methanol-to-hydrocarbons: process technology

Microporous and Mesoporous Materials 29 (1999) 49–66

Review

Methanol-to-hydrocarbons: process technology

Frerich J. Keil *Technical University of Hamburg–Harburg, Department of Chemical Engineering, Eissendorfer Str. 38,

D-21073 Hamburg, Germany

Received 4 February 1998; received in revised form 30 June 1998; accepted 15 July 1998

Abstract

This review presents methanol-to-hydrocarbons processes which have reached industrial applications, either on acommercial or on a pilot plant scale. The determination of kinetic expressions for various methanol conversionreactions is given. The processes discussed are: Mobil’s methanol-to-gasoline (MTG) plant in New Zealand, thefluidized bed MTG and methanol-to-olefins process; Mobil’s olefin-to-gasoline/distillate (MOGD) process; the MTOplant developed by UOP and Norsk Hydro; Haldor Topsøe’s TIGAS process. The developments of a liquid phasedimethyl ether synthesis (LP-DME) process by the Ahron University are also presented. © 1999 Elsevier Science B.V.All rights reserved.

Keywords: Commercial plants; Kinetics; Methanol-to-gasoline; Methanol-to-hydrocarbons; Methanol-to-olefins; Review

1. Introduction Mobil Oil’s Central Research Laboratory inPrinceton were attempting to methylate isobutenewith methanol in the presence of ZSM-5. NeitherMobil’s novel synthetic gasoline process, basedreaction proceeded according to expectation.on the conversion of methanol to hydrocarbonsInstead, aromatic hydrocarbons were found.over zeolite catalysts, was the first major newMobil’s Central Research team tried to find outsynfuel development in the 50 years since thewhether methanol could serve as a precursor to adevelopment of the Fischer–Tropsch process. This‘C1 olefin’ in alkylating isobutane, to form, pre-process is known as the methanol-to-gasolinesumably, neopentane. ZSM-5 was the first catalyst(MTG) process. It provided a new route fromtried for this hypothetical reaction. An equimolarcoal or natural gas to high-octane gasoline.mixture of methanol and isobutane was passedAccording to Chang and Silvestri [1], two teamsover HZSM-5. Methanol was quantitatively con-of Mobil scientists working on unrelated projectsverted, whereas only about 27% of isobutanediscovered by accident the formation of hydro-reacted. An experiment carried out with purecarbons from methanol over the synthetic zeolitemethanol also showed a complete conversion ofZSM-5 [2,3]. The group at Mobil Chemical inmethanol. A careful material balance revealed thatEdison, New Jersey, had been trying to convertthe overall reaction stoichiometry could be repre-methanol to ethylene oxide, while workers atsented as

* E-mail address: [email protected] (F.J. Keil ) CH3OH�[CH

2]+H

2O (1)

1387-1811/99/$ – see front matter © 1999 Elsevier Science B.V. All rights reserved.PII: S1387-1811 ( 98 ) 00320-5

50 F.J. Keil / Microporous and Mesoporous Materials 29 (1999) 49–66

where [CH2] is the average composition of the field). The ICI low-pressure methanol process wasemployed. Two trains, each capable of producinghydrocarbon product. More detailed investigations

suggested that the main reaction steps are: 2200 tons per day were installed. The methanolproduct is passed to the MTG plant. In 1986 thestartup phase of the project was completed with2CH

3OH P

+H2O

−H2O

CH3OCH

3�

−H2O

light olefinsfull commercial production with a capacity of570 000 tons of gasoline per year being achieved.�higher olefins+n/iso−paraffins (2)The final gasoline produced does not need further

+aromatics+naphthenesrefining and attains the quality of unleaded pre-mium gasoline. Owing to some reasons presentedAs can been seen from the reaction scheme,

methanol is first dehydrated to dimethylether below, from 1981 and 1984, Mobil, UnionRheinische Braunkohlen Kraftstoff AG (RBK )(DME). The equilibrium mixture of methanol,

DME and water is then converted to light olefins and Uhde (Dortmund, Germany) have operated ademonstration plant for a Fluid Bed Mobil(C2–C4). A final reaction step leads to a mixture

of higher olefins, n/iso-paraffins, aromatics and Process. The plant was located at the RBK facili-ties in Wesseling (Germany), and was operatednaphthenes. An interruption of the reaction leads

to a production of light olefins instead of gasoline. from December 1982 to the end of 1985. Thisplant has successfully demonstrated the perfor-An appropriate process for this purpose was devel-

oped by Mobil, the so-called methanol-to-olefin mance of the fluid bed reaction system for MTGand MTO technology [5,6 ]. Union Carbide alsoprocess (MTO).

Therefore, the discovery of the MTG reaction developed a process to convert methanol to olefinsin 1986 using a silicoaluminophosphate (SAPO)was an accident. This discovery gave occasion to

a tremendous amount of detailed investigations of catalyst. The olefins yield exceeded 90%, and theyreport that the process could be modified for highthe reaction mechanisms and optimization of the

catalysts. Furthermore, new types of zeolites for ethylene and propylene yield (about 60%) [7–10].As the oil price dropped again over the 1980sthe MTO and MTG reaction were synthesized. A

review of these investigations is presented by further developments of commercial processeswere stopped for the time being. Nevertheless,Stocker [4].

