene than the s m a l l e r , o r i g i n a l c a t a l y s t . Such a change brings about l e s s r e a c t o r pressure drop, and thus o v e r a l l , a more e f f i c i e n t p l a n t . Commercial Fixed-Bed Plant Design. The commercial fixed-bed MTG plant i s very s i m i l a r i n design concept t o the 4 B/D demonstration p l a n t . A t y p i c a l design of the r e a c t i o n s e c t i o n of a commercial p l a n t i s shown i n Figure 10. The feedstock may be d i s t i l l e d o r crude methanol. The feed i s vaporized by r e a c t o r e f f l u e n t heat-exchange and enters i n t o the dehydration r e a c t o r . The dehydration r e a c t o r e f f l u e n t i s mixed w i t h preheated r e c y c l e gas and enters the conversion r e a c t o r s . Although Figure 10 shows f o u r p a r a l l e l o r "swing" conversion r e a c t o r s , a l e s s e r o r g r e a t e r number of r e a c t o r s may be used depending upon the feed rate and regeneration frequency d e s i r e d .


38


Recycle Gas

Gas Make

Com pressor cT] • LPG Deethanizer


Gasoline Stabilizer

Steam Product Separator

^^Steam -Stabilized Gasoline

Raw Gasoline

Waste Water, Cooled Reactor Effluent

Figure 9.

Fixed-bed MTG separations s e c t i o n . DME Effluent Header

MeOH Superheater

DME Reactor r

f

—

H

MeOH Vaporizer Crude MeOH

Hot Excess Effluent From ZSM-5 Reactors

To Duplicate Reactor/Effluent Exchanger Trains ZSM-5 Reactors

MeOH JrlPreheater

Cooling Water To Separator

/ t [ \ Effluent/Recycle \[ y Exchanger Cooled Effluent From ZSM-5 Reactor

,1 1,1

I

Cooler Compressed Recycle Gas From Separator

J-UJ

Recycle Gas Header

Figure 10. Flowsheet o f the r e a c t i o n s e c t i o n o f the fixed-bed MTG process.



2.

PENICK E T A L .


39

We have found that to e f f i c i e n t l y recover the high heat content i n the r e a c t o r e f f l u e n t i t i s d e s i r a b l e t o s p l i t the stream i n t o two p a r t s . The m a j o r i t y of the e f f l u e n t from a p a r t i c u l a r r e a c t o r i s used t o heat the r e c y c l e gas t o that r e a c t o r . The excess hot e f f l u e n t s from a l l the r e a c t o r s are combined and used f o r generating steam and heating the methanol feed to the dehydration r e a c t o r . A f t e r these s e r v i c e s , the r e a c t o r e f f l u e n t s are combined and cooled by c o o l i n g water before e n t e r i n g the product separator. Considerable engineering judgment and e f f o r t are needed t o ensure that the heat recovery i s e f f i c i e n t , yet has low pressure drop. Since a l a r g e p o r t i o n of the heat contained i n the r e a c t o r e f f l u e n t has to be t r a n s f e r r e d back t o the c o l d r e c y c l e gas, t h i s exchanger arrangement received our s p e c i a l attention. We have a l s o found i t a d v i s a b l e to vaporize the methanol using a v a p o r - l i q u i d drum, analogous t o a steam drum. Such an arrangement increases the loop pressure drop because the excess hot e f f l u e n t flows through separate exchangers i n s e r i e s f o r preheating, v a p o r i z i n g , and superheating the methanol. However, t h i s arrangement has two advantages: •

A blowdown of l i q u i d methanol can be taken i f the crude methanol contains any s o l i d s .

•

As the c o n t r o l valves t o the various MTG r e a c t o r s are adjusted, pressure waves may be transmitted t o the methanol exchangers. I f one exchanger i s used f o r preheating, v a p o r i z i n g and superheating the methanol feed, the pressure waves may cause the area used f o r v a p o r i z a t i o n to f l u c t u a t e thereby l e a d i n g t o p o t e n t i a l i n s t a b i l i t y i n the v a p o r i z a t i o n r a t e .

Gas, l i q u i d hydrocarbon, and water phases are separated i n the product separator. Most of the gas i s r e c y c l e d . The net gas make and l i q u i d hydrocarbon streams are sent t o conventional petroleum f r a c t i o n a t o r s f o r separation and s t a b i l i z a t i o n as shown i n Figure 9. When methanol breaks through a c a t a l y s t bed, that r e a c t o r i s taken o f f stream and the c a t a l y s t i s regenerated by burning o f f the deposited coke. A i r i s used as the source o f oxygen t o burn o f f the coke and n i t r o g e n d i l u t i o n i s used t o c o n t r o l the peak burning temperature. Most of the r e a c t o r e f f l u e n t gases during regeneration are r e c y c l e d back to the r e a c t o r a f t e r water i s removed. The r e a c t i o n engineering i n the MTG plant centered on how deep the c a t a l y s t bed had t o be t o avoid premature methanol break-through. The d e c i s i o n that the minimum bed depth had t o be 2.5 meters e l i m i n a t e d r a d i a l r e a c t o r s from c o n s i d e r a t i o n .



