A process for producing 2,6-dialkylnaphthalene from a hydrocarbon feedstock that contains at least one component selected from the group consisting of dialkylnaphthalene isomers, monoalkylnaphthalene isomers, polyalkylnaphthalenes, and naphthalene, is provided that includes the following steps:
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1. A process for producing 2,6-dialkylnaphthalene from a hydrocarbon feedstock, comprising the following steps:
I. separating the hydrocarbon feedstock and/or a dealkylation product fed from step III into a fraction comprising naphthalene, a fraction comprising monoalkylnaphthalene, a fraction comprising dialkylnaphthalene and a fraction comprising remaining products;
II. separating and purifying 2,6-dialkylnaphthalene from the dialkylnaphthalene fraction of step I;
III. dealkylating the hydrocarbon feedstock and/or the remaining products fraction of step I and feeding the dealkylation product to step I;
IV. alkylating the naphthalene and monoalkylnaphthalene fractions of step I;
wherein the hydrocarbon feedstock is fed to step I and/or step III.
2. The process of
3. The process of
5. The process of
II-1. separating the dialkylnaphthalene fraction of step I into a 2,6-rich-dialkylnaphthalene fraction and a 2,6-lean-dialkylnaphthalene fraction;
II-2. purifying 2,6-dialkylnaphthalene from the 2,6-rich-dialkylnaphthalene fraction from step II-1.
6. The process of
7. The process of
12.36±0.4
11.03±0.2
8.83±0.14
6.18±0.12
6.00±0.10
4.06±0.07
3.91±0.07
3.42±0.06.
8. The process of
9. The process of
10. The process of
11. The process of
12. The process of
13. The process of
16. The process of
18. The process of
20. The process of
23. The process of
24. The process of
12.36±0.4
11.03±0.2
8.83±0.14
6.18±0.12
6.00±0.10
4.06±0.07
3.91±0.07
3.42±0.06.
25. The process of
12.36±0.4
11.03±0.2
8.83±0.14
6.18±0.12
6.00±0.10
4.06±0.07
3.91±0.07
3.42±0.06.
26. The process of
0. 27. A process for preparing a polyester resin comprising:
producing 2,6-dialkylnaphthalene from a hydrocarbon feedstock by the process of
oxidizing the 2,6-dialkylnaphthalene to form 2,6-naphthalenedicarboxylic acid; and
manufacturing the polyester resin from the 2,6-naphthalene-dicarboxylic acid.
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This application is a continuation in part of application Ser. No. 08/948,299 filed on Oct. 10, 1997
where k=2,6-DMN/2,7-DMN at the feed of cooling crystallizer.
Therefore, it is most especially preferable to increase the ratio of 2,6-DMN/2,7-DMN for the higher yield of 2,6-DMN. The fixed bed adsorption can increase the ratio from 1.0 at the feed to 2.0 and more at output, which results in higher separation yields and much lower internal recycling amounts of the overall process.
Preferably, for more effective production, the separation and purification of step II can be separated into separation section step II-1 and purification section step II-2, as can be seen from FIG. 5. In step II-1, the DAN fraction of step I is separated into 2,6-rich-DAN and 2,6-lean-DAN, and in step II-2, 2,6-DAN is purified from the 2,6-rich-DAN fraction from step II-1.
For example, separation of step II-1 may be preferably conducted by using distillation, and purification of step II-2 may be conducted by using cooling crystallization and/or crystallization under high pressure. By such a system, the 2,6-lean-DAN which contains little 2,7-DAN is separated by step II-1, and the 2,6-lean-DAN which contains much 2,7-DAN is separated by step II-2 as a remaining product of purification.
The conditions of HDA of step III include a temperature of about 200 to 1,000° C., and preferably 300 to 700° C., and a pressure of 0 to 250 atmospheres and preferably 5 to 150 atmospheres, a hydrogen circulation rate of from about 500 to 3,000 scf/bbl. The reaction is suitably accomplished utilizing a feed space velocity of about 0.1 to 10.0 hr−1.
One example of a suitable catalyst for HDA, is an activated alumina supported catalyst bearing an oxide of Group VIII metal, preferably, chromium.
One more example of a suitable catalyst for HDA, is an activated alumina supported catalyst bearing an oxide of Group VIII metal and a Group VI metal, preferably, cobalt and molybdenum. The said oxide may be preferably pre-treated at a temperature of 600 to 1000° F. in the presence of organic sulfide.
