Polymerization of Hydrocarbon Gases to Motor Fuels

Polymerization of Hydrocarbon Gases to Motor Fuels P. A. MASCHWITZ and L. M. HENDERSON

Downloaded by UNIV OF CALIFORNIA SAN DIEGO on June 2, 2015 | http://pubs.acs.org Publication Date: January 1, 1951 | doi: 10.1021/ba-1951-0005.ch009

The Pure Oil Co., Chicago, Ill.

Processes for the polymerization of hydrocarbon gases to motor fuel were developed to a commercial level in the early 1930's. Thermal polymerization plants, employing temperatures of 9 0 0 ° to 1200° F. and pressures of 60 to 3000 pounds per square inch, were developed first, closely followed by catalytic units operating at temperatures of 280° to 4 7 5 ° F. and pressures of 200 to 1200 pounds per square inch. Currently, thermal polymerization finds its greatest application in combination with thermal reforming of naphtha. Catalytic polymerization has proved highly successful, as is indicated by the fact that one company alone has licensed over 150 plants to date.

T h e past twenty-five years have seen a great increase i n the volume of hydrocarbon gases produced i n petroleum refineries as a consequence of the enlarged cracking capacity of the industry. E a r l y during this period refiners used cracked and natural gases (also being produced i n increasing volume) as fuel and i n some cases wasted the excess gas. W i t h the availability of this low priced charge stock as an incentive, research was directed simultaneously b y several organizations toward the conversion of these low molecular weight hydrocarbons to motor fuel b y polymerization. Development work first bore fruit i n 1931 when a unit capable of producing up to 100 barrels a day of gasoline from refinery gas started operation at the Toledo Refinery of The Pure O i l C o . (24). T h e process was subsequently licensed b y Alco Products, Inc. This was followed i n 1934 b y a unit of about 1000 barrels a day at the Alamo Refinery of the Phillips Petroleum C o . (#), Borger, Tex., operating on either cracked or natural gas fractions. T h e Polymerization Process Corp. was formed to license the process which incorporated the results of research by Phillips Petroleum, Standard O i l (Indiana), Standard O i l (New Jersey), Texas Co., and M . W . Kellogg C o . These were both thermal units relying on comparatively high heat and pressure to effect conversion. I n 1935 catalytic polymerization of olefins to motor fuel boiling fractions was accomplished on a commercial level at the East Chicago R e finery of the Shell O i l Co., using the polymerization process of the Universal O i l Products C o . and employing more moderate conditions. The use of catalytic polymerization spread uniformly throughout the following decade until almost every refinery having cracking facilities employed polymerization of cracked gases. There are now more than 150 Universal O i l Products (U.O.P.) catalytic polymerization plants i n operation. A combination of the thermal polymerization process and the U . O . P . catalytic process was introduced i n 1937 at the Shamrock O i l and Gas Co., Sunray, Tex. (23). I n 1934 the Shell Development C o . introduced the cold acid process (13), which selectively polymerizes isobutene, using sulfuric acid as catalyst. The hot acid process was also developed b y them and differed from the cold acid process i n polymerizing all C 4 olefins. B o t h products are predominantly the dimer. The cold acid process produces a large pro83

In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.


84

ADVANCES IN CHEMISTRY SERIES

portion of almost pure iso-octene which readily hydrogenates to iso-octane. The hot acid process produces a mixture of isomeric octenes which i n turn m a y be hydrogenated to hydrocodimer. Iso-octane and hydrocodimer from sulfuric acid plants, as well as from some U . O . P . catalytic plants designed to selectively polymerize isobutenes or butènes, were of notable importance during W o r l d W a r I I as components of aviation gasoline. During 1931-41, additional plants of a l l these types were built, b u t the logical combination of catalytic dehydrogenation of paraffins to olefins, followed b y catalytic polymerization of olefins, was not commercially realized although the literature abounds with references t o dehydrogenation catalysts (6). I t was not until 1941, under the stress of war, that the H o u d r y process for dehydrogenating isobutane to isobutene was first used b y Imperial Chemical Industries i n England. Today, dehydrogenation of propane or butane still finds no place as a step i n making motor gasoline. A further application of polymerization lies i n the valuable combination of the thermal polymerization process with thermal reforming of naphtha for octane number appreciation. A n outgrowth of such combination was first used at the Gulf Oil Corp.'s Pittsburgh refinery i n 1936 and called Polyforming. A gas reversion process was brought out b y Phillips Petroleum C o . and installed i n their Kansas C i t y Refinery i n 1937. These processes, which will be explained later, largely replaced thermal polymerization and found considerable application until the advent of catalytic reforming i n the last few years.