The oil crisis 1973 and the second oil crisis in investigations on a bench scale were pursued, andapplications for patents are still submitted.1978 initiated the development of a commercial

MTG process. In response to a request from the As Stocker has reviewed the details of catalystand reaction mechanisms of the methanol-to-New Zealand Government, Mobil Research and

Development Corporation built a 640 l per day hydrocarbons (MTHC) processes, especially theMTG, MTO and MOGD (Mobil’s olefin-to-gaso-(four barrels per day), fixed-bed pilot plant to

demonstrate the feasibility of the gas-to-gasoline line and distillate) processes, the present paper willfocus on the technology of the MTG, MTO and(GTG) process. Reports to New Zealand’s Liquid

Fuels Trust Board (LFTB) by Lurgi and Badger MOGD processes. Some new developments, suchas direct conversion of methane to fuels will beindicated confidence that the process would scale

successfully from the pilot plant to commercial discussed. First, some kinetic equations will bepresented. Second, the technology of the respectivesize. The project concepts were developed under

the terms of a 1980 Government Mobil processes will be discussed.Memorandum of Understanding. Mobil wasresponsible for the overall project management,Bechtel acted as Project Services Contractor. Davy 2. KineticsMcKee, Foster Wheeler and New ZealandEngineering Consultants contributed to the design For the design of chemical reactors the kinetic

expressions must be known. They should enableand engineering of the project. In a first step,methanol is synthetized from natural gas (Maui the designer to simulate various reactor types and

51F.J. Keil / Microporous and Mesoporous Materials 29 (1999) 49–66

various modes of operation. The kinetic models Autocatalysis was supported by measurements exe-cuted by Chang et al. [16 ] and Ono et al. [17,18].can be grouped into two main classes: (a) lumped

models which are a compromise between simplicity Chang [19] extended the model using the followingassumptions:and representation of the reality of the process;

and (b) detailed models that take into account (1) Methanol and DME are always at equilibriumand can be treated as a single kinetic species.individual reaction steps. In general, it is very time

consuming or even nearly impossible to find the (2) Generation of the reactive intermediate is firstorder in oxygenates.kinetic expressions of type (b). For design purposes

type (a) kinetics will do in most cases. These kinetic (3) Consumption of the reactive intermediate isfirst order in oxygenates.expressions have to cover the whole range of

reactor operation with respect to temperatures, (4) Olefins can be treated as a single kineticspecies.pressures and feed compositions. It was found in

the 1970s that over a wide range of conversions (5) Disappearance of olefins is first order inolefins.the initial step of ether formation is much more

rapid than the subsequent olefin-forming step, and These assumptions lead to the following set ofreactions:is essentially at equilibrium [11,12]. Thus the

equilibrium oxygenate mixture can be convenientlytreated as a single kinetic species. Based on these A�

k1 B (7)

facts, Voltz and Wise [13] have developed a lumpedkinetic model which described the rate of

A+B�k2 C (8)methanol–DME disappearance in process and

pilot plant studies of methanol conversion to gaso-line. Chang and Silvestri [14] have postulated a

B+C�k3 C (9)mechanism of hydrocarbon formation from

methanol–DME. They have supposed a concertedbimolecular process involving carbonoid interme- C�

k4 D (10)

diates. The intermediates then undergo sp3 inser-tion into C–H bonds, forming higher alcohols or where A represents the oxygenates, B :CH2 groups,ethers, which can dehydrate to form olefins and C olefins, and D paraffins+aromatics. The formalwhich can add :CH2 to form higher olefins. Chen kinetics are:and Reagan [15] discovered that the oxygenatedisappearance is autocatalytic over ZSM-5. They −dA/dt=k

1A+k

2AB (11)

proposed the following scheme:dB/dt=k

1A−k

2AB−k

3BC (12)

A�k1 B (3) dC/dt=k

2AB+k

3BC−k

4C (13)

Assuming the steady-state condition for B andA+B�

k2 B (4) eliminating time resulted in an expression like this:

−du/dA=1/A[(1+K1u)/(2+K

1u)(1−K

2u)+u]B�

k3 C (5)

(14)where A represents oxygenates, B olefins, and C

where u=C/A, K1=k3/k2 and K2=k4/k1.aromatics+paraffins. The rate of oxygenate disap-Anthony [20] improved Chang’s kinetic expres-pearance is first order in oxygenates [13].

sion. Doelle et al. [21] studied both sorption andThe experimental data could be fitted accordingreaction kinetics of methanol and DME conversionto the expressionover ZSM-5. In the range between 115 and 200°Cthe kinetics of methanol conversion followed the−dA/dt=k

1A+k

2AB (6)


rate law catalyst. The following two lumped models wereable to treat the experimental data consistently:

r=k1PCH

3O/(1+k

2PH

2O) (15)

Model I:Schipper and Krambeck [22] have obtainedresults in a pilot plant with an adiabatic fixed-bedreactor, which was operated under reaction–regen- A�

k1 B (21)

eration cyles and under conditions in which therewas irreversible deactivation. They defined catalyst

A+B�k2 B (22)activity for a given time on stream, b, as the

product of two activity terms, which correspondto the remaining activity due to permanent deacti- A+C�

k3 D (23)

vation, a, and to the remaining activity due toreversible coking deactivation, f:

andb=a · f (16)

B+E�k4 F (24)The reaction rate of each individual step, r

i, is

the product of the reaction rate for the freshcatalyst r

ioand the activity b: C+E�

k4 F (25)

ri=r

io· b (17)