Other preferred catalysts for HDA include catalysts including a metal selected from the group consisting of noble metal, nickel, and combinations thereof, and a synthetic zeolite characterized by an X-ray diffraction pattern including interplanar d-spacing and relative intensity I/Io×100 as follows:
12.36 ± 0.4
M-VS
11.03 ± 0.2
M-S
8.83 ± 0.14
M-VS
6.18 ± 0.12
M-VS
6.00 ± 0.10
W-M
The preferable noble metal is selected from the group consisting of platinum, palladium, and combinations thereof.
The conditions of alkylation of step IV preferably include a temperature of about 0 to 500° C., and preferably 240 and 450° C., and a pressure of between 0 to 250 atmospheres and preferably 1 to 50 atmospheres. The mole ratio of alkylating agent to feed of monoalkylnaphthalene or naphthalene can be from about 20:1 to 1:20, preferably from 10:1 to 1:10. The reaction is suitably accomplished utilizing a feed space velocity of about 0.1 to 10.0 hr−1.
Preferred alkylating agents include alcohols, olefins, aldehydes, halides, and ethers. For example, methanol, dimethylether and polyalkylbenzene are preferred. Methanol and dimethylether are especially preferred.
A suitable catalyst for alkylation is a synthetic zeolite characterized by an X-ray diffraction pattern including interplanar d-spacing and relative intensity I/Io×100 as follows:
12.36 ± 0.4
M-VS
11.03 ± 0.2
M-S
8.83 ± 0.14
M-VS
6.18 ± 0.12
M-VS
6.00 ± 0.10
W-M
4.06 ± 0.07
W-S
3.91 ± 0.07
M-VS
3.42 ± 0.06
VS
A suitable catalyst is described in U.S. Pat. No. 5,001,295, as MCM-22, the entire contents of which are hereby incorporated by reference.
Preferably, the alkylation can be carried out in any of the known reactors usually employed for alkylation. For example, a tubular reactor with a downflow of reactants over a fixed bed of catalyst can be employed.
In order to maintain high feedstock conversion, the injection of methanol to reactor can be performed, preferably, in multiple stages, and more preferably two stages. For example, one reactor with top and middle methanol feed, or two reactors in series with top and intermediate methanol feed are preferably used.
In a preferred embodiment, the 2,6-lean-DAN fraction from step II-1 may be subjected to isomerization conditions to provide for a dialkylnaphthalene fraction which has a greater content of 2,6-dialkylnaphthalene, as can be seen from FIG. 6. Preferably, the product of the isomerization may be fed to step I and/or step II-2 for more efficient recovery.
Preferred isomerization conditions are generally disclosed in co-pending application U.S. application Ser. No. 08/661,114 (the entire contents of which are hereby incorporated by reference), and are suitable for conducting simultaneous transalkylation of dialkylnaphthalene and naphthalene, and isomerization of dialkylnaphthalenes.
A preferred catalyst for isomerization is a synthetic zeolite characterized by an X-ray diffraction pattern including interplanar d-spacing and relative intensity I/Io×100 as follows:
12.36 ± 0.4
M-VS
11.03 ± 0.2
M-S
8.83 ± 0.14
M-VS
6.18 ± 0.12
M-VS
6.00 ± 0.10
W-M
4.06 ± 0.07
W-S
3.91 ± 0.07
M-VS
3.42 ± 0.06
VS
A suitable catalyst is described in U.S. Pat. No. 5,001,295, as MCM-22, the entire contents of which are hereby incorporated by reference.
Preferably, isomerization is conducted at a weight hourly space velocity (WHSV) of dialkylnaphthalenes of 0.1 to 10, preferably 0.5 to 5 h−1, more preferably 0.75 to 1.5 h−1.
Preferably, isomerization is conducted at a temperature of from 100 to 500° C., preferably 150 to 350° C., more preferably 200 to 300° C.
Preferably, isomerization is conducted at a pressure of atmospheric to 100 kgf/cm2, preferably atmospheric to 30 kcf/cm2.
During isomerization it is optionally preferable to co-feed hydrogen in an amount of 0.1 to 10 mol-H2/mol-hydrocarbons.