Thermal Polymerization Processes Thermal polymerization processes were developed simultaneously and independently by the Phillips Petroleum C o . and The Pure O i l Co., the former dealing more particularly with the conversion of C3 and C 4 hydrocarbons containing a low proportion of olefins, and the latter a somewhat higher proportion. Operating condition limits for the former were 800 to 3000 pounds per square inch gage and 900° to 1100° F . with the flow arrangement shown i n Figure 1. The latter company employed conditions of 60 pounds per square inch gage and 1150° F . to 1200 pounds per square inch gage and 1100° F . with flow a r rangement as i n Figure 2. T h e exothermic conditions of the polymerization with olefins make i t preferable to have the reaction coil i n a separate setting to allow controlled heat

GASOLINE -FRACTIONATING

«AS0LINE — J

Figure 1.

Unitary Polymerization Process


MASCHWITZ AND HENDERSON—POLYMERIZATION OF HYDROCARBON GASES

85

RECYCLE GAS

»6AS TO ABSORBER

HEATING

COIL

REACTION /

COIL

TAR SEPARATOR

7s


t- RECEIVER

Ψ2

^ L

-CHARGE

Figure 2.

-POLYMER GASOLINE

TO TAR STRIPPER

rRIPPER

1

«

Thermal Polymerization of Olefin-Containing Gas

dissipation. A t the lowest pressure and highest temperature when charging a gas con taining 3 5 % C and C4 olefins (17), the total liquid product is largely aromatic and has the true boiling point ( T . B . P . ) distillation shown i n Figure 3, curve A, having distinct plateaus a t benzene and toluene. B y increasing the pressure to 400 pounds per square inch gage and reducing the temperature to 1070° F . , a polymer having a T . B . P . distillation, curve B, is made which no longer exhibits marked aromaticity. F o r com parison the T . B . P . distillation curve of the crude polymer produced from a C8-C4 feed containing 6 4 % olefins, using a phosphoric acid catalyst at 400° F . and 200 pounds per square inch gage is shown as curve C on Figure 3 {12). I n the operating range required to produce maximum yields of motor fuel (which increases with operating pressure) the two thermal processes become technically indistinguishable. Although liquid polymers boiling i n the gasoline range can be made from methane, ethane, and ethylene, the yield has been too low and the operating conditions too severe to make gasoline from them economically. In predicting yields, an empirical mechanism has to be resorted to since the chemistry of the pyrolysis of mixtures of C3 and C4 hydrocarbons is still at least partly conjecture. A method for computing ultimate yields and product distribution worked out b y H . C . Schutt and others of Stone & Webster Engineering Corp., heretofore unpublished, involves the estimation of the hydrogen content of the residue gas, the total liquid polymer as a function of the operating conditions, subsequent simultaneous application of an over-all material and hydrogen balance to determine the yield of gasoline and tar products from a given charging stock. Weight % total polymer is found by solving 3

PT

= 100 X

(Hg (HG -

HF) HP)

where Ρ τ = weight % total polymer; Ho = weight % hydrogen i n Crfree residue gas; Hp = weight % hydrogen i n C -free polymer; and HF = weight % hydrogen i n net fresh feed. Weight % hydrogen i n the Crfree residue gas is obtained from Figures 4, 5, and 6 which relate weight % C i n the furnace feed to mole % C2 i n the Crfree residue gas, operating pressure to ethylene to ethane ratio, and operating temperature to hydrogen to 4

2


86


methane ratio, respectively. Although the correlations given i n Figures 4 and 6 apply quite generally, the validity of the relationship i n Figure 5 is limited to the temperature range prescribed thereon; however, practical operating temperatures are included.

140


300

tto 180

100 •0

1

1»

20

30

40

S9

β0

79

β9

M

% OVEftMEAO

Figure 3.