D+E�k4 F (26)For the irreversible deactivation kinetics Schipper

and Krambeck [22] proposed an empirical equa-tion similar to that used for permanent deactiva-

Model II:tion in catalytic cracking:

da/dt=−K

a0

exp(−Ea/RT ) am ; m>1 (18) A�

3k1 B+C+D (27)

The loss of activity due to coke deposition isexpressed as: B+E�

k2 F (28)

df/dt=−Kcokbn ; n>1 (19)

C+E�k3 F (29)By combination of the previous equations, a total

deactivation rate is obtained

db/dt=a{−Kcokbn−Km−2a

b} (20) D+E�k4 F (30)

This equation indicates that the rate of changewhere A¬Methanol+DME, B¬ethene, C¬of total catalyst activity is related not only to thepropene, D¬butene, F¬paraffins.total activity of the catalyst, b, but also to the

The authors found that an exponential activityclean-burned activity, a. When the catalyst is fresh,function of the typea is quite high so that the rate of activity loss is

high. After the first reaction–regeneration cycle, a ki=k

ioexp(−b

iHc/W ) (31)

is much lower. Thus, the rate of overall aging islower on the second cycle. can be applied to the observed values, where

Hc/W refers to the cumulative average amount ofSedran et al. [23] tested the kinetic model byChang [19] including the modifications proposed hydrocarbons produced per unit mass of catalyst

that was generated by the catalyst correspondingby Anthony [20] over the 302–370°C temperaturerange. In a further paper Sedran et al. [24] com- to each run. Therefore, the total hydrocarbon

formation is responsible for catalyst deactivationpared different lumped kinetic models for metha-nol conversion with hydrocarbons on a ZSM-5 by coking.


Benito et al. [25] considered the effect of com- where a is the activity defined asposition on the deactivation. This model has beenproven to be suitable by means of a wide experi- a=r

i(t)/r

io(t=0) (43)

mental study carried out in an isothermal fixed-bed integral reactor. ZSM-5 with a Si/Al ratio of The kdi are the kinetic constants for deactivation

by coke formation for lump i. The experimental24 was employed. The temperature range was300–375°C. The best fitting model was the one results revealed that the following equation can

describe the change of activity with time:proposed by Schipper and Krambeck [22]:

−da/dt=(kdAXA+kdCX

C+kdDX

D)a (44)

MEOH/DME(A)�k1 light olefins (C) (32)

The kdi were found to be different. In a furtherpaper Gayubo et al. [26 ] state that the models

2C�k2 products (D) (33)

of Chen and Reagan [15] and Schipper andKrambeck [22] adequately fit the experimentalresults. The authors emphasize the following

A+D�k3 D (34)

aspects: (1) ethylene and propylene are identifiedas primary products present in the gas phase; (2)an oligomerization step of light olefins is estab-

C+D�k4 D (35)

lished; and (3) a methylation step of products isestablished. Although the kinetic model ofSchipper and Krambeck [22] is slightly more com-The light olefins (ethylene, propylene) can poly-

merize to form products in the gasoline boiling plex than the one of Chen and Reagan [15], it ismore suitable because its kinetic scheme is closerrange D. Benito et al. [25] found the following

kinetic constants: to the real mechanism of the MTG process.Bos et al. [27] developed a kinetic model for the

methanol-to-olefins process [MTO], based on ak1=0.733×1013 exp(−33358/RT ) (36)

SAPO-34 catalyst. This catalyst makes ethene asa main product. The model is based on dedicatedk

2=0.127×108 exp(−17633/RT ) (37)

experiments in a pulse-flow, fixed-bed reactor. Thekinetic model was implemented in mathematicalk

3=0.204×1012 exp(−27987/RT ) (38)

models of various reactors for the estimation ofproduct selectivities and main dimensions. Thek

4=0.634×106 exp(−15855/RT ) (39)

experiments showed that the MTO reactions on afresh catalyst are very fast, with an overall first-The kinetic equations are:order rate constant of roughly 250 m3gas m−3cat s−1.The coke content of the catalyst is the main factorr

A0=−k

1XA−k

3XA

XD=dX

A/dt (40)

governing the selectivity and activity of the cata-lyst. In order to achieve an ethene-to-propene ratior

C0=k

1XA−k

2X2C−k

4XC

XD=dX

C/dt (41)

of 1 or higher, at least 7–8 wt.% of coke must bepresent on the catalyst. Consecutive reactionsThe X

irepresent weight fractions of lump i (on a

water-free basis), t is the space time. cannot be neglected. The net effects of these arean increase of ethene and propane selectivity andThe authors proposed a deactivation kinetic

model of separable functions dependent on the a decrease of mainly propene selectivity. Bos et al.[27] have tested several kinetic schemes.concentration of the three lumps of the reaction

scheme that are possible coke precursors: The final kinetic network of 10 first-order andtwo second-order reactions describes the experi-mental results satisfactory. The final reaction−da/dt=[∑ (k

diXi)]ad (42)


scheme for model discrimination was the (10) The aromatics undergo condensation.(11) The aromatics undergo alkylation withfollowing:

methanol.(12) The paraffins undergo demethanization form-

ing olefins and methane.The authors found a satisfactory agreement

between the experimental and calculated results.A further detailed model was developed byIordache et al. [29].