According to the preferred embodiment of
In this process, 2,6-lean-DAN as remaining product of separation/purification of step II is dealkylated and fed to separation of step I. So, 2,6-DAN isomers in 2,6-lean-DAN may be changed to MMN or NL and can be alkylated in step IV.
As for the preferred embodiment in
The process scheme of
Having generally described this invention, a further understanding can be obtained by reference to certain specific examples, which are provided herein for purposes of illustration only and are not intended to be limiting unless otherwise specified.
A 153 g amount of MCM-22 catalyst is charged into a tubular reactor (volume:370 cc). As a feedstock for alkylation, 1-MMN, 2-MMN and naphthalene are used, and mixed at a molar ratio of 2.2 of 2-MMN/1-MMN, and a weight ratio of 3.0 of MMN's (1-MMN+2-MMN)/naphthalene.
Thereupon, the feedstock is supplied to the reactor (254 C, 5 kg/cm2) at a rate of 153.4 g/hr and 1.0 hr−1 in WHSV with a feed of hydrogen at the rate of 1.8 ft3/hr. Four hours later, methanol, as an alkylating agent, is introduced into the reactor at 35.5 g/hr, and alkylation is conducted for 20 hours. The product obtained is analyzed by gas chromatography, and the results are summarized in Table 1.
TABLE 1
(Alkylation of Monomethylnaphthalene and Naphthalene)
Before Reaction
After Reaction
Component (wt %)
dimethylnaphthalene
0
17.19
2,6-DMN
0
1.72
2,7-DMN
0
1.20
other isomers
0
14.27
monomethylnaphthalene
73.63
60.10
2-MMN
50.55
40.32
1-MMN
23.08
19.78
naphthalene
25.28
18.67
other component
1.00
3.91
Evaluation
NL conversion (%)
—
26.15
2-MMN/1-MMN
2.2
2.04
MMN conversion (%)
—
18.37
2,6-DMN/total DMN (%)
—
10.02
2,6-DMN/2,7-DMN
—
1.44
As can be seen from Table 1, the ratio of 2,6-DMN/2,7-DMN is over 1.1 and the ratio of 2-MMN/1-MMN is over 2.0.
153 g of MCM-22 were charged in the tubular reactor (volume: 370 cc). As a feedstock for alkylation, 1-MMN (purity 95.5%) and 2-MMN (purity 96.6%) were used, and mixed at the molar ratio of 2.2 of 2-MMN/1-MMN. Feedstock was supplied in the reactor (350° C.) at the rate of 76.7 g/hr and 0.5 hr−1 in WHSV for 4 hours. Thereafter, methanol was started to be supplied in the reactor at the rate of 17.3 g/hr and the reaction was proceeded for 20 hours. The obtained product was analyzed by gas chromatography, and the result is summarized in Table 2.
TABLE 2
(Alkylation)
before reaction
after reaction
component (wt %)
dimethylnaphthalene
0
35.45
2,6-DMN
0
5.12
2,7-DMN
0
4.44
other isomers
0
25.89
monomethylnaphthalene
98.66
41.16
2-MMN
67.61
28.84
1-MMN
31.05
12.32
naphthalene
0
0.19
other component (mainly PAN)
1.53
23.20
evaluation
2-MMN/1-MMN
2.2
2.3
MMN conversion (%)
—
58.28
2,6-DMN/total DMN (%)
—
14.45
2,6-DMN/2,7-DMN
—
1.16
As can be seen from Table 2, the ratio of 2,6-DMN/2,7-DMN is over 1.1 and the ratio of 2-MMN/1-MMN is over 2.0.
Alkylation of MMN and naphthalene has been carried out for several months in the same manner described in Example 1 and about 400 kg of the product is collected. Distillation of the product is carried out by using a batch type distillation tower with a packed column. A number of theoretical trays of the tower is expected to be at least 50. The operation pressure at the top of the column is controlled between 15 and 36 Torr and distillation proceeds at a reflux ratio of 50 to 75.
The product is separated into 17 fractions by differences in boiling points as shown in Table 3.