True Boiling Point Curves of Polymers A and 8 « Thermal C = Catalytic

Figure 7 permits the estimation of the hydrogen content of the Crfree polygasoline as a function of operating pressure and weight % total Crfree polymer per pass on com bined throughput including the C recycle. The specific gravity of the Crfree 400° F . end point gasoline is also given. T h e relationship clearly illustrates the beneficial effect of pressure on yield a n d the deceleration of this benefit as pressure increases, so that a dis creet limit is soon reached. The range of total polymer yield per pass is limited by the rate of carbon deposition on the high end and the size of equipment on the low end. The range given is that residing between these two uneconomic extremes. 2

WT % V "

Figure 4.

'UKNACC 'CEO

Ethane-Ethylene in Residue Gas

T a r is defined as the part of the polymer product boiling above 400° F . end point gasoline, and the weight % may be determined from Figure 8 as a function of total polymer In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.


87

per pass and operating pressure for net fresh feed compositions shown on Figure 8. F u r ther, Figure 8 illustrates the penalty of high conversions b y overpolymerization. Pressure promotes a favorable gasoline to tar ratio. The hydrogen and material balances may be completed b y finding the hydrogen content and specific gravity of the tar as a function of operating temperature from Figure 9. Time and temperature are to some extent interchangeable to attain a certain con version per pass. Preferred reaction time is, however, i n the 50- to 100-second range. I ~" ASSOCIATED TEMP, β PRESSURE* PRESS. LB. SQ.M.TEMP. "F TEMP TOLERANCE


l§S8

ISOO 1790 2000

m •m 1039 1030 1029

î

ί

. *

Ï 7 β

i

\9

-T867S ΠΛδ" IBoTP 0PERATIN0 PRESSURE- L B / f e »Ν·

Figure 5.

Residue Gas Composition

Gas Reversion Process. T h e reforming of n a p h t h a for octane improvement is conducted under temperatures a n d pressures of the same order as those for t h e r m a l p o l y m e r i z a t i o n . Whereas reforming is a once-through operation, the t e r m P o l y forming applies when the C fraction produced i n the reforming process and that part of the C fraction not required for gasoline blending are recycled, i n a sense combining re forming and thermal polymerization. Polyforming can be extended to include outside C and C stocks. 3

4

3

4

3

* ο

S κ 1

4Γ"

, >ooV

Figure 6.

Hydrogen-Methane Ratio

I n the gas reversion process the recycle and outside C3-C4 stocks are heated separately for partial conversion before admixture with the naphtha. T h i s bridges thé difference i n reaction velocity between the two types of charge and is helpful since the conversion rate of naphtha is approximately four times that of propane and twice that of butane, thereby decreasing the volume of C3-C4 recycle. Figure 10 shows a simplified flow diagram of a typical gas reversion operation. The extent to which outside C3-C4 stocks can be utilized is not limited, and the process can revert to thermal polymerization as the proportion is


88

ADVANCES IN CHEMISTRY SERIES ASSOCIATED TEMP PRESS. LB/SO IN. 500 750 1000 ISOO 2000

A PRESSURES TEMP. "F 1080 - 1100 1062 - 1078 1048 - 1082 1090- 1040 1020 - 1090

FRESH FEED> C C MIXTURE} 1 5 - 4 3 % UNSATURATION f l

MOTE- FOR OPERATING TEMP. OUTSIOE RANGE GIVEN ,A TEMP. COR RECTION MUST BE APPLIED. .73

Λ

= Τ Γ = «•«•MM*.


SPEC. GRAVITY .74 .7β

.78

—

^

S s

4

η

TOTAL POLYMER (C FREE) PER PASS WT.% ON FURNACE CHARGE

SPEC. GRAVITY

4

Figure 7.

39 57 A. P. I. GRAVITY

"55

Composition of G a s o l i n e — 4 0 0 ° F. End Point, C -Free 4

increased to the final exclusion of naphtha. Gas oil may be substituted for naphtha, par ticularly if of a refractory nature, such as recycle stocks from catalytic cracking. I n the combination of processes both the reforming severity and octane appreciation may be i n creased before rate of carbon deposition becomes a limiting factor. Also, the rate of con version of C3 fraction and C4 fraction to motor fuel is increased b y alkylation and syn thesis reactions with the naphtha.

WT. % TOTAL C FREE POLYMER PER PASS 4

Figure 8.