3. Fixed-bed methanol-to-gasoline (MTG) process

Ten years after Mobil announced a process[1,30] for converting methanol to high-octane gas-oline from non-petroleum sources, a commercialplant was in operation in New Zealand. The plantconverts natural gas from the Maui and Kapuni

(45) fields into methanol and then into ca. 700 000 tonsper day of gasoline via Mobil’s fixed-bed MTG

Eqs. (8) and (12) are of second order. The process. The gasoline produced is fully compatibleformation of ethene from propene is of first order with conventional gasoline. The conversion ofin methanol and propene. The rate of formation methanol to hydrocarbons and water is virtuallyof ethene from butene depends on butene and complete and essentially stoichiometric. The reac-methanol. tion is exothermic with a heat of reaction of about

Besides the lumped models, a few far more 1.74 MJ kg methanol−1. The adiabatic temper-detailed kinetic models were developed. Mihail ature rise is about 600°C. A simplified blocket al. [28] include 53 reactions which were grouped diagram of the MTG process is given in Fig. 1.into 12 subgroups. These 12 major steps are: Methanol is vaporized and fed into the fixed-(1) The etherification reaction takes place concur- bed DME reactor. In the DME reactor, the metha-

rently with the thermal decomposition of the nol is catalytically equilibrated to a mixture ofmethanol into hydrogen and carbon monox- dimethyl ether, methanol and water. The reactoride. The ether generates the carbene. contains a special alumina catalyst. The reaction

(2) The carbene attacks the ether and the alcohol, takes place at a reactor inlet temperature offorming light olefins. 310–320°C and about 26 bar pressure. Approx-

(3) The carbene attacks the olefins, forming imately 15–20% of the heat of reaction is releasedhigher olefins. in this first step, which is controlled by chemical

(4) The carbene attacks the hydrogen, forming equilibrium. The DME reactor effluent is mixedmethane. with recycle gas (see later) to moderate the temper-

(5) The light olefins generate carbeniums ions. ature rise over the second reactor, and then fed(6) The carbenium ions attack the light olefins into the ZSM-5 reactor. The inlet temperature

forming higher olefins (oligomerization). range of this reactor is 350–370°C. About 85% of(7) The carbenium ions attack the higher olefins the reaction heat is released in the conversion

forming paraffins and dienes. reactor. After cooling, the effluent is separated into(8) The carbenium ions attack the dienes forming three phases: gas, liquid water, and liquid hydro-

paraffins and cyclodienes. carbons. Most of the gas is recycled to the ZSM-5(9) The carbenium ions attack the cyclodienes reactor. The water contains a small amount of

oxygenates which are treated in a biological wasteforming paraffins and aromatics.


Fig. 1. Block diagram of the fixed-bed MTG process [31].

water treatment plant. The hydrocarbon product In an adiabatic reactor an S-shaped temperatureprofile along the reactor is formed. There is a veryis distilled. It contains mainly raw gasoline, dis-

solved hydrogen, carbon dioxide and light hydro- small change in the shape of the profile as thecatalyst ages. This means that coke formationcarbons (C1–C4). The non-hydrocarbons, C1–C3

and a part of the C4 hydrocarbons are removed downstream of the main reaction zone is low orits level does not appear to affect catalyst activity.by distillation to produce gasoline. The raw MTG

gasoline contains considerable amounts of durene After some time, the reaction zone approaches thereactor outlet and significant quantities of metha-(1,2,4,5-tetramethylbenzene). The heavy gasoline

treating unit (HGT) removes most of the durene nol appear in the water product. After regenera-tion, the catalyst can be returned to conversionwhich causes driveability problems, since the freez-

ing point of durene is relatively high (79°C). The service. Owing to band aging and catalyst deactiva-tion, methanol is processed at a continuouslytreated heavy gasoline is blended with other gaso-

line components to give specification finished higher effective space velocity as the cycle pro-gresses. Gasoline yields are greatest in the vicinitygasoline.

The development of the MTG process has been of methanol breakthrough. Catalyst aging leads toa change in hydrocarbon products. As aging pro-described in the literature, see for example Refs.

[31–33]. Bench-scale studies of the MTG process gresses, production of normal paraffins decreasesand the isoparaffin content increases. The aro-were executed for fixed-bed and fluid-bed reactors.

In Fig. 2 a two reactor configuration is shown matics content decreases, and the content of olefinsand naphthenes increases. The isoparaffins com-[31]. The fixed-bed reactor configuration is quite

simple and requires minimum scale-up studies. pensate for the aromatics, so that the gasolineoctane number varies only little with time. OneWith fresh catalysts, the reaction occurs over a

relatively small zone. The reaction front moves should keep in mind that the fixed-bed processresults in a slightly changing product composition.down the catalyst bed as the coke deposits first

deactivate the front part of the bed. Use of a Under MTG reaction conditions, the ZSM-5 cata-lyst undergoes two types of aging: a reversible losssufficient catalyst volume permits a fixed-bed

design in which on-stream periods are long enough results from coke formed on the catalyst as areaction by-product and the reaction productto avoid frequent regeneration cycles. A synthetic

crude methanol blend containing 83 wt.% metha- stream also causes a gradual loss of activity. Hightemperature enhances this type of deactivation. Asnol (commercial, pure) and 17 wt.% distilled water

was used as the charge. This is a typical water different segments of the catalyst bed are subjectedto varying degrees of water partial pressure andcontent of crude methanol made from natural gas.