TABLE 3
(Alkylation and Distillation)
Amount
DMN
2,6-DMN
(kg)
concentration (%)
concentration (%)
Fraction-1-10
270.8
0.0
not analyzed
Fraction-11
30.9
0.5
not analyzed
Fraction-12
8.8
38.9
not analyzed
Fraction-13
11.0
64.8
11.2
Fraction-14
6.3
92.3
25.4
Fraction-15
15.7
99.6
4.3
Fraction-16
4.8
98.7
0.0
Fraction-17
5.3
41.6
0.0
Residue
21.2
0.0
0.0
A part of Fraction-17 and Residue shown in Table 3 are mixed to prepare the feedstock(Blend-A) for hydrodealkylation. A 50 g amount of Cr2O3/Al2O3 type catalyst produced by Sud-Chemie AG is charged into a tubular reactor. The reactor is heated gradually from ambient temperature to 662° F. to dry the catalyst while supplying hydrogen gas. Thereupon Blend-A is fed to the reactor at the rate of 50 g/hr and 1.0 hr−1 in WHSV, while supplying hydrogen gas at 1.2 scf/hr. Hydrodealkylation is carried out at 887° F. and 854 psig. The product is analyzed by GC and the results of hydrodealkylation are summarized in Table 4 below.
As shown in Table 4, Cr2O3/Al2O3 type catalyst is effective to enrich 2,6-DMN from 2,6-DMN lean feed.
TABLE 4
(Hydrodealkylation)
Feed (Blend-A)
HDA Product
naphthalene
0.0
3.11
2-MN
0.0
15.68
1-MN
0.0
3.71
2-EN
0.0
0.57
1-EN
0.0
0.26
2,6-DMN
0.0
4.86
2,7-DMN
0.0
4.58
1,3- + 1,7-DMN
0.0
7.58
1,6-DMN
0.0
2.92
2,3- + 1,4-DMN
0.14
4.84
1,5-DMN
0.0
0.39
1,2-DMN
20.18
9.42
1,8-DMN
0.0
0.12
Unknowns before first DMN
0.0
2.50
Unknowns between DMN
1.34
0.39
Heavies Including
78.34
39.08
Polymethylnaphthalenes
Total DMNs (%)
20.28
34.71
2,6-DMN/Total DMNs (%)
0.0
13.20
A 25 g amount of MCM-22 catalyst is charged into the tubular reactor (volume: 200 cc). The reactor is heated gradually from ambient temperature to 400° C. to dry the catalyst while supplying nitrogen gas, and the flow of nitrogen gas is ceased when the temperature becomes stable at 400° C. Thereupon, 2,6-lean-DMN is supplied to the reactor at the rate of 25 g/hr and 1.0 hr−1 in WHSV, and isomerization of DMN is carried out for four hours. The contents of the obtained product are analyzed by gas chromatography, and the results are summarized in Table 5.
TABLE 5
(Isomerization)
before reaction
After reaction
Component (%)
dimethylnaphthalene
98.09
80.10
2,6-DMN
6.21
13.96
2,7-DMN
8.48
8.66
other isomers
83.40
57.48
monoethylnaphthalene
0.20
9.77
2-MMN
0.03
6.71
1-MMN
0.17
3.06
naphthalene
0
0.78
other component
1.71
9.35
evaluation
2,6-DMN/total DMN (%)
6.3
17.4
2,6-DMN/2,7-DMN
0.73
1.61
(1) Crystallization under High Pressure Crystallization
A 1,505 g amount of DMN isomers is supplied into the high pressure crystallizer (KOBELCO 1.5 L type), and 236 g of 2,6-DNN crystals (purity 87%) are separated under the condition of 2,000 kgf/cm2 and 45° C.
(2) Cooling Crystallization
Using a vessel for crystallization (3 liter), 2,001 g of DMN isomers is cooled quickly from 50 to 40° C. with slow stirring. Then, 0.5 g of seed crystals are charged to the vessel which is kept at a temperature at 40° C. for an hour. Thereupon, the feedstock is cooled to 10° C. at 2° C./min. A 360 g amount of 2,6-DMN crystals (purity 68%) is separated by filtration under pressure.
The results of separation by both crystallization under high pressure and cooling crystallization are summarized in Table 6.