Tar Yields

Recent literature contains detailed correlations of the yield and quality of the gasoline product with operating conditions, quality of naphtha charge, and extent of outside addi tions of C and C fractions (7,14). The method of calculation employs several correlations of C and lighter fractions ex pressed as weight % of heater feed, known as "severity factor." The first correlation de termines the yield of C , C , C , and C to 400° F . end point fractions from the naphtha 3

4

2

3

4

5

6


MASCHwrrz AND HENDERSON—POLYMERIZATION O F HYDROCARBON GASES

89

alone, followed by others determining the yields of gasoline components from the C and C fractions from the naphtha and from outside gas which are added to those of the naphtha.


3

4

REACTION TEMPERATURE V

Figure 9.

Tar Composition

I n gas reversion the outside gas is subjected separately to mild thermal polymeriza tion, limited b y tar production, insufficient to advance materially the rate of carbon de position i n the coil after naphtha admixture. B y combining the yield mechanism given for thermal polymerization with that for Polyforming, the process may be evaluated. I n general, preferred operating conditions are 1020° to 1120° F . and 1500 to 2000 pounds per square inch gage i n the reversion section, followed b y 1025° to 1125° F . and 1000 to 2000 pounds per square inch gage i n the naphtha Polyform section. According to Bogk, Ostergaard, and Smoley (1) the y i e l d - C F R - A S T M octane number relationship of straight reforming, Polyforming, and gas reversion of 40 octane number straight r u n gasoline is given i n Figure 11.

CHARGE TO HEATER

3 ^

SATURATED Cj β C± CHARGE FROM NATURAL GASOLINE PLANT C , a C CHARGE FROM CRACKING STILLS' 4

QUENCH PUMP

Figure 10.

Gas Reversion Process


90


Gasoline boiling range distillates from thermal polymerization and gas reversion are not suitable for use as motor fuels without further treatment owing to high gum content and low induction periods. Fimshing is usually accomplished b y percolation through fuller's earth i n the vapor phase (Gray process) as shown i n Figure 10.


Catalytic Polymerization Processes The more significant commercial catalytic polymerization processes have proved to be phosphoric acid processes and sulfuric acid processes. The Universal O i l Products C o . process is based on the successful development of phosphoric acid catalysts b y Ipatieff, initially as a liquid (8, 9, 11), later on a solid a d sorbant (10). T h e first U . O . P . polymerization plants operated at pressures of about 200 pounds per square inch, temperatures of about 400° F . with a 100° F . temperature rise through the catalyst bed. A number of catalyst towers were used i n series for the catalyst had to be regenerated, usually one tower at a time while the remaining ones continued on stream. T h e regeneration was accomplished b y oxidizing the coke and heavy polymer deposit with a controlled concentration of oxygen i n an inert gas followed b y steaming of the catalyst for the purpose of restoring the water of hydration necessary for high conversion and extended life of the catalyst.

MO

ISO

ISO

110

100

to

to

70

OO

% 6AS0LINE YIELD 0Y VOLUME I A S E 0 ON NAPHTHA CHARGE

Figure 11. Yield-Octane Relationship for NaphthaGas Reversion and Reforming of Straight-Run Gasoline Gray and Hutchinson County, Tex., Crudes

B y 1939 a M i d g e t chamber-type polymerization process (20) had been developed that economically employed feed rates as low as 250,000 cubic feet per day, whereas the older units were uneconomical at feed rates below about 2,000,000 to 3,000,000 cubic feet per day. Midget units proved successful because i t was found that when operating the polymerization reaction at pressures of 500 pounds per square inch a sufficient amount of feed remained i n a dense phase to wash the catalyst clear of most of the heavy polymer, thus extending catalyst life and eliminating catalyst regeneration. Figure 12 is a simplified flow diagram of a chamber-type unit. I n this instance the feed is taken only from the cracking plant stabilizer overhead although some plants also include the absorber overhead gas i n the feed. I f a predominantly C3-C4 charge is polymerized the recovery section can be modified to yield a propane stream for liquefied petroleum gas sale as well as butane b y the use of a de-ethanizer (if required), a depropanizer, and a debutanizer. A further improvement was the development of the reactor type unit using multiple tube-and-shell reactors for better temperature control (26). T h i s type of reactor proved useful i n larger installations and for selective polymerization of either C3 or C4 olefins only. The larger reactor type units are so arranged that the steam produced i n the reactor shell by the exothermic reaction is used to preheat the feed to the proper inlet temperature. The tubes usually are from 2 to 6 inches i n diameter. In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.