40

Fluid-Bed Development


From the small bench-scale t e s t s , the f l u i d - b e d MTG development has proceeded to two phases of scale-up. The i n i t i a l scale-up was v e r t i c a l to a 4 B/D c a p a c i t y , and the current phase i s a h o r i z o n t a l scale-up to 100 B/D c a p a c i t y . V e r t i c a l Scale-Up of F l u i d Bed. The f l u i d - b e d process was f i r s t scaled up v e r t i c a l l y from the 45 cm t a l l bench-scale r e a c t o r to the 760 cm t a l l 4 B/D r e a c t o r shown s c h e m a t i c a l l y i n Figure 11. The 4 B/D r e a c t o r i n t e r n a l diameter was 10 cm. The operation of t h i s 4 B/D f l u i d p i l o t plant was very s u c c e s s f u l . E s s e n t i a l l y complete conversion of methanol was achieved at the design c o n d i t i o n s f o r s u p e r f i c i a l gas v e l o c i t i e s up to 0.5 m/sec. E x c e l l e n t s t a b i l i t y of the o p e r a t i o n was demonstrated over a 75-day run. Higher g a s o l i n e y i e l d and octane number than f o r fixed-bed processing were v e r i f i e d (Table I I ) . A t y p i c a l temperature p r o f i l e i n the r e a c t o r i s shown i n Figure 12. Despite the h i g h l y exothermic nature of the r e a c t i o n and the unusually high l e n g t h - t o diameter r a t i o (greater than 75) of the r e a c t o r , a very uniform temperature p r o f i l e was e s t a b l i s h e d . There was a complete absence of any troublesome "hot spots". In a d d i t i o n , the t r a n s i e n t temperature p r o f i l e s during heat-up and cool-down were a l s o uniform and s t a b l e . With these encouraging r e s u l t s , i t was decided to scale-up the f l u i d - b e d process f u r t h e r . A 100 B/D f l u i d - b e d program has been formulated to demonstrate h o r i z o n t a l scale-up. H o r i z o n t a l Scale-Up of F l u i d Bed. The 100 B/D p i l o t p l a n t w i l l provide a d d i t i o n a l i n f o r m a t i o n f o r design of a commercials c a l e , f l u i d - b e d process. Key data yet to be obtained a r e : •

r e a c t o r i n t e r n a l b a f f l e design to maintain complete conversion of methanol i n a l a r g e diameter f l u i d i z e d vessel.

•

c a t a l y s t a t t r i t i o n and u l t i m a t e l i f e .

A schematic diagram of the r e a c t o r system of the 100 B/D p l a n t i s shown i n Figure 13. There are three major v e s s e l s : r e a c t o r , regenerator, and e x t e r n a l c o o l e r . The r e a c t o r c o n s i s t s of a dense f l u i d - b e d s e c t i o n (60 cm ID χ 13.2 m height) l o c a t e d above a d i l u t e phase r i s e r . Two modes w i l l be studied to remove r e a c t i o n heat :



( y

OO "I

Catalyst Recirculation Line

Catalyst Storage Vessel

Air

Regenerator

Flue G a s

Water

H2O Tank

h B/D f l u i d - b e d p i l o t p l a n t .

Reactor10.2 c m χ 7.6 m

F i g u r e 11.

Vapor Feed

Liquid F e e d

)

Water

Tank Car

l eF ëeëc¥T Preparation

oo

Methanol

Disengager

Filter -

Gas

Liquid Hydrocarbons

HC Prod Tank

Separator

Off

Condenser


42


ENGINEERING

Table I I . Process Conditions and Product Y i e l d s from MTG Processes Fixed Bed

F l u i d Bed

83/17

83/17


Conditions Methanol/Water Chg. (W/W) Dehydration Reactor I n l e t Temperature (°C) Dehydration Reactor O u t l e t Temperature (°C) Conversion Reactor I n l e t Temperature (°C) Conversion Reactor O u t l e t Temperature (°C) Pressure (kPa) Recycle R a t i o (mol/mol chg.) Space V e l o c i t y (WHSV)

316 404 360

413

415

413

2,170 9.0 2.0

275 1.0

Y i e l d s (Wt. % of Methanol Charged) Methanol + Ether Hydrocarbons Water CO, C 0 Coke, Other 2