Fig. 2. Schematic of fixed-bed bench-scale plant [31].

temperature, a permanent activity gradient results unit with a capacity of 0.636 m3 day−1 ofmethanol was built and operated.in the catalyst bed. In an 8-month aging test in

the bench-scale plant under realistic process condi- The major objective of the demonstration plantwas to verify the bench-scale results. The onlytions the following results were obtained [31]:

(1) Start-of-cycle (SOC ) gasoline yields increase different variable between a bench-scale unit anda commercial-size reactor is the linear velocity offrom cycle to cycle as a consequence of perma-

nent aging. the reactants. The catalyst bed diameter was 5 cm,the bed length was 3 m. The corresponding values(2) End-of-cycle (EOC) gasoline yields are fairly

constant. for the ZSM-5 reactors were 10 cm and 2.4 m. Thelinear velocities in these beds were about 10 times(3) As the catalyst ages, the change in gasoline

yield within a cycle decreases, and the cycle those in the bench-scale unit. Heat transfer alongthe reactor wall is negligible. The product yields,average gasoline yield increases. The cycle

lengths stabilize. selectivities, adiabatic temperature rise and theband aging behavior were nearly the same com-(4) The propane/propene ratio can be used to

track catalyst activity. pared with the bench-scale results. The cyclelengths were about 50% longer. This effect can be(5) SOC propane/propene yields show a sharp

decline over the first 50 days of operation related to the slower rate of movement of thecatalyst bed temperature profile. As the raw MTGfollowed by a gradual decline, whereas the

EOC values appear to approach a constant gasoline contains about 5.5 wt.% durene, a heavyMTG gasoline treatment (HGT) was developedvalue.

(6) Inherent gasoline selectivity did not change [34,35]. In the HGT process a 177+°C cut ofMTG gasoline, comprising primarily aromatics,throughout the aging test.

(7) The selectivity of the ZSM-5 catalyst for the is processed over a multifunctional metal–acidcatalyst. The following reactions occur: dispro-desired gasoline product increases as it ages.

This is a peculiarity of the MTG process. portionation, isomerization, transalkylation, ringsaturation, and dealkylation/cracking. The durene(8) After the bench-scale tests a demonstration


content is reduced to less than 2 wt.%. The gasoline methanol is partially dehydrated over a specialalumina catalyst to an equilibrium mixture ofproduced in the demonstration plant was used in

an automobile test fleet. These tests included inves- methanol, dimethyl ether and water. This reactiontakes place at a reactor inlet temperature oftigations with New Zealand-type cars, US cars,

Japanese and European cars. The tests were con- 310–320°C and 27 bar and releases 15–20% of theoverall heat of reaction. The DME reactor effluentducted under a wide range of ambient conditions.

The performance of MTG gasoline was equivalent is split into four parallel streams, mixed withheated recycled gas and passed into four parallelto that of conventional gasolines of similar

volatility. conversion reactors. The recycle stream controlsthe temperature rise. These reactors contain aAfter the successful MTG process development

a Joint Executive Committee (JEC), installed on ZSM-5 catalyst, where the conversion to hydro-carbons and water is completed. This type ofan agreement between the New Zealand

Government and Mobil in 1980, was to prepare a reactor is presented in more detail in Fig. 5.The conversion reactor inlet temperatures arereport which included a plan for the design, con-

struction, and operation of a plant to manufacture controlled to be 350–366°C. The inlet pressuresare 19–23 bar. In this reactor the main part ofgasoline, and an assessment of the viability of such

a project. Mobil was responsible for overall project reaction heat is released. Five conversion reactorsare installed of which only four are on-stream.management, Bechtel Petroleum Inc. was

employed as Project Services Contractor. The JEC The fifth reactor is thus either in the regenerationmode or on standby. The hot reactor effluent isreport was completed in July 1981. It concluded

that the project was technologically feasible and first cooled by generating steam, then fresh crudemethanol and recycle gas is heated. The reactorcommercially attractive. The synfuel plant was

mainly commissioned during 1985. In November effluent is then further cooled to 25–35°C at 16 barin a bank of water coolers and enters a three-1985 the first MTG gasoline was sent to the

Ministry of Energy tank farm near Port Taranaki. phase product separator, where gas, liquid hydro-carbons and water separate. The water phase,The New Zealand MTG plant is sited within

180 hectares of land at Motunui, Taranaki. It is which contains trace amounts of oxygenatedorganic compounds is passed to the water treat-designed to convert 52–55 PJ per annum of natural

gas into 570 000 tons per year of gasoline. The ment. The gas phase contains mostly light hydro-carbons, hydrogen, CO and CO2. This gas isdetails of the commissioning of the MTG plant

were outlined by Maiden [36], Bem [37], and recycled with the aid of a recycle compressor tothe conversion reactors. The raw gasoline whichChang [38]. Some aspects of plant design and

scale-up considerations were presented by Krohn contains dissolved hydrogen, carbon dioxide andlight hydrocarbons (C1–C4) is sent to theand Melconian [39]. A simplified block diagram

of the New Zealand plant is given in Fig. 3. de-ethanizer. The off-gases, including methane,ethane and some C3, together with a purge gasThe plant consists of three main process units:

two methanol trains, each capable of producing stream from the product separator are scrubbedin a sponge absorber in order to retain any gasoline2200 tons per day, and the MTG conversion plant.