91

CRACKING PLANT


1

CATALYST TOWERS

3>

- c CRACKEO GASOUNE^

Figure 12.

CAUSTIC β WATER WASH TOWERS

U.O.P. Catalytic Polymerization Process—Chamber Type

The principal operating variables i n the U . O . P . catalytic polymerization process are feed composition, pressure, temperature, space velocity, and water content of the feed. The type of feed that can be used varies widely. I n nonselective polymerization i n chamber-type plants a C3-C4 olefin content of about 20 to 2 5 % , is practical. When the olefin content of the fresh feed rises above 2 5 % , enough spent gas of low olefin content must be recycled i n order to limit the olefin content of the total feed. Reactor-type plants can employ a higher olefin content i n the feed, the exact value depending on the cooling efficiency of the reactor. The ease of polymerization of olefins, i n the order of increasing reactivity, is ethylene, propylene, 1-butylène, 2-butylene, and isobutylene. The presence of isobutylene a p parently accelerates the polymerization of n-butylenes, and the pre ence of butylènes has a similar effect on propylene (8,11, 25). Polymerization is accompanied b y a decrease i n volume and therefore is promoted by elevated pressures. Also, raising the pressure lowers the temperature at which satisfactory conversion can be obtained; and low temperatures are preferred because they minimize the formation of heavy polymers and coke, thus increasing catalyst life. Older units operated at about 200 pounds per square inch and 400° F . at the inlet, using cracked gases without further compression. The newer chamber-type units operate at about 500 pounds per square inch, inlet temperatures of 350° to 400° F . , and 450° to 475° F . at the outlet, with temperature control b y means of recycling and quenching at intermediate points with spent gas. Reactor-type units operate at 700 to 1200 pounds per square inch with about 400° F . inlet and limit the temperature rise to approximately 15° F . When selectively polymerizing C olefins, the temperature of a reactor-type unit may be as low as 280° to 375° F., depending on the pressure used and whether dimers only or also trimers are desired. The space velocity is generally chosen to give conversions of 85 to 9 5 % of the olefins charged. Higher conversions are possible, but uneconomical, as more catalyst is required and catalyst life is reduced. Space velocities, based on total feed, may range from about 0.3 to 0.5 gallon per hour per pound of catalyst or 4.5 to 7.0 cubic feet per hour per pound of catalyst, depending on olefin content of feed and other process variables. The addition of the proper amount of water to the feed is of great importance. A n 4



9 2

*EFFLUENT C

REACTOR/

^

ISOBUTYLENE POLYMER

^MIXER

rib CAUSTIC SODA SETTLER

EXTRACT/ SETTLER


fcb

4

EXTRACT SETTLER!

i b ^ r

7

CAUSTIC SOOA SETTLER

" Π

Τ _ F E E D CONTAININO ISOBUTYLENE

Figure 13.

It

CAUSTIC SOOA"

- 6 5 % SULFURIC

ACID

Shell Cold Acid Polymerization Process

underhydrated catalyst promotes heavy polymer and coke deposition on the catalyst (26), thus shortening catalyst life, decreasing its activity, and increasing the pressure drop. Overhydration results i n softening of the catalyst, often to the point of plugging the reactor. Reduction i n activity of the catalyst may be caused b y polymer and coke deposition due to excessive temperature, low pressure, insufficient water i n the feed, or too low a space velocity (26). T h e presence of diolefins, oxygen, caustic, or nitrogen bases, such as ammonia or amines i n the feed, also causes loss of catalyst activity. For chamber-type plants the useful catalyst life expected is 100 gallons of polymer per pound of catalyst and about twice that for reactor-type plants (26). The quality of operating procedures can affect the catalyst life to a considerable degree. The polymerization of light olefins using copper pyrophosphate is licensed by The M . W . Kellogg C o . under patents of the Polymerization Process Corp. The process is essentially the same as the U . O . P . process but instead uses a copper pyrophosphate catalyst (18). T h efirstplant was built i n 1939 (22) and several more have been put into operation since that time. A correlation of operating variables for this process was published in 1949 (21); i t shows how conversion is influenced b y catalyst activity, temperature, ratio of propene plus n-butene to isobutene, and the space velocity of olefins and of total feed when operating at 900 pounds per square inch gage pressure. A catalyst life of 100 to 150 gallons of polymer per pound of catalyst is claimed (15). The catalyst is sometimes diluted with charcoal i n the ratio of 1:1 to 2:1 of catalyst to charcoal (21). The charcoal acts as an adsorbent for the phosphoric acid released under operating conditions and distributes the acid over a larger portion of the bed. The phos phoric acid acts as the actual catalytic agent. The reactor inlet temperature is about 400° F . for nonselective operation or about 350° to 370° F . for codimer production from butènes. The temperature rise is of the order of 20° to 50° F . A recent commercial development i n catalytic polymerization of olefins has been made b y the California Research Corp. This process employs a reaction chamber filled with crushed quartz which is coated with liquid phosphoric acid i n situ. The catalyst is reported to have a life of 100 gallons of polymer per pound of 7 5 % b y weight phosphoric In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.