0.0 43.4 56.0 0.4 0.2 100.0

0.2 43.5 56.0 0.1 0.2 100.0

1.4 5.5 0.2 8.6 3.3 1.1 79.9 100.0

5.6 5.9 5.0 14.5 1.7 7.3 60.0 100.0

85.0 13.6 1.4 100.0

88.0 6.4 5.6 100.0

93

97

Hydrocarbon Product (Wt. %) L i g h t Gas Propane Propylene Isobutane n-Butane Butènes C 5 + Gasoline

Gasoline ( i n c l u d i n g A l k y l a t e ) [RVP-62 kPa ( 9 p s i ) ] LPG Fuel Gas

Gasoline Octane (R+O)


2.

PENICK E T A L .


43

6h


Reactor Elevation, Meters Above Feed Inlet

100

Figure 12.

Figure 13.

300 500 Temperature, °C

700

F l u i d - b e d r e a c t o r temperature p r o f i l e .

100 B/D f l u i d - b e d MTG demonstration p l a n t .


C H E M I C A L R E A C T I O N ENGINEERING


44

•

c i r c u l a t i o n of c a t a l y s t through an e x t e r n a l c a t a l y s t c o o l e r , and

•

i n t e r n a l heat exchange pipes immersed i n the dense bed reactor.

To m a i n t a i n a constant c a t a l y s t a c t i v i t y i n the r e a c t o r , a small f r a c t i o n of "coked" c a t a l y s t w i l l be c o n t i n u o u s l y regenerated and returned to the r e a c t o r . The mechanical design b a s i s of the 100 B/D u n i t has been v e r i f i e d on a f u l l - s c a l e Cold Flow Model (CFM). This non-reacting model proved very u s e f u l f o r o p t i m i z i n g b a f f l e design and c a t a l y s t c i r c u l a t i o n s t r a t e g i e s . S e v e r a l d i f f e r e n t b a f f l e designs — h o r i z o n t a l and v e r t i c a l arrangements (Figure 14) — have been t e s t e d i n the CFM using s e v e r a l experimental techniques: gas t r a c e r , capacitance probes, bed expansion a n a l y s i s , and v i s u a l o b s e r v a t i o n . Integrated residence time behavior of the bed and q u a n t i t a t i v e , l o c a l g a s / s o l i d s d i s t r i b u t i o n and bubble s i z e d i s t r i b u t i o n were measured. Experimental r e s u l t s i n d i c a t e that h o r i z o n t a l b a f f l e s are e f f e c t i v e i n breaking bubbles. In the t u r b u l e n t f l u i d i z a t i o n regime, h o r i z o n t a l b a f f l i n g provides a l s o the e f f e c t of s t a g i n g the f l u i d bed and l i m i t i n g the formation of gross c i r c u l a t i o n p a t t e r n s w i t h i n the bed. Thus, h o r i z o n t a l b a f f l e s were s e l e c t e d f o r i n s t a l l a t i o n i n the p i l o t p l a n t . The b a f f l e design was optimized to ensure s u f f i c i e n t c a t a l y s t f l u x through the b a f f l e d s e c t i o n . An o v e r l y r e s t r i c t i v e design can cause a marked d e n s i t y gradient i n the bed, w i t h most of the c a t a l y s t accumulating above the top b a f f l e s e c t i o n . The f i n a l b a f f l e design chosen f o r the 100 B/D p i l o t p l a n t d i d not e x h i b i t t h i s c a t a l y s t segregation phenomenon. We b e l i e v e that the 100 B/D p l a n t represents the f i n a l development step before a commercial-size r e a c t o r can be designed w i t h confidence. Future Development The fixed-bed MTG process designed f o r New Zealand represents a simple but h i g h l y r e l i a b l e design concept. The development of the HGT, which r e l a x e s the durene c o n s t r a i n t , opens the process v a r i a b l e window c o n s i d e r a b l y . For example, i f the r e c y c l e r a t i o i s reduced from 9:1 to 7:1, the s i z e of the r e c y c l e c i r c u i t i s reduced and i t becomes more e f f i c i e n t , l e a d i n g to reductions of about 30% i n heat t r a n s f e r area and compressor horsepower. However, the octane of the g a s o l i n e product i s reduced by about one number. S e v e r a l more advanced fixed-bed process arrangements which reduce processing c o s t s without reducing octane are being considered and subjected to p r e l i m i n a r y screening t e s t s .



PENiCK E T A L .

Figure Ik.


Examples o f f l u i d - b e d r e a c t o r i n t e r n a l b a f f l e s .