TABLE 6
(Separation and Purification)
Component (g)
before crystallization
crystal
filtrate
CRYSTALLIZATION UNDER HIGH PRESSURE
2,6-DMN
301
205
96
2,7-DMN
232
22
210
other DMN
972
9
963
TOTAL
1505
236
1269
2,6-DMN/2,7-DMN
1.3
—
0.5
2,6-DMN/total DMN
20.0%
—
7.6%
purity of crystal
—
87%
—
recovery of 2,6-DMN
—
68%
—
yield of 2,6-DMN
—
13.6%
—
COOLING CRYSTALLIZATION
2,6-DMN
400
244
156
2,7-DMN
308
67
241
other DMN
1293
49
1244
TOTAL
2001
360
1641
2,6-DMN/2,7-DMN
1.3
—
0.65
2,6-DMN/total DMN
20.0%
—
9.5%
purity of crystal
—
68%
—
recovery of 2,6-DMN
—
61%
—
yield of 2,6-DMN
—
12.2%
—
“Recovery of 2,6-DMN” means the content of 2,6-DMN in the crystals against the content of 2,6-DMN in the feedstock.
“Yield of 2,6-DMN” means the content of 2,6-DMN in the crystal against the total weight of feedstock.
As shown in Table 6, the yield of 2,6-DMN by crystallization under high pressure is much higher than by cooling crystallization. Further, the 2,6-DMN/total-DMN ratio of the filtrate by crystallization under high pressure is less than 8%. Therefore, the filtrate is more effective as a feedstock for transalkylation and isomerization of 2,6-lean-DMN.
Furthermore, when an attempt is made to increase the purity of crystals by cooling crystallization, the yield of 2,6-DMN decreases drastically.
Pre-condensation of 2,6-DMN from DMN mixture (Table 7) was tried by cooling crystallization and a 2,6-DMN rich cake, which is to be used as feedstock for the crystallization under high pressure, was separated by bench scale pressure filtration unit.
Purification of 2,6-DMN from the 2,6-DMN rich cake was carried out by the crystallization under high pressure method using Kobelco's HPC test machine.
Several series of experiments were performed and results are summarized in FIG. 9.
As can be seen in
TABLE 7
(Composition of DMN Mixture)
DMN mixture
Sulfur (ppm)
43.0
Nitrogen (ppm)
140
2,6-DMN
13.8
2,7-DMN
14.1
1,6-DMN
7.1
1,3- & 1,7-DMN
20.5
1,4-DMN
—
1,2- & 1,5-DMN
0.4
2,3-DMN
0.5
Others
43.6
Two types of batch distillation tower are used for the separation of alkylnaphthalenes from LCO. One of the distillation tower (Fractioneer-A) has 167 liters still and 32 foot long column with PRO-PAK (Scientific Development Company) and the other distillation tower (Fractioneer-B) has 27 liters still with an 11 foot long column with PRO-PAK. 164 kg of LCO is charged into the Fractioneer-A and distillation is carried out at a reflux ratio of 50 and a pressure of 60 Torr. 80.5 liters are taken at a take off rate of 0.7 liters per hour.
Then 25 kg of the residue in the still of Fractioneer-A is taken out after the first distillation and charged into Fractioneer-B. Another batch distillation is carried out at a reflux ratio of 50 and a pressure of 50 Torr. 14 liters are taken at a take off rate of 125 ml per hour. The components of the product (Blend-B) obtained from the above-mentioned two step distillation are shown in Table 8.
A 70 g amount of Cr2O3/Al2O3 type catalyst produced by Sud-Chemie AG is charged into a tubular reactor. The reactor is heated gradually from ambient temperature to 932° F. to dry the catalyst while supplying hydrogen gas. Thereupon distillation product (Blend-B) obtained from Example 8 is supplied to the reactor at the rate of 70 g/hr and 1.0 hr−1 in WHSV, while supplying hydrogen gas at 0.98 scf/hr. The hydrodealkylation reaction is carried out at 933° F. and 1138 psig. The product is analyzed by GC and the results of hydrodealkylation are summarized in Table 8 below.
A 70 g amount of CoO/MoO3/Al2O3 type catalyst produced by Akzo Chemicals Inc. is charged into a tubular reactor. The reactor is heated gradually from ambient temperature to 300° F. with nitrogen flow at 5 scf/hr. Then the flow gas is switched to hydrogen at 2 scf/hr and pressure is increased to 500 psig. Catalyst is contacted with an organic sulfide (Kerosene with 1.0% of Dimethyldisulfide) for sulfiding while supplying hydrogen gas and then temperature is raised to 650° F. Thereupon distillation product (Blend-B) obtained from Example 8 is fed to the reactor at the rate of 70 g/hr and 1.0 hr−1 in WHSV, while supplying hydrogen gas at 0.98 scf/hr. Hydrodealkylation is carried out at 932° F. and 1425 psig. The product is analyzed by GC and the results of hydrodealkylation are summarized in Table 8 below.