93

acid solution; i t is renewed by washing with water and steam and recoating with 7 5 % phosphoric acid solution without disturbing the quartz. The catalyst is understood to be rugged, permitting operation without recycle on feed of any olefin content obtained i n normal refinery operations. The exothermic heat of reaction is controlled b y quenching with cold feed at spaced points throughout the bed as required. I t is claimed that the process is easy to control, requires a minimum of pretreatment of feed stock, and accom plishes conversion of 9 0 % or better of the olefins to polymer with most normal feed stocks. P o l y m e r i z a t i o n w i t h S u l f u r i c A c i d . I n the early 1930'& development work o n sulfuric acid-catalyzed polymerization was undertaken b y a number of research organizations. The most widely used process is that developed by the Shell companies Downloaded by UNIV OF CALIFORNIA SAN DIEGO on June 2, 2015 | http://pubs.acs.org Publication Date: January 1, 1951 | doi: 10.1021/ba-1951-0005.ch009

OS).

The cold acid process was developed first and effects polymerization of isobutene only. Figure 13 shows a simplified flow diagram of the process. T w o contact stages are generally used i n counterflow. I n the second stage the freshly regenerated acid contacts the hydrocarbons from the first stage, and the intermediate acid contacts the fresh feed in the first stage. The acid, with absorbed isobutene, is heated to cause polymerization, then cooled; the polymer which forms a separate phase is withdrawn, and the acid then is ready for re-use. The absorption step occurs at a temperature of about 68° to 104° F . and whatever pressure is necessary to maintain the hydrocarbon feed i n the liquid phase. When using a 6 5 % acid about 90 to 9 5 % of the isobutene is absorbed. Polymerization takes place at a temperature of 200° to 220° F., producing approximately 75 to 8 0 % dimer, the rest trimer. Thus about 6 7 % of the isobutene in the feed is converted to iso-octenes. The hot acid process was developed as a logical outgrowth of the cold acid process (13,15,16). The hydrocarbon is contacted with acid, varying from 63 to 7 2 % in strength, at a temperature of 167 ° to 212 F . A t this temperature the n-butenes as well as isobutene are absorbed, polymerized almost at once, and the higher molecular weight olefins so formed are preferentially reabsorbed b y the hydrocarbon phase, minimizing further polymeriza tion. Figure 14 is a flow diagram for a typical plant. T o promote the cross polymerization of olefins the concentration of the more reactive isobutene is kept low by recycling a stream of hydrocarbon-acid emulsion having a low iso0

L

RESIDUAL BUTANE TO STORAGE

•CAUSTIC SOOA

OIMER PROD. TO STORAGE^] HYDROLYZER TIME TANK DlMERv COLUMN^

TIME TANK

Y

FINAL ACID SETTLER

±5U

WEACTOR SEPARATOR

WIUW

A f ^REACTOR COOLER

ι

ι

ACID ENTRAINED

REACTOR CIRCULATING PUMP

1\ REBOILER^-

HY0ROLYZER FEEP PUMP - FRESH ACID FEED CHARGE• BUTANE - BUTYLENE FRACT.

HYDROLYZER SEPARATOR, -j

HYDROLYZER HEATER

HYDROLYZER CIRCULATING PUMP

SPENT-4 CAUSTIC BOTTOMS TO STORAGE

Figure 14. Shell Hot Acid Polymerization Process In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.