46

CHEMICAL

REACTION

ENGINEERING

The f l u i d - b e d process being developed i n the 100 B/D p i l o t plant has good p o t e n t i a l f o r becoming an important part of our f u t u r e technology. I t has inherent p o t e n t i a l advantages over the fixed-bed process as mentioned e a r l i e r . I t i s b a s i c a l l y f l e x i b l e f o r easy a d a p t a t i o n t o petrochemicals production and b e t t e r c o n t r o l of the a u t o c a t a l y t i c r e a c t i o n . We are l o o k i n g a t c a t a l y s t m o d i f i c a t i o n s both f o r improved y i e l d / o c t a n e c h a r a c t e r i s t i c s and f o r improved steam s t a b i l i t y . I t has been our experience that new, improved c a t a l y s t s are an i n t e g r a l part of r e a c t i o n engineering. To keep processes v i t a l and c o m p e t i t i v e , i t i s necessary t o c o n t i n u a l l y look f o r improved c a t a l y s t s . In a d d i t i o n t o the production o f g a s o l i n e , the prospects look promising f o r manufacturing d i e s e l f u e l s and chemicals using r e l a t e d technology. R e f e r r i n g back t o Figure 3, which shows the r e a c t i o n steps i n v o l v e d i n methanol conversion t o hydrocarbons, we observe that l i g h t o l e f i n s form f i r s t as intermediate products. As the r e a c t i o n proceeds, the l i g h t o l e f i n s react f u r t h e r t o form i s o p a r a f f i n s and aromatics. I n the g a s o l i n e production mode, we are operating a t the f a r r i g h t where the g a s o l i n e components dominate. By m o d i f i c a t i o n of c a t a l y s t and/or process c o n d i t i o n s , one can operate where the s e l e c t i v i t y t o l i g h t o l e f i n s i s h i g h . We have learned to i n c r e a s e the ethylene y i e l d to about 30 wt% of hydrocarbon, f o r i n s t a n c e . This process i s perhaps the most promising new route f o r the production of ethylene from s y n t h e s i s gas. In our e x p l o r a t o r y s t u d i e s , we have now found good leads f o r y i e l d i n g as much as 75-80% C 2 - C 4 o l e f i n product, which i n t u r n can be converted to good q u a l i t y d i s t i l l a t e . ( _ 5 > 6 ) We can now consider the co-production of g a s o l i n e and d i s t i l l a t e from methanol as shown i n Figure 15. Engineering and development s t u d i e s are i n the i n i t i a l phases of e v a l u a t i n g such concepts. The z e o l i t e - c a t a l y z e d methanol conversion technology, whether the d e s i r e d product i s g a s o l i n e , d i e s e l , j e t f u e l o r ethylene f o r petrochemicals, w i l l provide new o p p o r t u n i t i e s f o r s y n f u e l s i n the coming decades.



2.

PENICK E T A L .


47

» Ethylene for Chemicals

Methanol

Methanol To Olefins

çr,Pr,c = Olefins 4

MTG 371-427°C 172-2068 kPa

Gasoline

MOGD Distillate 204-316°C 3447-5516 kPa F i g u r e 15.

P r o d u c t i o n o f e t h y l e n e , g a s o l i n e , and d i s t i l l a t e

methanol.

American Chemical Society Library 1155 16th St., N.W.

In Chemical Reaction Engineering—Plenary Lectures; Wei, J., et al.; Washington, D.C. 20036 ACS Symposium Series; American Chemical Society: Washington, DC, 1983.

from


48

Literature Cited 1. Meisel, S. L., McCullough, J. P., Lechthaler, C. Η., and Weisz, P. B., CHEMTECH, 1976, 6, 86-9. 2. Kokotailo, G. T., Lawton, S. L., Olson, D. Η., and Meier, W. Μ., Nature, 1978, 272, 437-8. 3. Chang, C. D., and Silvestri, A. J., J. Catal., 1977, 47, 249-59. 4. Fitch, F. B., and Lee, W., "Methanol-to-Gasoline, An Alternative Route to High Quality Gasoline," presented at the International Pacific Conference on Automotive Engineering, Honolulu, Hawaii, Nov., 1981. 5. Garwood, W. E., "Conversion of C -C Olefins to Higher Olefins Over Synthetic Zeolite ZSM-5," presented at Am. Chem. Soc. Mtg., Las Vegas, Nevada, March, 1982. 6. Meisel, S. L. and Weisz, P. B., "Hydrocarbon Conversion and Synthesis Over ZSM-5 Catalysts", presented at Advances in Catalytic Chemistry II Symposium, Salt Lake City, Utah, May, 1982. 2


ENGINEERING

RECEIVED March 10,

10

1983


Development of the Methanol-to-Gasoline Process - American

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