The gas-to-methanol plant employs the ICI low- components. Then the gas is sent to the fuel gassystem. The de-ethanizer bottom product is sentpressure methanol process. Details concerning the

methanol plant are described by Allum and to the stabilizer where C3 and part of the C4components are removed overhead to the fuel gasWilliams [40]. The MTG unit is based on the

Mobil ZSM-5 catalyst. A more detailed presenta- system. C4/C5 components are withdrawn as asidestream. The bottom product is fed into ation of this plant is given in Fig. 4.

Depressurized crude methanol is pumped to gasoline splitter where it is separated into lightand heavy gasoline fractions. Each stream is cooledreaction pressure and vaporized against ZSM-5

reactor effluent before flowing into the first-stage and stored. As can be seen from Fig. 3, the heavygasoline fraction, which contains durene, is passeddehydration (DME) reactor. In this reactor crude


Fig. 3. Simplified block flow diagram of the New Zealand GTG plant [37].

Fig. 4. New Zealand MTG unit [40].


the fluidized bed reactor are given in Figs. 6 and7, respectively.

To study fluid dynamics and verify the mechan-ical design basis, a full-scale cold flow model(CFM ) was employed. This non-reacting modelproved very useful for optimizing baffle design andcatalyst circulation strategies. Several differentbaffle designs, horizontal and vertical arrange-ments, were tested in the CFM using differentexperimental techniques like tracer tests, capaci-tance probes, and bed expansion analysis. Theexperiments indicated that horizontal baffles areeffective in breaking bubbles. When the catalystfines concentration is sufficient (higher than 15%<40 mm), bubbles are small and the bed shows ahomogeneous appearance. Edwards and Avidan[41] employed a homogeneous, one-dimensionaldispersion model, combined with reaction kineticsto calculate the fluidized bed reactor performance.The Peclet numbers were calculated from SF6Fig. 5. ZSM-5 conversion reactor [39].tracer experiments. The superficial gas velocitiesranged from 0.3 to 1 m s−1. It turned out that theaxial dispersion model can be used to predict

to the HTG reactor. The durene level is cut down conversion in a turbulent fluid-bed reactor underthere to about 2 wt.%. The blended gasoline pro- the condition of an overall homogeneity of theduct has a research octane number (RON) of 92.2 bed. This is achieved in a turbulent bed of Groupand a motor octane number (MON) of 82.6. A powder by operation at superficial velocities

above 0.3 m s−1, with a minimum of fines. Ascale-up of the 15.9 m3 day−1 fluid-bed MTG pro-cess was possible from bench-scale without loss of4. Fluid-bed MTG and MTO processconversion efficiency. The fluid-bed process hasthe following advantages compared to the fixed-Mobil Research and Development Corp., Union

Rheinische Braunkohlen Kraftstoff AG and Uhde bed process:(1) Excellent heat transfer properties of a fluidizedGmbH (Dortmund, Germany) jointly designed,

engineered and operated a fluid-bed MTG demon- bed permit direct steam generation in coilsimmersed in the reactor.stration plant which was located in Wesseling near

Cologne in Germany. The project was partly (2) Continuous regeneration of the catalyst (con-stant catalyst activity) and a uniform bedfinanced by the American and German

Governments. Details of this project are described temperature result in a constant gasolinequality.by Gierlich et al. [5], Grimmer et al. [6 ], Penick

et al. [34], and by Edwards and Avidan [41]. (3) Transient temperature profiles during heat-upand cool-down are also uniform and stable.Bench-scale experiments on the fluid-bed MTG

process were executed at Mobil in the 1970s [42]. (4) The specific throughput in a fluid-bed systemis higher.Subsequently, a 0.1 m diameter, 7.6 m tall,

0.636 m3 day−1 pilot plant was successfully oper- (5) Higher octane numbers were found (seeTable 1).ated. From December 1982 to the end of 1985 a

15.9 m3 day−1 demonstration plant produced gas- (6) The yield of gasoline including alkylate is atleast 7.5% higher.oline in Wesseling. Details of the entire plant and


Fig. 6. Fluid-bed MTG demonstration plant [49].

(7) The durene content is lower (maximum5 wt.%)

(8) Liquid injection, a unique feature in the fluidbed, provides the flexibility to tailor the steambalance as per requirement.

(9) Specific investment cost is lower.A simplified scheme of the fluidized bed plant

in Fig. 6 shows that crude methanol is vaporizedand fed into the reactor. To improve the overallprocess economy, the water can be removed fromthe crude methanol. The heat of reaction can beused to generate high pressure steam. The heatexchanger can be operated inside or outside thereactor. Coils immersed inside the reactor turnedout to be most efficient. The catalyst is continu-ously withdrawn and regenerated by partiallyburning off the coke. The rate of catalyst circula-tion through the regenerator determines theaverage activity of the catalyst in the reactor. Abank of different cyclones remove catalyst dustfrom the reactor effluent. The hydrocarbon pro-ducts are processed in a gas fractionation unitto produce C+

5hydrocarbons, alkylation feed

(C3/C4 fractionation) and olefin recycle. Heavygasoline treatment will be required only if a further

Fig. 7. Fluidized bed reactor of the demonstration plant. reduction of durene content is necessary. A com-


mercial concept of the fluid-bed MTG process was ation. MOGD olefin product distribution is deter-mined by thermodynamic, kinetic, and shape-developed by Uhde [6 ]. At a gas price of