94

Table I. Process

Thermal Polymerization Pure Oil Co. data

07)

Temperature, ° F . Pressure, lb./sq. inch Feed Unsaturates, %


Inspection of Typical Thermal Polymer Gasolines

Inspection of gasoline Gravity, ° A P I A S T M Diet. I.B.P., ° F . 10% 50% 90% E.p., ° F . RVP&, lb./sq. inch Composition, vol. % Olefins Aromatics Naphthenes Paraffins Octane No. Motor, clear With 3 cc. T E L Research, clear With 3 cc. T E L

(2)

a

1040 1225

Gas Reversion U)

1154 61 Ci-C 47

1090 625-500 C3-C4

C3-C4

34.4

69.4

68.9

62.5

62.0

120 167 212 298 399

106 123 143 228 348 9.1

100 116 152 245 378 10.7

104 127 179 303 428 8.9

93 123 210 345 406 8.5

36.1 3.4 31.3 29.2

40 5 17 38

72.2 78.6

77.0

77.5

69.0 80.5

4

*47

35-45

Naphtha: gas ratio 1:1

35

89*2

Clay treated. * Reid vapor pressure.

a

butene content. A t contact times of 10 to 15 minutes, about 85 to 9 0 % of the C4 olefins are converted to polymer containing up to 90 or 9 5 % octenes. Sulfuric acid polymerization plants found their greatest application i n the aviation gasoline program. C o l d acid plants were generally converted to the hot acid type because the yield of the latter is much higher with only a very small sacrifice i n octane value of the hydrogenated dimer.

Properties of Polymer Motor Fuel Some of the properties of typical polymer gasolines are shown i n Tables I and I I . Inspections are given for thermal as well as catalytic polymer gasolines. Catalytic polymer gasolines are characterized b y their high olefin content, high octane values i n straightrun blends, and their high gum-forming tendencies. However, the gasoline is very susceptible to the action of inhibitors and requires no rerunning. N o noticeable effect of space velocity on the octane number of nonselective polymers is observable (3). T h e composition of thermal gasolines is influenced primarily by reaction temperature.

Mechanism of Polymerization of Alkenes (Olefins) Thermodynamic calculations have shown the areas of temperature and pressure i n which polymerization will tend to take place. The free energy ( ÂF°) of polymerization of normal olefins (above ethene) has been expressed b y AF° « - 20,320 + 3 3 . 2 6 Γ {19) Other thermodynamic equations indicate that pressure favors polymerization. Although the free energy of polymerization is more favorable at low temperatures, rather high tem peratures, or powerful catalysts must be used to overcome the passivity of olefins and to i n crease the rate of polymerization. The mechanism of catalytic polymerization of alkenes to motor fuel recently has been ably discussed b y Schmerling and Ipatieff i n one chapter of a current book on catalysis M a n y theories have been proposed but three have received considerable attention: Whitmore's carbonium ion theory {26) postulates that a carbonium ion (positive hydro carbon ion) adds to an olefin to form a higher molecular weight carbonium ion which then yields the olefin polymer by elimination of a proton (H+). W i t h acid catalysts—for ex ample, sulfuric acid—the initial carbonium ion is formed b y addition of a hydrogen ion from the acid to the extra electron pair i n the double bond of the olefin. A second proIn PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.

95


Table II.

(Phosphoric acid catalyst) Nonselective Nonselective Nonselective {26) {26) {26)

Process Temperature, ° F . Pressure, lb./sq. inch Feed Unsaturates, %


Inspection of Typical Catalytic Polymer Gasolines

Inspection of gasoline Gravity, ° A P I A S T M Dist. I.B.P., ° F . 10% 50% 90% E.p., ° F . R V P a , lb./sq. inch Composition, vol. % Olefins Aromatics Naphthenesl Paraffins J Octane No. Motor, clear With 3 cc. T E L Research, clear With 3 cc. T E L

1000 c

1000 C

375-450 500 C -C 35

62.4

64.3

67.3

144 204 266 330 402 4

78 152 258 379 416 11

90 143 225 367 422 11

4

8

5

Selective {25) 330-350 900 C 56 4

4

Dimer 61.3

Triraer 51.5 261 286 329 379 434

210 223 228 234 257

94.1*> 0.8 5.1 82.5 85.0 97.0 100.0

95.I

e

Reid vapor pressure. *> Ce -f- gasoline. Hydrogenated dimer.