As shown in Table 8, both Cr2O3/Al2O3, and CoO/MoO3/Al2O3 type catalyst are effective to enrich DMN isomers from DMN lean feed.
TABLE 8
(Hydrodealkylation)
Example 7
Distillation
Product
Example 8
Example 9
(Blend-B)
HDA Product
HDA Product
Catalyst
Cr2O3/
CoO/MoO3/
Al2O3
Al2O3
naphthalene
0.0
1.33
4.75
2-MN
0.0
4.31
9.48
1-MN
0.0
1.05
2.44
2-EN
0.0
1.56
2.13
1-EN
0.0
0.36
0.44
2,6-DMN
0.0
2.46
3.53
2,7-DMN
0.0
2.56
3.76
1,3- & 1,7-DMN
0.0
2.55
3.82
1,6-DMN
0.0
1.43
2.17
2,3- & 1,4-DMN
0.0
1.18
1.44
1,5-DMN
0.0
0.24
0.34
1,2-DMN
0.04
0.40
0.32
1,8-DMN
0.04
0.25
0.47
Unknowns before
0.0
8.67
17.68
first DMN
Unknowns
0.0
0.54
0.68
between DMN
Heavies including
99.92
71.11
46.56
polymethylnaphthalenes
Total DMNs (%)
0.08
11.07
15.85
2,6-DMN-Total DMNs
0.0
22.22
22.27
164 kg of LCO is charged into the Fractioneer-A and distillation is carried out at a reflux ratio of 50 and a pressure of 60 Torr. 120 liters are taken at a take off rate of 0.7 liters per hour and prepared for hydrotreating feedstock as Blend-C. NiO/MoO3/Al2O3 type catalyst produced by Akzo Chemicals Inc. is chosen as a hydrotreating catalyst and charged into a tubular reactor.
After drying and sulfiding catalyst, then Blend-C is fed into the reactor at the rate of 0.43 hr−1 in WHSV and hydrotreating is carried out at 400 psig and 726° F., while supplying hydrogen gas at 3495 scf/bbl. The results of hydrotreating are summarized in Table 9 below.
As shown in Table 9, NiO/MoO3/Al2O3 type catalyst is effective to reduce the nitrogen and/or sulfur compounds in LCO with minimum loss of DMN isomers.
TABLE 9
(Distillation and Hydrotreating)
400
400
500
F. +
F. +
F. +
Total
Nit.
Sul
Conv.
Conv.
Conv.
DMNs
ppm
ppm
API
(%)
(%)
(%)
wt %
Blend-C
230
388
18.7
—
—
—
26.56
HDT
7
5
20.6
0.6
7
21
21.24
Product
2,6-DMN and 2,7-DMN are mixed and dissolved into iso-octane at 2.0 wt % of concentration respectively. Then DMN-isooctane solution is fed to the adsorption column (4.6 mm ID and 500 mmL) packed with K-Y zeolite at the rate of 0.50 ml/min, while column temperature is controlled at 158° F. Time course data of the DMN concentration in the effluent is gathered by GC analysis and breakthrough curve in adsorption step is obtained.
After the adsorption step, liquid feed is switched to the pure iso-octane and effluent is also analyzed to gather the time course data of the DMN concentration in the effluent.
The results of the breakthrough curve and desorption curve are summarized in FIG. 10. As shown in the figure, 2,7-DMN is priorly adsorbed in zeolite compared to 2,6-DMN and it is obviously possible to improve 2,6-/2,7-DMN by contacting DMN isomers with the K-Y type zeolite.
Another adsorption test is carried out in the same manner described in Example 12 except that mesitylene is used as a solvent. The results of the breakthrough curve and desorption curve are summarized in FIG. 11.
Obviously, numerous modifications and variations of the present invention are possible in light of the above teachings. It is therefore to be understood that within the scope of the appended claims, the invention may be practiced otherwise than as specifically described herein.
This application is based on U.S. application Ser. No. 08/948,299, filed Oct. 10, 1997, the entire contents of which are hereby incorporated by reference.
Yamamoto, Koji, Motoyuki, Masahiro, McWilliams, John Paul, Sapre, Ajit Vishwanath, Donnelly, Susan Patricia
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