US$1 GJ−1 and a GTG plant of 6180 ton day−1 selective limitations. A large-scale MOGD test runwas executed in a Mobil refinery in 1981. Acapacity, 1 l of unleaded premium gasoline will

cost US$0.19 (data for 1987). The fluid-bed tech- commercially produced zeolite catalyst wasemployed. A simplified scheme of this process isnology is ready for commercialization. The plant

at Wesseling was also operated for the production presented in Fig. 8 [49].In general, four fixed-bed reactors, three on-lineof olefins at a pressure between 2.2 and 3.5 bar

and a temperature of about 500°C. At steady state and one in regeneration are used. The three on-linereactors are operated in series with interstageconditions the olefin yield was more than 60%,

although the catalyst was not tuned to olefin cooling and liquid recycle to control the heat ofreaction. The olefins feed is mixed with a gasolineproduction. Lurgi has considered a tubular fixed-

bed MTG process, which offers some advantages recycle stream and passed, after heating, throughthe three reactors. In order to generate a gasoline-over the New Zealand fixed-bed process [43].

However, the fluid-bed MTG process has over- rich stream for recycle to the reactors, a fraction-ation is used. The recycle improves distillate selec-taken this concept.tivity. The MTO and MOGD process can becombined. A possible process flow-sheet is pre-sented in Fig. 9.5. The methanol-to-olefins (MTO) and the mobil

olefin-to-gasoline/distillate (MOGD) process High-octane MTO gasoline is partially split offbefore the MOGD section and is later blendedwith MOGD gasoline. The raw distillate is hydro-As light olefins are intermediates in the MTG

reaction scheme, methanol can also be employed treated and can be fractionated into various pro-ducts. Typical distillate and gasoline yields fromto produce light olefins. Higher reaction temper-

atures, lower pressures and a high ratio of ( lower the olefins yield obtained in the 15.9 m3 day−1MTO plant at Wesseling are 50/50 w/w. This ratioacidity zeolites) favor the production of light

olefins. Chang et al. [44,46 ], Chang [45], can be variied considerably. The gasoline is olefinicand aromatic, and of better quality than FCCSchoenfelder et al. [47], and Tshabalala and

Squires [48] report bench-scale measurements on gasoline. The MON and RON are 93.0 and 85.0,respectively. The durene content is very small. Itsthe MTO process. As already mentioned, an MTO

demonstration plant was also operated. Union physical properties, such as flash point, boilingrange and viscosity, are comparable with conven-Carbide developed a process to convert methanol

to olefins using a SAPO catalyst. The olefins yield tional distillate fuels. MOGD diesel has a densityof 0.8 instead of 0.86. MOGD can also be used asexceeded 90%. The process can be modified for

high ethylene and propylene yield (about 60%) jet fuel.[7,8,10]. The basic flow-sheet of the MTO processis the same as that of the fluidized bed MTGprocess. Avidan [49] described the results obtained 6. UOP/HYDRO MTO processin the Wesseling plant.

A further development is the Mobil olefins-to- UOP (Des Plaines, Illinois) and Norsk Hydro(Oslo, Norway) developed a new gas-to-olefinsgasoline/diesel (distillate) (MOGD) process

[50,51]. In this process, gasoline and distillate (GTO) and MTO process which produces a veryhigh yield of ethylene (48%) and propylene (33%).selectivity is greater than 95% of the olefins in the

feed and gasoline/distillate product ratios range As a catalyst an attrition resistant SAPO-34 wasemployed. A simplified process flowsheet is pre-from 0.2 to >100. In order to obtain high octane

numbers, shape-selectivity is tuned such that sented in Fig. 10 [52,53].The UOP/HYDRO MTO unit employs a fluid-mostly methyl-branched iso-olefins (C5–C20) are

produced. The C10 to C20 fraction needs hydrogen- ized-bed reactor coupled to a fluidized-bed regener-


Table 1Process conditions and product yields from MTG processes

Fixed-bed Fluid-bed

Methanol/water charge (w/w) 83/17 83/17Dehydration reactor inlet temperature (°C) 316 –Dehydration reactor outlet temperature (°C ) 404 –Conversion reactor inlet temperature (°C) 360 413Conversion reactor outlet temperature (°C) 415 413Pressure (kPa) 2170 275Recycle ratio (mol/mol ) charge 9.0 –WHSV (h−1) 2.0 1.0

Yields (wt.%) of methanol chargedMethanol+ether 0.0 0.2Hydrocarbons 43.4 43.5Water 56.0 56.0CO, CO2 0.4 0.1Coke, other 0.2 0.2

100.0 100.0

Hydrocarbon product (wt.%)Light gas 1.4 5.6Propane 5.5 5.9Propylene 0.2 5.0Isobutane 8.6 14.5n-Butane 3.3 1.7Butenes 1.1 7.3C5+gasoline 79.9 60.0

100.0 100.0

Gasoline (including alkylate)[RVP-62 kPa (9 psi)] 85.0 88.0LPG 13.6 6.4Fuel gas 1.4 5.6

100.0 100.0Gasoline (RON) 93 97

RVP=Reid vapour pressure.WHSV=weight hourly space velocity.RON=research octane number.LPG=liquified petroleum gas.

ator. The heat of reaction is controlled by steam to the product recovery section, which includes ademethanizer, a deethanizer, a C2 splitter, a C3generation. The catalyst is sent continuously to

the regenerator, where the coke is burned off and splitter, and a depropanizer. Polymer-grade ethy-lene and propylene are produced from these frac-steam is generated to remove the heat resulting

from burning. tionation columns along with methane, ethane,propane, and C4 product streams. TheAfter heat recovery, the reactor effluent is

cooled, and some of the water is condensed. After UOP/HYDRO MTO process can be economicallyviable in different scenerios [52,53]:compression, the effluent passes through a caustic

scrubber to remove CO2 and to a dryer to remove (1) Production of methanol at a remote gas fieldsite and transportation of the methanol to anwater. The reactor section is quite similar to the

Mobil/Uhde process. The effluent then proceeds MTO plant located at the olefins user’s site.