β

e

posed mechanism is quite similar to the carbonium ion mechanism but differs from i t i n that i t assumes that the condensation takes place b y way of addition of an alkyi ester to the olefin to form a higher molecular weight ester which then dissociates to form the polymer and regenerate the acid. For the polymerization of olefins catalyzed by phosphoric acid, Ipatieff (8) has proposed that the polymerization involves the interaction of two molecules of phosphoric acid ester with the elimination of phosphoric acid «and the production of polymer by the union of the two hydrocarbon radicals. Farkas and Farkas (4) have introduced a promising technique for the unraveling of the mechanism of polymerization. These investigators employed a phosphoric acid catalyst i n which the three hydrogens were replaced b y deuterium (heavy hydrogen). The deuterium thus served as a tracer element and thereby enabled the investigators to obtain more specific information about olefin polymerization. They suggest a modifica tion of IpatiefFs theory. Their experiments are not entirely conclusive, but they suggest a very promising method of attack for future researches on the mechanism of polymeriza tion of alkenes.

Future Research The polymerization of gaseous olefins is a relatively simple process. Its increased utilization will depend on an ample supply of light olefins. This suggests further investigation into the problem of producing light olefins from gaseous paraffins at low cost.

Acknowledgment Grateful acknowledgment is made to H . Hennig for his assistance i n preparing this paper, also to the Petroleum Refiner for permission to reproduce Figures 10, 11, 13, and 14, the Oil and Gas Journal for Figure 1, Petroleum Processing for Figure 12, and the California Research Corp. for specific information relative to their phosphoric acid process.

Literature Cited Bogk, J. E., Ostergaard, P., and Smoley, E . R., Refiner Natural Gasoline Mfr., 19, 393 (1940). Carey, J . S., Ibid., 15, 549 (1936). Egloff, G., Oil Gas J., 34, No. 44, 140 (1936). Farkas, Α., and Farkas, L., Ind. Eng. Chem., 34, 716 (1942). Frankenburg, W. G., et al. (editor), "Advances in Catalysis and Related Subjects," Vol. II, p. 21 et seq., New York, Academic Press, Inc., 1950. (6) Frolich. P. K., and Wiezewich, P. J . , Ind. Eng. Chem., 27, 1055 (1935).

(1) (2) (3) (4) (5)


96

(7) (8) (9) (10) (11) (12) (13) (14) (5) (16) (17)


(18) (19) (20) (21) (22) (23) (24) (25) (26)

ADVANCES IN CHEMISTRY SERIES Hirsch, J. H., Ostergaard, P., and Offutt, W. C., Petroleum Refiner, 25, 570 (1946). Ipatieff, V . N . , Ind. Eng. Chem., 27, 1067 (1935). Ipatieff, V . N . , U. S. Patent 1,960,631 (May 29, 1934). Ibid., 2,018,065-6 (Oct. 22, 1935). Ipatieff, V . N . , and Corson, Β. B., Ind. Eng. Chem., 27, 1069 (1935). Ipatieff, V . N . , Corson, Β. B., and Egloff, G., Ibid., 27, 1077 (1935). McAllister, S. H . , Proc. Am. Petroleum Inst.,(III)18, 78 (1937). Offutt, W. C., Ostergaard, P., Fogle, M . C., and Beuther, H., Petroleum Refiner, 25, 554 (1946). Petroleum Refiner, 28, No. 9, 188, 192 (1949). Refiner Natural Gasoline Mfr., 18, 362 (1939). Ridgway, C. M., Wagner, C. R., and Swanson, H . R., Natl. Petroleum News, 28, 47 (Nov. 4, 1936). Ruthruff, R. F., U . S. Patent 2,189,655 (Feb. 6, 1940). Sachanen, A . N . , "Conversion of Petroleum," 2nd ed., p. 47, New York, Reinhold Publishing Corp., 1948. Shanley, W . B., and Egloff, G., Refiner Natural Gasoline Mfr., 18, 227 (1939). Steffens, J . H . , Zimmerman, M . U . , and Laituri, M .J.,Chem. Eng. Progress, 45, 269 (1949). Van Voorhis, M . G., Nat. Petroleum News, 32, R-230 (1940). Wade, H . N . , Ibid., 29, R-347 (1937). Wagner, C. R., Ind. Eng. Chem., 27, 933 (1935). Weinert, P. C., and Egloff, G., Petroleum Processing, 3, 585 (1948). Whitmore, F. C., Ind. Eng. Chem., 26, 94 (1934).

R E C E I V E D April 22,

1951.


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