Fig. 8. MOGD demonstration plant [49].

Fig. 9. MTO/MOGD process [6 ].


Fig. 10. UOP/Hydro MTO process for polymer-grade products [53].

(2) An integrated GTO complex at the gas field costs determine a high proportion of the cost ofproduction due to the low energy price. Thesite and transportation of olefins or polyolefins

products to customers. TIGAS process was developed for just this pur-pose. A block scheme is presented in Fig. 11.(3) Increased olefins production and feedstock

flexibility at an existing naphtha or ethane– In the TIGAS process the two process steps,MeOH synthesis and the MTG process, are integ-propane cracker facility by installing an MTO

reactor section and feeding into a revamped rated into one single synthesis loop without isola-tion of MeOH as an intermediate. Thecracker fractionation section.

(4) A smaller MTO unit using methanol produced experimental program for the TIGAS processlasted 3 years and was terminated in January 1987.in a single-train methanol plant to meet the

local demand for olefins or polyolefins or both. A demonstration plant at Houston was operatedfor 10 000 h. The purpose of the process develop-(5) The UOP/HYDRO MTO process and catalyst

have been successfully demonstrated and are ment work on the integrated gasoline synthesiswas to modify the three process steps – synthesiscurrently available for license.gas production, oxygenate synthesis and the MTGprocess — in order to be able to operate all stepsat the same pressure and the last two steps in one7. Haldor Topsøe TIGAS process, Akron

LP-DME-to-gasoline process (DTG) single synthesis loop [54]. By selecting combinedsteam reforming and autothermal reforming forthe synthesis gas production, and by using a multi-Topsøe has developed a low investment process

for the conversion of natural gas to gasoline [54]. functional catalyst system, producing a mixture ofoxygenates instead of only MeOH, the front-endAs many future synthetic fuel plants will be built

in remote areas where the price of natural gas is and the oxygenate synthesis can operate at thesame pressure (~20 bar). The TIGAS processvery low and not related to gasoline, low invest-

ment cost is essential, as the investment-related avoids the compression of syngas to about


Fig. 11. Topsøe TIGAS process [54].

50–100 bar required by a conventional methanol Parkyns et al. [58] discuss some further aspects ofsynfuel production.plant. This reduces the capital and operating costs

of the combined synthesis and conversion loops.The overall reaction of the TIGAS process is

8. Conclusion3CO+3H2PCH

3OCH

3+CO

2(46)

The DME can be converted into hydrocarbon A broad variety of well-tested processes for theproducts in a separate MTG reactor. The opera- production of hydrocarbons from methanol istion of the MTG process and the oxygenate synthe- available. Their future usage is determined bysis in one loop will then only call for minor changes natural gas and methanol prices.in the MTG process. A separation unit leads tothe products.

The University of Akron has developed a pro-Referencescess which converts syngas directly to DME using

LP-DME synthesis [55]. One-step conversion of[1] C.D. Chang, A.J. Silvestri, CHEMTECH 10 (1987) 624.syngas to DME improves the per-pass conversion[2] C.D. Chang, Catal. Rev.-Sci. Engng 25 (1983) 1.and reactor productivity over syngas to methanol.[3] C.D. Chang, Catal.-Rev.-Sci. Engng 26 (1984) 323.

A dual catalyst system is based on a combination [4] M. Stocker, Microporous Mesoporous Mater. (1999) 3of Cu/ZnO/Al2O3 catalyst and gamma-alumina (this issue).

[5] H.H. Gierlich, K.H. Keim, N. Thiagarajan, E. Nitschke,catalyst. This conversion offers some advantages.A.Y. Kam, N. Daviduk, Paper presented at the 2nd EPRIFirst, the liquid phase is lean in methanol becauseConference Synthetic Fuels – Status and Directions, Sanof in-situ conversion to DME over gamma-alu-Francisco, CA, 1985.

mina. Second, water produced by both methanol [6 ] H.R. Grimmer, N. Thiagarajan and E. Nitschke, in: D.M.synthesis (CO2 hydrogenation) and DME synthesis Bibby, C.D. Chang, R.W. Howe, S. Yurchak (Eds.),

Methane Conversion, Studies in Surface Science and(methanol dehydration) constantly shifts the for-Catalysis, vol. 36, Elsevier, Amsterdam, 1988, p. 273.ward water gas shift reaction. This is the special

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DME. The one-step conversion of syngas to DME [9] S.W. Kaiser, Arabian J. Sci. Engng 10 (1985) 361.improves the volumetric productivity by as much [10] G. Pop, G. Musca, D. Ivanescu, E. Pop, G. Maria, E.

Chirila, O. Muntean, Chem. Ind. 46 (1992) 443.as 100% over that of syngas to methanol conver-[11] C.D. Chang, J.C.W. Kuo, W.H. Lang, S.M. Jacob, J.J.sion. This is because conversion of syngas to DME

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methanol-to-hydrocarbons: process technology

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